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The role of Mg-Al hydrotalcite derived mixed oxides ascatalytic support materials : Applications in thetransesterification of vegetable oils for biodiesel

production and in the steam reforming of glycerol forhydrogen production

Eliane Dahdah

To cite this version:Eliane Dahdah. The role of Mg-Al hydrotalcite derived mixed oxides as catalytic support materi-als : Applications in the transesterification of vegetable oils for biodiesel production and in the steamreforming of glycerol for hydrogen production. Catalysis. Université du Littoral Côte d’Opale; Uni-versité de Balamand (Tripoli, Liban), 2018. English. �NNT : 2018DUNK0498�. �tel-02137402�

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N° d'ordre : ULCO

THÈSE de doctorat en Cotutelle

Pour obtenir le grade de Docteur délivré par

L’UNIVERSITÉ DU LITTORAL CÔTE D’OPALE

et

UNIVERSITY OF BALAMAND

Spécialité : Chimie

Présentée par

Eliane DAHDAH

Titre:

The role of Mg-Al hydrotalcite derived mixed oxides as catalytic support materials:

Applications in the transesterification of vegetable oils for biodiesel production and in the

steam reforming of glycerol for hydrogen production

Soutenue publiquement le 17 decembre 2018 devant le jury composé de :

Jean Luc BLIN Professeur, Université de Lorraine Rapporteur

Madona LABAKI Professeur, Université Libanaise de Beyrouth Rapporteur

Mouïn HAMZÉ Professeur, Secrétaire général du CNRS Liban Examinateur

Nancy ZGHEIB Docteur, Université Saint-Esprit de Kaslik Examinateur

Jane ESTEPHANE Professeur, Université de Balamand Encadrant

Cédric GENNEQUIN Maître de conférences, ULCO Encadrant

Samer AOUAD Professeur, Université de Balamand Directeur de thèse

Edmond ABI AAD Professeur, ULCO Directeur de thèse

ACKNOWLEDGEMENTS

The journey behind this thesis has been an unforgettable chapter of my life that I will always

treasure. I have been blessed with an experience that has made me grow, not only scientifically,

but also as a person.

This experience would have never been possible without the ARCUS E2D2 project which granted

me the opportunity as well as the financial support required to accomplish this thesis between

France and Lebanon. I would also like to thank the University of Balamand research council

through the BRG5/2016.

There are several individuals that have been involved in this journey that I would sincerely like to

thank.

I would first like to thank Pr. Edmond ABI-AAD, my thesis director in France, for his trust,

kindness and guidance throughout these three years. I sincerely thank him for seeing my potential.

I would like to express my sincere gratitude to Dr. Samer AOUAD, my thesis director in Lebanon,

for his availability and valuable scientific advice. It has been an honor to be his student starting

from my bachelor’s degree all the way up to my Ph.D thesis.

Thank you to Dr. Cedric GENNEQUIN, my thesis supervisor in France, for his encouragement,

scientific advice and help. I thank him for always having an answer to all of my questions.

My sincere thanks to Dr. Jane ESTEPHANE, my thesis supervisor in Lebanon, for her motivation,

kindness and availability. I am very grateful for the opportunity to have been supervised by her

during this Ph.D.

I would sincerely like to thank the jury members: Pr. Jean-Luc BLIN, Pr. Madona LABAKI, Pr.

Mouïn HAMZÉ and Dr. Nancy ZGHEIB for the interest they have shown in this work and the time

they have devoted to reading this manuscript.

My gratitude goes to Pr. Dominique COURCOT, director of the U.C.E.I.V at the Université du

Littoral Côte D’Opale, Dr. Jihad ATTIEH, Dean of the Faculty of Sciences at the University of

Balamand and Dr Henri ZAKHEM, chairperson of the Chemical Engineering Department at the

University of Balamand for their warm welcomes. A special thank you to Dr. Hanna EL NAKAT,

chairperson of the Department of Chemistry at the University of Balamand, for his continuous

support and encouragement throughout my university years.

A special thank you to Pr. Antoine ABOUKAIS for his advice and encouraging words.

Heartfelt thanks to Miss Waed ATTIEH.

I would sincerely like to thank the teams working at the University of Balamand. In the Chemistry

Department: Miss Amal AOUN, Mr Bilal EL-KHOURY, Miss Sherine SALIBA and Miss Dima

MOUSSA. I would also like to thank the team working at the Chemical Engineering Department:

Miss Katia CHEDID, Miss Vanessa KHOURY and Miss Hala FALLAH.

I would also like to thank all the members of the U.C.E.I.V at the Université du Littoral Côte

D’Opale including Pr Stéphane SIFFERT, Dr Fred LEDOUX, and Dr Renaud COUSIN. My

sincere thanks to Dr. Lucette TIDAHY for her consistent help with the proper functioning of the

experiments. I would also like to thank Dr. Christophe POUPIN for conducting the XRD

experiments.

Thank you to all the people I have met along this journey that I am proud to call my colleagues

and friends, Aida, Carole, Cynthia, Eliane Bsaibess, Eliane Saliba, Elie, Ghida, Hadi, Jihane,

Joudia, Joyce, Maria, Muriel, Nancy, Nathalie, Pamela, Raya, Rebecca, Rita, Romain, Sami,

Sandy, Sara, Tarek, Tracy, Yara and Yasmine. Each and every one of them has been a part of so

many memories that will always make me smile.

Special thanks to Reem HAYDAR and Yara YOUSSEF for their contribution to the biodiesel results

of this work.

This work would never have been complete without the support and the presence of my family who

have always believed in me and motivated me to do my best. To my parents and younger siblings:

Mum, Dad, Elio, Odile, Rebecca and Antonio, you are all my greatest blessing and I dedicate this

work to you all. Thank you all for always being there.

I humbly end my acknowledgments with my gratitude to the Lord Almighty who has been with me

through every step of this journey.

ABSTRACT

This work aims to study the transesterification of sunflower oil for biodiesel production and the

steam reforming of glycerol for renewable hydrogen production. Mg-Al hydrotalcite derived

mixed oxides were chosen as catalytic support materials given their known basicity, thermal

stability and low cost. Only one active phase was studied for biodiesel production (Ca) and two

different active phases (Ru and Ni) for glycerol steam reforming. The various prepared catalysts

were characterized by different techniques such as X-Ray Diffraction (XRD), specific surface area

determination by BET method, Fourier Transform Infrared Spectroscopy (FTIR), H2-Temperature

Programmed Reduction (H2-TPR), CO2-Temperature Programmed Desorption (CO2-TPD) and

Simultaneous Thermogravimetric- Differential Scanning Calorimetry (TG-DSC).

For biodiesel production, the effects of thermal treatment and the preparation method for Ca

incorporation were both studied. The catalyst that showed the best catalytic performance was an

uncalcined Mg4Al2 support impregnated with 40 wt% Ca followed by a calcination at 600oC (Ca

600/Mg4Al2HT) due to its enhanced basic properties. After several optimization steps over this

catalyst, the optimum conditions for biodiesel synthesis were: a methanol to oil molar ratio of 15:1,

a catalyst to oil ratio of 2.5 wt% and a reaction time of 6 hours which gave a FAME yield of 95%.

The properties of the produced biodiesel were studied and found to be in good agreement with

ASTM (American Society for Testing and Materials) requirements.

For the catalytic steam reforming of glycerol, the activity of Ni-based pure metal oxides (Ni/CeO2,

Ni/Y2O3, Ni/ZrO2) was first studied as pure metal oxides are commonly used as support materials.

The effect of the support was evaluated and the Ni/ZrO2 catalyst was found to be the most efficient

for hydrogen production. To study the effect of the zirconia phase, a tetragonal Ni/ZrO2 was also

prepared. The tetragonal catalyst was less active for hydrogen production compared to the

monoclinic catalyst. A stability test at 600oC also showed the deactivation of the tetragonal

Ni/ZrO2 after 6 hours on stream.

For the catalytic steam reforming of glycerol using Mg-Al hydrotalcite derived mixed oxides, Ru-

based Mg-Al catalysts were first studied. The effect of the preparation method (impregnation vs

grafting) on the catalytic activity of Ru-Mg-Al catalysts was studied. The catalyst prepared by the

impregnation method resulted in a better catalytic activity than the catalyst prepared by the grafting

method as it allowed a higher accessibility of the active phase.

For the Ni-based Mg-Al catalysts, the hydrotalcite supports were modified with La to study the

effects of promoter addition on catalytic properties and activity. The bimetallic effect (1%Ru-

5%Ni) and effect of a higher Ni loading (15 wt%) were also studied. The 5 wt% Ni impregnated

on a La modified Mg-Al support (Ni/Mg6Al1.6La0.4) catalyst was the most efficient for hydrogen

production due to its enhanced basicity and metal-support interaction.

The activities of the most efficient catalysts, Ni/ZrO2 and Ni/Mg6Al1.6La0.4, were compared. Both

catalysts produced similar hydrogen yields. Differences in glycerol conversion to gaseous products

were attributed to a higher formation of liquid by-products over the hydrotalcite support compared

to the zirconia support. Nevertheless, less coke was produced over the hydrotalcite catalyst given

its higher basicity. A stability test at 600oC showed the deactivation of the Ni/Mg6Al1.6La0.4

catalyst after 6 hours on stream. Upon optimization of the reaction conditions on Ni/Mg6Al1.6La0.4,

a significant improvement in the stability was observed as the catalyst lasted for 24 hours on

stream. Therefore, Ni/Mg6Al1.6La0.4 could be a promising candidate for industrial application.

Moreover, the lower cost of the hydrotalcite support compared to a pure metal oxide support makes

it additionally more attractive for industrial application.

Keywords: hydrotalcite, biodiesel, calcium, glycerol, ruthenium, nickel, lanthanum, steam

reforming

RÉSUMÉ

Ce travail vise à étudier la transestérification de l'huile de tournesol pour la production de biodiesel

et le vaporeformage du glycérol pour la production d'hydrogène renouvelable. Les oxydes mixtes

Mg-Al prépares par voie hydrotalcite ont été choisis comme supports catalytiques. Pour la

production de biodiesel, les effets du traitement thermique et de la méthode de préparation pour

l'incorporation du calcium ont été étudiés. Le catalyseur Ca600/Mg4Al2HT calciné à 600°C,

préparé par impregnation de 40 wt% de Ca sur le support Mg4Al2 non calciné, a montré les

meilleures performances catalytiques en raison de ses propriétés basiques améliorées.

Pour le vaporeformage catalytique du glycérol, l'activité des oxydes de métaux purs imprégnés par

le Ni a d'abord été étudiée, puisque les oxydes de métaux purs sont souvent utilisés comme

supports catalytiques. Le catalyseur Ni/ZrO2 s'est révélé le plus efficace pour la production

d'hydrogène. Pour le vaporeformage du glycérol en présence d’oxydes mixtes obtenus par voie

hydrotalcite, les catalyseurs au Ru-Mg-Al ont été étudiés. L'effet de la méthode de préparation

(imprégnation ou greffage) sur l'activité catalytique a été étudié. Le catalyseur préparé par

imprégnation permettait une plus grande accessibilité à la phase active. Pour des catalyseurs Ni-

Mg-Al, les supports ont été modifiés par du lanthane. Le catalyseur Ni/Mg6Al1.6La0.4 était le plus

performant en raison de sa basicité accrue et de son interaction métal-support. Les activités des

catalyseurs Ni/ZrO2 et Ni/Mg6Al1.6La0.4, ont été comparées. Les deux catalyseurs ont produit des

rendements en hydrogène similaires. Moins de coke a été produit sur le catalyseur préparé par voie

hydrotalcite, en raison de sa basicité supérieure. Un test de stabilité à 600°C a montré la

désactivation du catalyseur Ni/Mg6Al1.6La0.4 après 6 heures. Lors de l'optimisation des conditions

de réaction sur le catalyseur Ni/Mg6Al1.6La0.4, une amélioration significative de la stabilité a été

observée, étendant ainsi sa durée de vie à 24 heures.

Mots-clés: hydrotalcite, biodiesel, calcium, glycérol, ruthénium, nickel, lanthane, vaporeformage

Introduction générale

La majeure partie de la demande énergétique mondiale actuelle provient des combustibles fossiles.

Cependant, les préoccupations croissantes liées au réchauffement de la planète, à la croissance

démographique et aux réserves de carburant limitées ont incité les chercheurs à rechercher des

carburants de substitution. L'utilisation de biocarburants renouvelables est une solution

prometteuse, car ils peuvent potentiellement remplacer les combustibles fossiles et sont en outre

neutres en carbone. Le biodiesel, en particulier, peut être produit par la transestérification d'huiles

végétales avec un alcool en présence d'un catalyseur. Le glycérol est obtenu en tant que sous-

produit de cette réaction. Avec la demande croissante pour la production de biodiesel, la production

de glycérol augmente également de manière spectaculaire. Étant donné les utilisations limitées du

glycérol, ces quantités excessives de glycérol provenant de la production de biodiesel pourraient

devenir un problème de gaspillage, le marché international du glycérol étant petit par rapport aux

quantités produites. Néanmoins, ce glycérol peut être valorisé en l’utilisant pour produire de

l’hydrogène. Par conséquent, l'excès de glycérol produit par l'industrie du biodiesel peut être

considéré comme une matière première potentielle pour la production d'hydrogène renouvelable.

Bien que diverses technologies de reformage du glycérol soient disponibles, le procédé de

vaporeformage est le plus largement utilisé et le plus compétitif pour la production d'hydrogène.

Le vaporeformage du glycérol combine la vapeur et le glycérol en présence d'un catalyseur pour

produire de l'hydrogène et du dioxyde de carbone. La présence d'un catalyseur est essentielle

puisqu’elle contribue à augmenter la vitesse de réaction et à contrôler la sélectivité en vue de la

formation des produits désirés. Un catalyseur idéal doit non seulement être actif, mais également

être stable et peu coûteux. Étant donné qu'un catalyseur est généralement composé à la fois d'un

support et d'une phase active, le choix des deux composants est crucial pour la recherche du

catalyseur idéal. Ce travail vise à concevoir des catalyseurs adéquats qui rendront rentables la

production de biodiesel et le vaporeformage du glycérol. Dans la littérature, il a été montré que les

catalyseurs ayant des propriétés basiques sont préférés pour les deux réactions. Pour la production

de biodiesel, les catalyseurs basiques assurent des rendements plus élevés. Pour le vaporeformage

du glycérol, les catalyseurs basiques aident à limiter la formation de coke. Ainsi, les hydrotalcites

à base de Mg-Al ont été choisies en tant que supports pour les deux applications. Les hydrotalcites

sont connues pour leur faible coût, leur stabilité thermique élevée et leur basicité élevée. Des

phases actives adaptées à chaque réaction ont également été sélectionnées. Le calcium a été

sélectionné pour la production du biodiesel, tandis que le nickel et le ruthénium ont été sélectionnés

pour la réaction de vaporeformage du glycérol. Tous les catalyseurs préparés ont été caractérisés

en utilisant diverses techniques et ont ensuite été testés dans leurs réactions respectives.

Ce manuscrit est divisé en 4 chapitres :

• Le premier chapitre consiste en une étude bibliographique qui couvre à la fois la production de

biodiesel par transestérification et le vaporeformage du glycérol. Les catalyseurs utilisés dans

chaque réaction sont revus et une description des hydrotalcites et de leurs propriétés est également

incluse.

• Le deuxième chapitre porte sur l’utilisation d’oxydes mixtes Mg-Al préparés par voie

hydrotalcite et imprégné par du Ca pour la production de biodiesel. Les effets du traitement

thermique et de la méthode de préparation pour l'incorporation du Ca (imprégnation contre

greffage) ont été étudiés. L'optimisation des conditions de réaction a été réalisée sur le catalyseur

le plus performant.

• Le troisième chapitre porte sur l’utilisation d’oxydes de métaux purs (CeO2, Y2O3 et ZrO2)

imprégnés par du Ni pour la réaction de vaporeformage du glycérol, puisqu’ils sont souvent utilisés

en tant que supports catalytiques. L'effet de la phase cristalline du ZrO2 imprégné par du Ni a

ensuite été étudié, suivi de tests de stabilité dans la réaction de vaporeformage du glycérol.

• Le quatrième chapitre porte d’abord sur l’utilisation d’oxydes mixtes Mg-Al préparés par voie

hydrotalcite et imprégnés par du Ru pour la réaction de vaporeformage du glycérol. L'effet de la

méthode de préparation a été étudié (imprégnation contre greffage). Ensuite, des catalyseurs

Ni/Mg-Al dont le support a été modifié par du lanthane ont été évalués dans la même réaction.

L'effet bimétallique (1% Ru-5% Ni) et l'effet d'une teneur en Ni plus élevée (15% en masse de Ni)

ont également été étudiés. Une comparaison entre les catalyseurs les plus performants identifiés

dans les chapitres 3 et 4 est ensuite effectuée. Des tests de stabilité et une optimisation des

conditions de réaction ont finalement été réalisés sur le catalyseur le plus performant.

Chapitre 1 : Étude bibliographique En raison de la croissance démographique constante, de la diminution des réserves de combustibles

fossiles et des préoccupations environnementales croissantes liées à l'utilisation des combustibles

fossiles, l'utilisation de sources d'énergie alternatives est devenue de plus en plus importante

[1],[2]. La recherche de carburants alternatifs moins chers, moins polluants, plus efficaces et

surtout susceptibles de remplacer les combustibles fossiles classiques a récemment conduit à une

focalisation intense sur les biocarburants [2]. Les biocarburants, par définition, sont des sources

d'énergie obtenues à partir de biomasse (cultures, arbres, graisses animales, etc.). En fait, ces

carburants apparaissent comme une solution à long terme car ils sont renouvelables et considérés

théoriquement comme neutres en dioxyde de carbone [3]. Comme le montre la figure 1.1, la

combustion de diesel conventionnel provenant de combustibles fossiles entraîne une accumulation

de CO2 dans l'atmosphère, tandis que le CO2 atmosphérique émis par la combustion d'un

biocarburant est recyclé via la photosynthèse des cultures utilisées pour produire le biocarburant.

Figure 1.1: Cycles d'émission de CO2 du diesel conventionnel et d'un biocarburant [3]

La demande et la production de biodiesel en particulier ont considérablement augmenté au cours

des dernières années [1]. Le biodiesel est non toxique, biodégradable et émet moins de polluants

(CO, SO2 et hydrocarbures) par rapport au diesel conventionnel [4]. Le biodiesel a également un

indice de cétane plus élevé, une teneur en soufre plus faible et un point d’éclair plus élevé [5].

Cependant, en termes de prix, le biodiesel n'est pas encore compétitif. Le biodiesel coûte 1,5 à 3

fois plus cher que le diesel conventionnel [6]. Un moyen de réduire les coûts de production du

biodiesel serait de valoriser son principal sous-produit, le glycérol, afin de lui apporter une valeur

supplémentaire [2].

Le biodiesel peut être produit par la transestérification de triglycérides avec un alcool en présence

d'un catalyseur [7]. Ces triglycérides peuvent être extraits d'huiles végétales telles que les huiles

de tournesol, de soja et de colza [3]. Les matières premières de deuxième génération comprennent

les huiles de cuisson, les graisses et les graisses animales. Au cours des dernières années, les

microalgues ont été utilisées comme matière première de biodiesel de troisième génération [8].

Lorsque l'alcool utilisé est du méthanol, le biodiesel produit est connu sous le nom d'esters

méthyliques d'huiles végétales (EMHV).

La principale réaction de transestérification est représentée dans la figure 1.2. Comme le montre

cette figure, selon la stoechiométrie, trois moles de méthanol sont nécessaires pour convertir une

mole de triglycéride en trois moles d’EMHV et une mole de glycérol en tant que sous-produit.

Figure 1.2: Réaction globale de transestérification des triglycérides avec du méthanol [9]

Comme le montre la figure 1.2, le glycérol est obtenu en tant que sous-produit principal de la

réaction de transestérification. En fait, 10% de la quantité de biodiesel produite est constituée de

glycérol, c'est-à-dire que, pour chaque tonne de biodiesel produite, 100 kg de glycérol sont

également produits [10].

Le glycérol est un produit chimique utilisé dans de nombreuses applications en raison de ses

propriétés. Les principales utilisations sont illustrées dans la figure 1.3. Près des deux tiers des

utilisations industrielles du glycérol sont dans le secteur alimentaire (23%), les soins bucco-

dentaires (20%), les cosmétiques (13%) et le tabac (12%) [1].

Figure 1.3: Utilisations du glycérol

Uses of Glycerol

Foods (23%) Oral Care (20%) Cosmetics (13%) Tobacco (12%)

Drugs (9%) Miscellaneous (9%) Plastics (8%) Alkyds (3%)

Meat casing (2%) Explosive (1%)

L’excès de glycérol sur le marché mondial nécessite des utilisations alternatives [11]. Une solution

prometteuse consiste à utiliser le glycérol comme matière première potentielle pour la production

d'hydrogène renouvelable. Avec les progrès de l'industrie des piles à combustible, la demande en

hydrogène continue de croître [12]. En fait, l'hydrogène peut être produit à partir de glycérol de

plusieurs manières [1], telles que la pyrolyse, l'oxydation partielle, le reformage auto-thermique,

le reformage en phase aqueuse et le vaporeformage.

Vaporeformage du glycérol

Le vaporeformage du glycérol (Équation 1) est un équilibre qui implique la décomposition ou la

pyrolyse du glycérol (Équation 2) suivi de la réaction de gaz à l’eau (WGS) (Équation 3):

C3H8O3 + 3H2O 7H2 + 3CO2 ΔH298K = +128 kJ/mol (1)

C3H8O3 4H2 + 3CO ΔH298K = +251 kJ/mol (2)

CO + H2O CO2 + H2 ΔH298K = -41 kJ/mol (3)

Le vaporeformage du glycérol (glycerol steam reforming, GSR): un moyen de valorisation

prometteur

Le vaporeformage en tant que méthode de valorisation du glycérol est préféré aux autres méthodes

puisqu’il peut être conduit à pression atmosphérique [13]. De plus, le vaporeformage est largement

utilisé pour la production d'hydrogène et par conséquent, les industries sont déjà équipées pour

s'adapter aux conditions de vaporeformage du glycérol.

Par conséquent, l'utilisation du glycérol comme matière première pourrait être à la fois prometteuse

et avantageuse pour les applications industrielles. Cependant, le processus GSR, comme tout autre

processus, se heurte à plusieurs obstacles qui doivent d'abord être surmontés pour permettre une

commercialisation efficace à l'échelle industrielle. L'un de ces obstacles est que la réaction

principale est souvent accompagnée de diverses réactions secondaires affectant la production et la

pureté de l'hydrogène [2]. La formation de coke est également un problème majeur car elle entraîne

la désactivation du catalyseur utilisé. Pour résoudre ces problèmes, il est donc crucial de contrôler

les conditions de fonctionnement et de concevoir de nouveaux catalyseurs capables de maximiser

la production d'hydrogène [3]. Ces catalyseurs doivent non seulement être actifs, mais également

stables et peu coûteux. Les études thermodynamiques aideront à sélectionner les conditions de

fonctionnement appropriées qui réduiront l’influence des réactions secondaires et du dépôt de

coke. Une étude bibliographique sur les catalyseurs testés dans la réaction GSR aidera également

à sélectionner la phase active et le matériau du support pour la conception d'un catalyseur

approprié.

Dans la littérature, les catalyseurs basiques semblent être des candidats prometteurs pour les

réactions de transesterification et de vaporeformage. Les oxydes mixtes Mg-Al préparés par voie

hydrotalcite sont très utiles en tant que catalyseurs basiques [14]. Les travaux porteront donc

principalement sur l’utilisation d’oxydes mixtes dérivés de l’hydrotalcite à base de Mg-Al comme

supports catalytiques. Les phases actives métalliques utilisées pour chaque réaction seront

différentes. Pour la réaction de production du biodiesel, le calcium sera utilisé comme phase active

car le CaO est bon marché et connu pour sa basicité. Alors que pour le vaporeformage du glycérol,

le ruthénium et le nickel seront utilisés comme phases actives compte tenu de leurs activités

connues. Pour les catalyseurs imprégnés par du Ni, les supports seront également modifiés avec

du lanthane pour étudier l'effet de l'addition de ce promoteur.

Les hydrotalcites ou composés apparentés à l’hydrotalcite sont des hydroxydes à structure double

lamellaire répondant à la formule générale: [M2+1-xM

3+x(OH)2]

x+(An-) x/n.mH2O où M2+ est un ion

de métal divalent, M3+ est un ion métallique trivalent et An- est un anion.

Les hydrotalcites sont principalement utilisées après calcination ou décomposition thermique, ce

qui conduit à la formation d'oxydes métalliques. Ces oxydes métalliques présentent plusieurs

propriétés intéressantes telles qu’une surface spécifique élevée, une basicité élevée, un effet

mémoire et une capacité d'échange anionique élevée, une stabilité au frittage, et un coût peu élevé.

Chapitre 2: Production de biodiesel sur des catalyseurs Ca/Mg-Al préparés par voie

hydrotalcite

Le procédé préféré à utiliser lors de la synthèse d'hydrotalcites est la technique de co-précipitation

à pH constant afin d'obtenir une phase homogène lors de la combinaison de plusieurs métaux. Dans

ce chapitre, les effets du traitement thermique et de la méthode de préparation pour l’incorporation

du Ca (greffage contre imprégnation) ont été comparés. Un résumé des échantillons préparés et de

leurs caractéristiques est présenté dans le tableau 2.1. En ce qui concerne la dénomination des

catalyseurs, l'abréviation HT représente les catalyseurs sous leur forme non calcinée, alors que le

« 600 » indique la température de calcination.

Table 2.1: Résumé des catalyseurs préparés

Catalyst

prepared

Calcination

Temperature

(˚C)

Impregnation Calcination

after

impregnation

Sample name

Mg4Al2 - - - Mg4Al2 HT

Mg4Al2 600oC - - Mg4Al2 600

Mg4Al2 - 40 wt% Ca 600˚C Ca600/Mg4Al2HT

Mg4Al2 600oC 40 wt% Ca 600˚C Ca600/Mg4Al2 600

Ca2Mg2Al2 - - - Ca2Mg2Al2HT

Ca2Mg2Al2 600˚C - - Ca2Mg2Al2600

Ca4Al2 - - - Ca4Al2 HT

Ca4Al2 600˚C - - Ca4Al2 600

Tous les catalyseurs préparés ont été testés pour la transestérification de l'huile de tournesol. Les

expériences ont été effectuées à une vitesse d'agitation de 400 tours par minute, un rapport molaire

méthanol/huile de 12:1, un rapport massique de catalyseur à l'huile de 2,5% et une température de

60°C. Les supports d'hydrotalcite ayant différentes compositions (Mg4Al2, Ca2Mg2Al2 et Ca4Al2)

et sous leur forme non calcinée (HT) et calcinée ont été testés dans la réaction de transestérification

pour la production de biodiesel afin d'étudier l'effet de la composition et de la calcination sur le

rendement en EMHV. Les résultats obtenus sont présentés dans la figure 2.1.

1.22

3.75

0.57 0.24

4.12

0.48

0

5

10

15

20

Ca 4Al 2

Ca 2Mg 2

Al 2

Calcined (600)

FA

ME

yie

ld (

%)

Uncalcined (HT)

Mg 4Al 2

Mg 4Al 2

Ca 2Mg 2

Al 2

Ca 4Al 2

Figure 2.1: Rendement en EMHV obtenu sur des catalyseurs non calcinés et calcinés Mg4Al2, Ca2Mg2Al2 et Ca4Al2

(vitesse d'agitation de 400 tr/min, MOMR 12:1, 2,5% en poids de CTOR et une température de 60°C pendant 4 heures)

Dans les mêmes conditions réactionnelles, le support Mg4Al2 a été testé à la fois sous ses formes

non calcinée (Mg4Al2HT) et calcinée (Mg4Al2600). Les résultats obtenus ont montré que les deux

catalyseurs ont une faible activité (rendement inférieur à 2%). Dans le but d'augmenter la basicité

du support, du Ca a été intégré dans la structure de l'hydrotalcite avec différentes compositions.

Deux catalyseurs ont été préparés, le Ca2Mg2Al2 et le Ca4Al2. Les réactions de transestérification

effectuées en présence du Ca2Mg2Al2 calciné et non calciné n'ont montré aucune activité

significative. Le rendement obtenu pour Ca2Mg2Al2 non calciné était de 3,75% et celui du calciné

de 4,12%. Même si la fraction de calcium était encore plus élevée dans le solide Ca4Al2, le

pourcentage en EMHV restait non significatif et était inférieur à 1% pour les formes calcinée et

non calcinée. Il a été rapporté que les hydrotalcites, en l'absence de calcination, ne fournissent

aucune activité catalytique pour la réaction de transestérification [15]. Le manque d'activité des

catalyseurs calcinés et contenant du Ca pourrait être attribué à l'inaccessibilité de la phase active

de calcium du fait du procédé de préparation. Le CaO est connu pour sa bonne activité catalytique

dans la réaction de transestérification, mais il semble que dans ces catalyseurs, le CaO ait été

enfermé dans la structure du support [16].

Par la suite, des supports Mg4Al2 calcinés et non calcinés ont été imprégnés de 40% en masse de

Ca, calcinés à 600°C, puis testés dans la réaction de production de biodiesel (figure 2.2). En

revanche, l’imprégnation au calcium a considérablement amélioré l’activité catalytique observée

pour le catalyseur Ca600/Mg4Al2HT en augmentant le rendement en EMHV à 70%. Le Ca

imprégné à la surface a clairement créé des sites actifs pour que la réaction se produise. CaO est

responsable de la haute activité du catalyseur en raison de la formation d'anions O2- qui servent de

centres actifs [17]. Les diffractogrammes XRD ont vérifié la présence de la phase CaO et du

spinelle MgAl2O4. L'ajout de Ca a amélioré les propriétés basiques.

1.22 0.24

70

0.910

20

40

60

80

100 Mg

4Al

2

Ca 6

00/M

g4A

l 2 6

00

Calc

ined

FA

ME

yie

ld (

%)

Uncalc

ined

Ca 6

00/M

g4A

l 2H

T

Ca impregnated Mg4Al

2

Figure 2.2: Rendement en EMHV obtenu sur Ca imprégné sur support Mg4Al2 (vitesse d'agitation de 400 tr/min, rapport

MOMR 12:1, 2,5% en poids de CTOR et température de 60°C pendant 4 heures)

Optimisation

Pour déterminer les conditions réactionnelles optimales, l’effet du rapport catalyseur/huile, du

rapport molaire méthanol/huile et du temps de réaction sur la transestérification de l’huile de

tournesol a été étudié sur le catalyseur le plus actif: Ca600/Mg4Al2HT.

Effet du rapport catalyseur/huile (CTOR)

La charge du catalyseur affecte de manière significative le rendement en EMHV. Une quantité

insuffisante de catalyseur entraîne une réaction incomplète [18]. Lorsque le pourcentage de

catalyseur est augmenté, le rendement est amélioré. A mesure que la masse du catalyseur

augmente, le nombre de sites O2- augmente, ce qui adsorbe davantage de H+ provenant de centres

actifs formant du méthanol [17]. Les réactifs auront donc plus de contact avec ces centres actifs,

augmentant ainsi le rendement. Dans ce travail, lorsque la quantité de catalyseur a augmenté de

2,5% à 10% en masse, le rendement est passé de 70% à 73%, comme indiqué sur la figure 2.3. Il

apparaît que la quantité supplémentaire de catalyseur ajoutée n'a pas eu d'effet significatif sur le

rendement. Dans certains cas, une plus grande quantité de catalyseur pour la même quantité d’huile

et de méthanol pourrait résulter en un mélange plus pauvre compte tenu de la viscosité plus élevée

du mélange de réactifs et du catalyseur. L'augmentation de la résistance au transfert de masse qui

en résulte pourrait limiter le rendement en EMHV [19]. La quantité optimale de catalyseur utilisée

dans les expériences suivantes pour la synthèse de biodiesel a donc été choisie pour être de 2,5%

en poids.

70 73

0

20

40

60

80

100Effect of CTOR: Ca 600/Mg

4Al

2HT

10 wt%

FA

ME

yie

ld (

%)

2.5 wt%

Figure 2.3: Effet du rapport catalyseur / huile sur le rendement en EMHV sur le catalyseur Ca600/Mg4Al2HT (vitesse

d'agitation de 400 tr/min, MOMR 12:1 et température de 60°C pendant 4 heures)

Effet du rapport alcool / huile (MOMR)

Le rapport molaire alcool/huile joue un rôle important dans la production de biodiesel. Pour que

la réaction de transestérification se produise, un rapport alcool/huile d’au moins 3:1 doit être

présent [20]. Un excès d’alcool ou de méthanol peut être ajouté pour faire progresser l’équilibre,

ce qui conduit la réaction presque à son terme et maximise la production du biodiesel. Il a été

rapporté que le rendement en biodiesel augmente avec l'augmentation de la quantité de méthanol

[21]. Dans ce travail, une augmentation du rapport méthanol/huile de 12:1 à 15:1 en fixant les

autres paramètres, améliore considérablement le rendement de 70% à 92% (figure 2.4).

70

92

0

20

40

60

80

100

Methanol:Oil=15:1

Effect of Alcohol to Oil ratio:

Ca 600/Mg4Al

2HT

FA

ME

yie

ld (

%)

Methanol:Oil=12:1

Figure 2.4: Effet du rapport alcool / huile sur le rendement en EMHV sur le catalyseur Ca600/Mg4Al2HT (vitesse

d'agitation de 400 tours/minute, 2,5% en poids de CTOR et température de 60°C pendant 4 heures)

Effet du temps de réaction

L'effet du temps de réaction sur le rendement a également été étudié (figure 2.5). La réaction

commence lentement quand on disperse l'alcool dans l'huile mais s'accélère après un certain temps.

A mesure que le temps de réaction augmente, la vitesse de conversion augmente considérablement.

Une expérience a été menée dans laquelle le temps a été augmenté de 4 à 6 heures en fixant les

autres paramètres. La conversion augmentait alors de 70% à 83%.

70

83

0

20

40

60

80

100Effect of Reaction Time:

Ca 600/Mg4Al

2HT

6 hours

FA

ME

yie

ld (

%)

4 hours

Figure 2.5: Effet du temps de réaction sur le rendement en EMHV sur le catalyseur Ca600/Mg4Al2HT (vitesse d'agitation

de 400 tr/min, 2,5% en poids de CTOR, 12:1 MOMR et une température de 60°C)

Effet combiné du rapport alcool / huile et du temps de réaction

Une transestérification a également été réalisée sur le catalyseur Ca600/Mg4Al2HT en utilisant un

rapport méthanol: huile de 15:1 et un temps de réaction de 6 heures (figure 2.6). La réaction a été

laissée pendant 6 heures avec 2,5% en masse de catalyseur, une vitesse d'agitation de 400 tr/min,

un rapport 15:1 de méthanol à huile et une température de 60 ° C. Les résultats obtenus ont montré

que l'augmentation du rapport méthanol/huile diminuait la viscosité du mélange et augmentait le

transfert de masse [8]. Le rendement le plus élevé de 95% a été obtenu en contrôlant ces deux

paramètres.

95

0

20

40

60

80

100

Ca 600/Mg4Al

2HT

FA

ME

yie

ld (

%)

2.5

wt%

Meth

anol:O

il =

15:1

6 h

ours

Figure 2.6: Effet combiné sur le rendement en EMHV sur le catalyseur Ca600/Mg4Al2HT (vitesse d'agitation de 400

tr/min, 2,5% en poids de CTOR, 12:1 MOMR et une température de 60°C)

Des oxydes mixtes dérivés d'hydrotalcite ont été étudiés dans la réaction de transestérification pour

la production de biodiesel. Le catalyseur Ca600/Mg4Al2HT a montré les meilleures performances

catalytiques et a prouvé que l’ajout de Ca est essentiel pour la création de sites actifs. De plus, les

résultats ont montré que le traitement thermique du support/catalyseur et le procédé de préparation

pour l'incorporation de Ca jouent tous deux un rôle majeur. L'imprégnation de la phase active de

Ca sur le support Mg4Al2 suivie d'un traitement thermique était clairement la meilleure méthode

de préparation car le rendement le plus élevé était obtenu sur ce catalyseur. Cela est attribué à une

amélioration des propriétés basiques. Après une série d’expériences, les conditions optimales pour

la synthèse de biodiesel étaient: un rapport molaire du méthanol à l’huile de 15:1, une teneur de

catalyseur de 2,5% en masse par rapport à l’huile et un temps de réaction de 6 heures qui donnait

un rendement en EMHV de 95%. Les propriétés du biodiesel produit ont également été analysées

et sont conformes aux exigences de la norme ASTM (American Society for Testing and Materials).

Dans les chapitres suivants, la valorisation du glycérol, sous-produit obtenu lors de la production

de biodiesel, par la réaction de vaporeformage pour produire de l'hydrogène sera étudiée.

Chapitre 3: Vaporeformage du glycérol sur des oxydes de métaux purs à base de Ni

Dans ce chapitre, l'activité des oxydes de métaux purs, imprégnés par du Ni, a été étudiée dans la

réaction de vaporeformage du glycérol. L'effet du support sur l'activité catalytique a d'abord été

étudié, où Ni/ZrO2 s'est avéré être le catalyseur le plus actif par rapport aux catalyseurs Ni/CeO2

et Ni/Y2O3. L’effet de la phase zircone a ensuite été étudié et le catalyseur monoclinique à base de

zircone s’est révélé plus actif et plus stable que le catalyseur à base de zircone tétragonale. Les

résultats de ce chapitre soulignent l’importance du choix du matériau de support car il détermine

les propriétés physico-chimiques du catalyseur, qui en définissent les performances dans la

réaction de vaporeformage du glycérol. Les supports utilisés dans ce chapitre étant assez coûteux,

il est important d’étudier l’utilisation de supports moins chers, plus attractifs pour une application

industrielle. Par conséquent, dans le chapitre suivant, des oxydes mixtes Mg-Al prépares par voie

hydrotalcite ont été utilisés comme supports.

Chapitre 4: Vaporeformage du glycérol (GSR) sur des catalyseurs dérivés de l'hydrotalcite

à base de Mg-Al et imprégnés par du Ru et/ou par du Ni

Deux méthodes de préparation différentes ont été utilisées pour la synthèse de catalyseurs Ru-Mg-

Al: l'imprégnation et le greffage. Leurs activités catalytiques ont été comparées. L'activité du

support Mg-Al seul sans phase active a également été étudiée. L'analyse DRX a vérifié

l'incorporation de Ru dans la structure d'hydrotalcite du catalyseur greffé. Cependant, le catalyseur

greffé présentait une activité catalytique inférieure à 600°C dans la réaction de GSR. Ceci peut

être attribué à la plus faible accessibilité de la phase active dans le catalyseur greffé du fait de la

préparation. Il semble qu'une partie de la phase active ait été «piégée» dans la majeure partie du

solide et est ainsi devenue moins accessible. Cela a été vérifié par une analyse H2-TPR qui a révélé

des interactions métal-support plus élevées dans le greffé, mises en évidence par des pics de

réduction de température plus élevés mais une consommation globale d'hydrogène plus faible. Les

catalyseurs greffés et imprégnés ont tous deux présenté une désactivation au cours d'un test de

stabilité à court terme (6 heures) en raison de la formation de carbone associée à la formation de

produit liquide. Les différences d'activité catalytique, la répartition des sous-produits liquides et la

formation de coke entre les deux catalyseurs sont une indication forte de la différence de taille des

particules de Ru résultant également du mode de préparation différent. Cette section a également

montré comment l’addition de la phase active Ru aux supports Mg-Al améliorait de manière

significative leur activité catalytique dans la réaction de vaporeformage du glycérol à la suite d’un

éventuel effet synergique. Les résultats de cette section ont été publiés dans le « International

Journal of Hydrogen Energy ».

Dans la section suivante, le nickel a été utilisé comme phase active. Le support Mg6Al2 a été

modifié avec du lanthane en différentes quantités à des fins de comparaison. Le lanthane a été co-

précipité lors de la synthèse des supports. L'influence de l'addition et de la quantité de La sur les

propriétés physico-chimiques et les activités catalytiques a été étudiée. Sur la base des résultats de

la section précédente, la méthode de préparation de l’imprégnation a été utilisée comme moyen

d’ajout de la phase active (Ni) aux supports.

Les supports Mg6Al2, Mg6Al1.8La0.2, Mg6Al1.6La0.4 et Mg6Al1.2La0.8 ont été synthétisés par co-

précipitation des métaux à pH constant. Les propriétés physico-chimiques des supports seuls ont

été étudiées avant et après la calcination afin de vérifier la présence de la structure d'hydrotalcite

et de s'assurer de sa destruction. L'analyse thermique a montré qu'une température de calcination

de 700°C était suffisante pour la destruction de la phase d'hydrotalcite. Les DRX ont vérifié la

présence de pics d'hydrotalcite avant la calcination pour tous les catalyseurs, ainsi que la

destruction de la phase d'hydrotalcite et l'apparition de nouvelles phases après la calcination à

700°C. Les spectres FTIR ont également confirmé la présence de la structure hydrotalcite et sa

destruction après traitement thermique. L'influence de la quantité de promoteur (La) sur les

caractéristiques du support a également été étudiée. Avant la calcination, l'ajout de lanthane a

entraîné une cristallinité plus faible de la phase hydrotalcite. Après calcination, la phase La2O2CO3

a été observée sur tous les supports contenant du La résultant de l’absorption du CO2

atmosphérique par le La2O3. Avec l'augmentation de la teneur en lanthane, les surfaces spécifiques

des supports ont diminué.

Après la préparation et la caractérisation des supports Mg6Al2, Mg6Al1.8La0.2, Mg6Al1.6La0.4 et

Mg6Al1.2La0.8, le Ni a été imprégné sur les supports calcinés. Les solides séchés ont ensuite été

stabilisés thermiquement par calcination à 600°C sous air pendant quatre heures.

Reformage à la vapeur du glycérol sur Ni/Mg6Al2, Ni/Mg6Al1.8La0.2, Ni/Mg6Al1.6La0.4 et

Ni/Mg6Al1.2La0.8

La figure 4.1 montre les rendements en H2, la conversion du glycérol en produits gazeux et les

sélectivités en CO2, CO, CH4 des catalyseurs Ni/Mg6Al2, Ni/Mg6Al1.8La0.2, Ni/Mg6Al1.6La0.4 et

Ni/Mg6Al1.2La0.8.

400 450 500 550 600 650 7000.0

0.2

0.4

0.6

0.8

1.0

1.2

1.4

1.6

1.8

2.0

2.2 Ni/Mg

6Al

2

Ni/Mg6Al

1.8La

0.2

Ni/Mg6Al

1.6La

0.4

Ni/Mg6Al

1.2La

0.8

Hydro

gen y

ield

(m

ol H

2/m

ol gly

cero

l)

Temperature (oC)

(a)

400 450 500 550 600 650 7000

10

20

30

40

50 Ni/Mg

6Al

2

Ni/Mg6Al

1.8La

0.2

Ni/Mg6Al

1.6La

0.4

Ni/Mg6Al

1.2La

0.8

Gly

cero

l C

onvers

ion G

aseous (

%)

Temperature (oC)

(b)

400 450 500 550 600 650 7000

10

20

30

40

50

60

70

80

90

100

CO

2 S

ele

ctivity (

%)

Temperature (oC)

Ni/Mg6Al

2

Ni/Mg6Al

1.8La

0.2

Ni/Mg6Al

1.6La

0.4

Ni/Mg6Al

1.2La

0.8

(c)

400 450 500 550 600 650 7000

10

20

30

40

50

60

70

80

90

100(d)

Ni/Mg6Al

2

Ni/Mg6Al

1.8La

0.2

Ni/Mg6Al

1.6La

0.4

Ni/Mg6Al

1.2La

0.8

CO

Sele

ctivity (

%)

Temperature (oC)

400 450 500 550 600 650 7000

10

20(e) Ni/Mg

6Al

2

Ni/Mg6Al

1.8La

0.2

Ni/Mg6Al

1.6La

0.4

Ni/Mg6Al

1.2La

0.8

CH

4 S

ele

ctivity (

%)

Temperature (oC)

Figure 4.1: (a) Rendement en H2 (b) conversion du glycérol en produits gazeux, sélectivités en (c) CO2 (d) CO (e) CH4

dans la réaction GSR sur Ni/Mg6Al2, Ni/Mg6Al1.8La0.2, Ni/Mg6Al1.6La0.4 et Ni/Mg6Al1.2La0.8 (débit: 0,025 mL/min,

rapport molaire vapeur / glycérol: 9)

En termes de rendement en hydrogène et de conversion du glycérol en produits gazeux, une

tendance claire est observée: Ni/Mg6Al1.6La0.4 > Ni/Mg6Al1.2La0.8 > Ni/Mg6Al2 >

Ni/Mg6Al1.8La0.2. L'augmentation de la conversion du glycérol en produits gazeux peut être

directement corrélée à l'augmentation du rendement en hydrogène.

Le La2O3 est un oxyde hydrophile qui améliore donc l'activation de l’eau adsorbée [22]. Cela

favorise la réaction de gaz à l’eau, particulièrement à des températures plus basses. Pour cela, à

des températures plus basses, tous les catalyseurs contenant du lanthane présentaient des

sélectivités en CO2 et en CO inférieures à celles du Ni/Mg6Al2. Les sites basiques observés pour

tous les catalyseurs dans les profils CO2-TPD favorisent davantage l'adsorption de CO2. Au-delà

de 550oC, le Ni/Mg6Al1.6La0.4 était le seul catalyseur qui démontrait la capacité de reformer le CH4

produit en H2 et CO par le biais de la réaction du vaporeformage du méthane

Il ressort de ces résultats que le rapport Ni/Mg6Al1.6La0.4 contenait la teneur optimale en lanthane

fournissant à la fois une bonne interaction métal-support ainsi qu'une bonne dispersion de Ni.

Les catalyseurs ont également été caractérisés après test pour étudier le dépôt de coke afin d’établir

des corrélations avec l’activité catalytique. Des faibles quantités de coke ont été déposées sur les

quatre catalyseurs lors de la réaction de vaporeformage du glycérol. En effet, les catalyseurs au

Mg-Al sont connus pour résister au dépôt de coke en raison de leurs propriétés basiques. Leurs

sites basiques améliorent l'oxydation des fragments de CHx adsorbés sur le métal actif. L'ajout de

La2O3 en tant que promoteur a également conduit à la gazéification du carbone [23]. Le caractère

hydrophile de La2O3 facilite la gazéification du coke. De plus, les sites basiques des catalyseurs

contenant du lanthane ont amélioré l'adsorption du CO2, ce qui contribue également à l'élimination

du coke via la réaction inverse de Boudouard [24].

De plus, le Ni/Mg6Al2 produit deux types de coke. Le premier type est oxydé à une température

élevée (580oC) et le second type à une température inférieure (460oC). Le coke oxydé à une

température plus basse (< 500°C) est généralement appelé coke encapsulant ou non filamenteux,

alors que le coke oxydé à une température plus élevée (> 500°C) est du coke en forme de whiskers

ou filamenteux. Par conséquent, le catalyseur Ni/Mg6Al2 produisait à la fois du coke filamenteux

et non filamenteux. Le Ni/Mg6Al1.8La0.2 a également produit du coke filamenteux et non

filamenteux, mais avec des températures d'oxydation inférieures. En fait, avec des quantités

croissantes de lanthane dans la composition de catalyseur, le coke formé est devenu plus amorphe

et plus facile à oxyder comme en témoignent les températures plus basses.

Les produits de déshydratation formés en particulier sur des sites acides se transforment en

précurseurs de coke pour encapsuler du coke. Cela peut être lié aux résultats, car tous les

catalyseurs produisaient du coke d'encapsulation résultant des fortes concentrations de produits de

déshydratation, d'acétaldéhyde et d'hydroxyacétone.

Après imprégnation des supports préparés avec 5% en masse de Ni et calcination à 600°C, les

solides obtenus ont été caractérisés. Les DRX ont montré que les périclases (MgO et NiO) étaient

endommagées dans une certaine mesure par l'incorporation de lanthane. L'analyse H2-TPR a révélé

des interactions et des dispersions métal-support variables en fonction de la teneur en lanthane.

Une petite quantité de La a entraîné une interaction plus forte métal-support alors que des quantités

plus importantes ont affaibli cette interaction. Cela a été attribué à une influence sur la dispersion

des métaux. Les profils de CO2-TPD ont montré une basicité accrue de tous les catalyseurs

contenant du lanthane, comme indiqué par des proportions plus élevées de sites de force moyenne

et de sites basiques forts lorsque la teneur en lanthane augmentait. Ainsi, l'ajout de La entraîne des

modifications structurelles qui affectent directement les propriétés acido-basiques du catalyseur.

À la suite de tests catalytiques, le catalyseur Ni/Mg6Al1.6La0.4 présentait les meilleures

performances catalytiques, car sa teneur en lanthane était optimale, ce qui donnait le rendement le

plus élevé en hydrogène et la conversion du glycérol en produits gazeux.

Comparaison à l'oxyde métallique pur à base de Ni

Par la suite, l’activité du catalyseur Ni/Mg6Al1.6La0.4 a été comparée au catalyseur le plus efficace

parmi les oxydes de métaux purs, le Ni/ZrO2. Les deux catalyseurs ont une teneur en Ni similaire

de 5% en masse. La figure 4.2 montre la comparaison des rendements en hydrogène et des

conversions de glycérol en produits gazeux.

Figure 4.2: (a) Rendement en H2 (b) Conversion du glycérol en produits gazeux dans la réaction GSR sur Ni/ZrO2 et

Ni/Mg6Al1.6La0.4 (débit: 0,025 ml / min, rapport molaire vapeur / glycérol: 9)

Les rendements en hydrogène étaient similaires entre les deux catalyseurs malgré les différents

supports utilisés. Néanmoins, le Ni/ZrO2 était clairement plus actif dans la conversion de la

molécule de glycérol en produits gazeux. Les concentrations de produit liquide (non présentées)

obtenues sur le Ni/Mg6Al1.6La0.4 étaient significativement supérieures à Ni/ZrO2. Ceci explique la

différence observée dans la conversion du glycérol en produits gazeux. Le support hydrotalcite

favorise clairement une formation plus élevée de sous-produits liquides alors que l'oxyde

métallique pur produit des quantités plus faibles. Néanmoins, les deux catalyseurs sont capables

de convertir ces produits liquides en produits gazeux à des températures plus élevées, comme le

montre la tendance à la baisse des concentrations de produits liquides. En comparant les résultats

des analyses thermiques des catalyseurs usés (figure 4.3), il est clair que la quantité de coke

produite est nettement inférieure pour le Ni/Mg6Al1.6La0.4 (2,5%) par rapport au Ni/ZrO2 (16%).

Ceci est attribué à la basicité plus élevée du support hydrotalcite, notamment en raison de l'ajout

de lanthane. En outre, les profils DSC montrent que le coke produit sur le support en hydrotalcite

est non filamenteux, alors que celui du support en oxyde de métal pur est filamenteux.

400 450 500 550 600 650 7000

1

2

3(a)

Hydro

gen y

ield

(m

ol H

2/m

ol gly

cero

l)

Temperature (oC)

Ni/ZrO2

Ni/Mg6Al

1.6La

0.4

400 450 500 550 600 650 7000

10

20

30

40

50(b) Ni/ZrO

2

Ni/Mg6Al

1.6La

0.4G

lycero

l C

onvers

ion G

aseous (

%)

Temperature (oC)

Test de stabilité

Un test de stabilité a été effectué à 600°C pour étudier l'activité du catalyseur Ni/Mg6Al1.6La0.4 à

une température fixe sur une période de temps plus longue. Les conversions de glycérol en produits

gazeux, les rendements en hydrogène et les sélectivités en CO2, CO et CH4 ont été suivis pendant

6 heures en fonctionnement à 600°C et les résultats sont présentés sur la figure 4.4 (a) et (b). Les

conversions totales de glycérol et les distributions de produits liquides à des intervalles de 2 heures

sont présentées sur la figure 4.4 (c) et (d).

Le catalyseur Ni/Mg6Al1.6La0.4 a connu une diminution dramatique de la conversion du glycérol

en produits gazeux et du rendement en hydrogène. La baisse est devenue moins exprimée après

les trois premières heures. Cette diminution d'activité pourrait être due à une accumulation de coke

à la surface du catalyseur. La sélectivité supérieure en CO2 par rapport au CO indique l’opération

de la réaction de gaz à l'eau.

L'analyse du produit liquide a permis de vérifier la perte d'activité avec le temps écoulé en

diminuant les valeurs de conversion du glycérol total. Les concentrations de tous les produits

liquides étudiés étaient en général basses à 600°C car la formation de produit liquide est plus

favorisée à des températures plus basses. A des températures plus élevées, ces produits forment

des intermédiaires qui conduisent à la production d'oxydes de carbone et d'hydrogène [25].

Néanmoins, la variation des concentrations de ces produits liquides au fil du temps confirme que

le dépôt de coke est associé aux réactions de déshydratation, de déshydrogénation et

d'hydrogénolyse qui ont lieu sur les sites actifs du catalyseur dans GSR [26].

Figure 4.3: Courbes (a) TG et (b) DSC obtenues pour Ni/ZrO2 et Ni/Mg6Al1.6La0.4 utilisés

100 200 300 400 500 600 700 80084

86

88

90

92

94

96

98

100

102

104(a)

16%

We

igh

t L

oss (

%)

Temperature (oC)

Ni/ZrO2

Ni/Mg6Al

1.6La

0.4

2.5%

100 200 300 400 500 600 700 800

(b) Ni/ZrO2

Ni/Mg6Al

1.6La

0.4

DS

C s

ign

al (a

.u.)

Temperature (oC)

L'analyse des TG après le test de stabilité (figure 4.5) a montré que la formation de coke sur le

catalyseur correspondait à une perte de masse de 8%. Le profil DSC a révélé un pic exothermique

à une température de 452°C. Cela peut être attribué au coke non filamenteux ou encapsulant. Selon

la littérature, l'encapsulation de coke peut bloquer les sites métalliques actifs, ce qui entraîne une

diminution de la conversion [27]. Les résultats de ce test de stabilité vérifient donc que la formation

de coke d'encapsulation affecte l'activité catalytique via le blocage des sites métalliques actifs.

Figure 4.4: (a) Sélectivité des produits C (CO2, CH4, CO) (b) Rendement en H2 et conversion du glycérol en produits

gazeux (c) Concentrations en produit liquide (d) Conversion du glycérol total de l’essai de stabilité sur Ni/Mg6Al1.6La0.4

(débit: 0,025 mL / min, rapport molaire vapeur / glycérol: 9)

0 60 120 180 240 300 3600

10

20

30

40

50

60

70

80

90

100(a) Ni/Mg

6Al

1.6La

0.4S

ele

ctivity to C

pro

du

cts

(%

)

Time (min)

CO2

CH4

CO

0 60 120 180 240 300 3600

10

20

30

40

50

60(b)

Time (min)Gly

cero

l con

vers

ion

to g

aseo

us (

%)

Ni/Mg6Al

1.6La

0.4

0

1

2

3

4

5

6

7 H2 y

ield

(mol H

2 /mol o

f gly

cero

l)

120 240 3600.00

0.05

0.10

0.15

0.20

0.25

0.30

0.35

0.40

0.45

0.50(c)

Co

ncentr

ation (

mol/L)

Time (min)

Acetaldehyde

Acetone

Hydroxyacetone

Ethylene Glycol

Ni/Mg6Al

1.6La

0.4

120 240 3600

10

20

30

40

50

60

70

80

90

100(d) Ni/Mg

6Al

1.6La

0.4

Tota

l G

lycero

l C

onvers

ion (

%)

Time (min)

Figure 4.5: Courbes TG et DSC obtenues pour le Ni/Mg6Al1.6La0.4 utilisé après le test de stabilité (débit: 0,025

mL/min)

100 200 300 400 500 600 700 80086

88

90

92

94

96

98

100

102

TG

DSC

Temperature (oC)

Weig

ht Loss (

%)

Ni/Mg6Al

1.6La

0.4

DS

C s

ign

al (a

.u.)

452oC

8%

Optimisation

Afin d'optimiser l'activité du catalyseur Ni/Mg6Al1.6La0.4 qui s'était rapidement désactivé dans le

test de stabilité précédent, le débit a été réduit à 0,008 mL/min. Le test de stabilité a été effectué à

600°C et les résultats sont présentés à la figure 4.6.

Plusieurs différences ont été notées, surtout que le catalyseur a pu résister pendant 24 heures de

test. De plus, les valeurs de conversion du glycérol et les rendements en hydrogène étaient

nettement supérieurs à ceux du test de stabilité précédent. Ceci peut être attribué au débit plus

faible. En réduisant le débit, la vitesse spatiale horaire du gaz a également été réduite. Des vitesses

spatiales horaires plus basses du gaz permettent plus de temps de contact entre les réactifs et le

catalyseur et entraînent donc une valeur de conversion du glycérol plus élevée [28],[29]. Des

sélectivités plus élevées en CO2 qu’en CO ont tout de même permis de confirmer l’apparition de

la réaction de gaz à l’eau. Il semble que dans ce test de stabilité, le La2O3 a pu gazéifier une partie

Figure 4.6: (a) Sélectivité des produits C (CO2, CH4, CO) (b) Rendement en H2 et conversion du glycérol en produits

gazeux (c) Concentrations en produit liquide (d) Conversion du glycérol total de l’essai de stabilité sur Ni/Mg6Al1.6La0.4

(débit: 0,008 mL / min, rapport molaire vapeur / glycérol: 9)

0 500 1000 15000

10

20

30

40

50

60

70

80

90

100Ni/Mg

6Al

1.6La

0.4

Time (min)Gly

cero

l con

vers

ion

to g

aseo

us (

%)

0

1

2

3

4

5

6

7(b)

H2 y

ield

(mol H

2 /mol o

f gly

cero

l)0 500 1000 15000

10

20

30

40

50

60

70

80

90

100(a) Ni/Mg

6Al

1.6La

0.4

Sele

ctivity to C

pro

du

cts

(%

)

Time (min)

CO2

CH4

CO

1260 14400.00

0.05

0.10

0.15

0.20

0.25

0.30

0.35

0.40

0.45

0.50(c) Ni/Mg

6Al

1.6La

0.4

Co

ncentr

ation (

mol/L)

Time (min)

Acetaldehyde

Acetone

Hydroxyacetone

Ethylene Glycol

1260 14400

10

20

30

40

50

60

70

80

90

100(d) Ni/Mg

6Al

1.6La

0.4

Tota

l G

lycero

l C

onvers

ion (

%)

Time (min)

du coke produit. Néanmoins, avec le temps, la capacité du catalyseur à gazéifier le coke produit a

diminué.

La perte d'activité globale est vérifiée par les conversions totales de glycérol obtenues à partir de

produits liquides recueillis au cours des 6 dernières heures du test de stabilité. Les conversions

totales de glycérol ont diminué de 69% à 51% au cours des 6 dernières heures. Les concentrations

de produit liquide au cours de ces 6 dernières heures ont révélé que les concentrations

d'hydroxyacétone augmentaient avec le temps.

L'analyse thermique après test (figure 4.7) a montré une faible quantité de coke correspondant à

une perte de masse de 3%. Les profils DSC ont révélé deux pics exothermiques distincts

correspondant à deux types de coke (filamenteux et non filamenteux). Les maxima des pics étaient

respectivement de 456 et 580°C. La présence de coke filamenteux indique une modification du

mécanisme de formation de coke résultant de la réduction du débit. Ce coke filamenteux s'est

formé en raison des réactions de Boudouard et de décomposition du méthane dans les nouvelles

conditions. Il semble que les fluctuations observées des sélectivités en CO2 et en CO soient

également dues à la réaction de Boudouard. Les sélectivités décroissantes en CH4 étaient donc

certainement dues à la réaction de décomposition du méthane. Le coke d'encapsulation non

filamenteux s'est formé en raison de la formation d'un produit liquide, en particulier

l'hydroxyacétone et l'acétaldéhyde. Comme il est connu que le coke encapsulant diminue l’activité,

la désactivation observée était finalement due à la formation de ce dernier. Néanmoins, moins de

coke d'encapsulation a été produit étant donné la formation supplémentaire de coke filamenteux,

ce qui explique le temps de désactivation plus long. La formation de coke filamenteux n’affecte

en fin de compte pas la stabilité du catalyseur.

Cette section a clairement démontré que la diminution de la vitesse spatiale horaire des gaz

entraînait un rendement en hydrogène plus élevé ainsi qu'une stabilité accrue [30]. Une quantité

de coke plus faible a également été observée.

Figure 4.7: Courbes TG et DSC obtenues pour le Ni/Mg6Al1.6La0.4 utilisé après le test de stabilité (débit: 0,008 mL/min)

100 200 300 400 500 600 700 80096

98

100

102

TG

DSC

Temperature (oC)

Weig

ht Loss (

%)

580oC

Ni/Mg6Al

1.6La

0.4

DS

C s

ignal (a

.u.)

3%

456oC

Conclusion

L'objectif principal de cette thèse était d'évaluer l'utilisation des oxydes mixtes Mg-Al prépares

par voie hydrotalcite en tant que supports catalytiques dans la réaction de transestérification pour

la production de biodiesel et le vaporeformage du glycérol pour la production d'hydrogène.

Différentes phases actives ont été utilisées pour chaque réaction: le calcium pour la production de

biodiesel, le ruthénium et le nickel pour le vaporeformage du glycérol.

Pour la production de biodiesel, l'activité des catalyseurs Mg-Al à base de Ca a été testée dans la

transestérification de l'huile de tournesol. L'imprégnation de la phase active de Ca sur le support

Mg4Al2 non calciné, suivie d'un traitement thermique (600°C) était clairement la meilleure

méthode de préparation car le rendement le plus élevé était obtenu sur ce catalyseur

(Ca600/Mg4Al2HT). De plus, les conditions réactionnelles optimales pour la production de

biodiesel sur Ca600/Mg4Al2HT se sont révélées être un rapport molaire de méthanol/huile de 15:1,

un rapport catalyseur/huile de 2,5% en masse et un temps de réaction de 6 heures, ce qui a donné

un rendement en EMHV de 95%.

L'activité de différents oxydes de métaux purs imprégnés par du Ni a d'abord été étudiée pour le

vaporeformage catalytique du glycérol. Le catalyseur le plus efficace pour la production

d'hydrogène était le catalyseur monoclinique Ni/ZrO2.

Pour le vaporeformage catalytique du glycérol à l'aide d'oxydes mixtes Mg-Al prépares par voie

hydrotalcite, les catalyseurs à base de Ru ont d'abord été étudiés. Deux catalyseurs Ru-Mg-Al ont

été synthétisés en utilisant deux méthodes de préparation différentes: l'imprégnation et le greffage.

Les propriétés physico-chimiques et les activités catalytiques des deux catalyseurs ont été

comparées. Le catalyseur greffé présentait une activité catalytique inférieure dans la réaction de

GSR. Il a donc été constaté que la technique d’imprégnation était plus efficace car elle permettait

une plus grande accessibilité de la phase active du ruthénium métallique. Cela a été vérifié par une

analyse H2-TPR qui a révélé des interactions métal-support plus fortes pour le catalyseur greffé.

L'activité des oxydes mixtes Mg-Al prépares par voie hydrotalcite imprégnés par du nickel a

ensuite été testée. Les supports Mg-Al ont été modifiés avec du lanthane. Les résultats de la

caractérisation des supports seuls et des supports imprégnés de Ni ont révélé l'effet de

l'incorporation de lanthane. Pour les supports seuls, avant la calcination, l'addition de lanthane a

entraîné une cristallinité plus faible de la phase hydrotalcite, du fait que des espèces de carbonate

de lanthane et d'oxyhydroxyde se sont formées. Avec l'augmentation de la teneur en lanthane, les

surfaces spécifiques des supports ont diminué. Pour les supports imprégnés par du Ni, la DRX a

montré que les périclases (MgO et NiO) étaient endommagées par l'incorporation du lanthane.

L'analyse H2-TPR a révélé des interactions et des dispersions métal-support variables en fonction

de la teneur en lanthane. Les profils CO2-TPD ont montré une basicité accrue de tous les

catalyseurs contenant du lanthane. Le catalyseur Ni/Mg6Al1.6La0.4 présentait la meilleure

performance catalytique, car sa teneur en lanthane était optimale, ce qui donnait le rendement le

plus élevé en hydrogène et la conversion du glycérol en produits gazeux. L'activité du

Ni/Mg6Al1.6La0.4 a ensuite été comparée au catalyseur Ni/ZrO2 (monoclinique). En termes de

production d'hydrogène, les deux catalyseurs ont présenté des activités similaires. Cependant, une

large quantité de sous-produits liquides a été produite par rapport sur le catalyseur dérivé de

l'hydrotalcite en raison de ses sites acides de Lewis. Néanmoins, la basicité plus élevée résultant

de la composition du support (Mg-Al-La) a conduit à une formation beaucoup plus faible de coke.

Le coût relativement bas du Ni/Mg6Al1.6La0.4 en fait également un candidat plus prometteur pour

le GSR. Une étape d'optimisation a permis au catalyseur Ni/Mg6Al1.6La0.4 de rester performant

pendant 24 heures.

Pour compléter ces travaux, il est important de réaliser des études supplémentaires susceptibles

d’aider à l’intensification des processus à l’échelle de l’usine. Cela rendra l'utilisation des

catalyseurs plus attrayante pour une éventuelle application industrielle. Par exemple, pour la

production de biodiesel, il serait intéressant d’optimiser davantage la composition du catalyseur et

/ ou le procédé de préparation afin d’obtenir des rendements en EMHV plus élevés durant des

périodes plus courtes. De plus, un catalyseur n'est considéré comme efficace que s'il peut être

réutilisé pendant plusieurs cycles. Par conséquent, une étude de réutilisabilité est essentielle pour

garantir une bonne activité sur plusieurs cycles de réaction.

Pour la réaction de vaporeformage du glycérol, il serait également intéressant de faire varier la

composition du catalyseur et les méthodes de préparation. Une optimisation plus poussée des

conditions de test pourrait également être réalisée pour augmenter les rendements en hydrogène et

les conversions de glycérol en produits gazeux. Pour une application industrielle, un catalyseur

doit être stable pendant de longues périodes. Par conséquent, une étude de désactivation

approfondie devrait également être menée. Des études mécanistiques pourraient également être

réalisées lorsque des intermédiaires réactifs sont identifiés (FTIR in situ) pour comprendre le rôle

de chaque composant catalytique (phase de support / active) dans la réaction. Tous les tests

catalytiques de cette étude ont été réalisés avec du glycérol pur acheté auprès de Sigma Aldrich.

Le glycérol obtenu directement à partir de la production de biodiesel est en fait considéré comme

«brut» car il contient plusieurs impuretés, notamment du méthanol et du savon. La purification du

glycérol brut étant assez coûteuse, il serait également intéressant de mener la réaction de

vaporeformage sur ce glycérol brut et d'étudier l'influence de ces impuretés sur l'activité des

catalyseurs.

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LIST OF FIGURES

Figure 1.1: CO2 emission cycles of conventional diesel and a biofuel [3] ............................................... 5

Figure 1.2: Worldwide biodiesel production from 2010-2016 [7] ........................................................... 6

Figure 1.3: Leading biodiesel producing countries of 2016 [7] ............................................................... 7

Figure 1.4: Triglyceride molecule [9] ........................................................................................................ 8

Figure 1.5: Overall transesterification reaction of triglycerides with methanol [10] ........................... 8

Figure 1.6: Transesterification reaction steps [11] ................................................................................... 9

Figure 1.7: Hydrolysis of triglycerides [10] .............................................................................................. 9

Figure 1.8: Saponification of FFA [14] .................................................................................................... 10

Figure 1.9: Glycerol molecule .................................................................................................................. 13

Figure 1.10: Boiling points and vapor pressures of glycerol solutions [28] ......................................... 14

Figure 1.11: Uses of glycerol .................................................................................................................... 15

Figure 1.12: Possible reactions for glycerol pyrolysis [25] .................................................................... 16

Figure 1.13: Equilibrium constants of possible reactions in GSR ........................................................ 21

Figure 1.14: Thermodynamic equilibrium calculations (T:400-700oC and WGFR 9:1) [41] ............ 22

Figure 1.15: GSR liquid product distribution [42] ................................................................................ 23

Figure 1.16: Schematic representation of the structure of hydrotalcites ............................................. 29

Figure 1.17: Formation, transformation and gasification of carbon on Ni [82] .................................. 34

Figure 1.18: Formation and transformation of coke on metal surfaces [82] ....................................... 35

Figure 2.1: XRD diffractograms of uncalcined hydrotalcite supports ................................................. 40

Figure 2.2: XRD diffractograms of calcined supports (Mg4Al2600, Ca2Mg2Al2600, Ca4Al2600) and

Ca impregnated Mg4Al2 support (Ca600/Mg4Al2HT, Ca600/Mg4Al2600) ........................................... 41

Figure 2.3: N2 adsorption/desorption isotherms of calcined catalysts .................................................. 42

Figure 2.4: Pore size distribution of calcined catalysts .......................................................................... 43

Figure 2.5: CO2-TPD profiles of calcined catalysts ................................................................................ 44

Figure 2.6: FAME yield obtained over uncalcined and calcined Mg4Al2, Ca2Mg2Al2 and Ca4Al2

catalysts (Stirring rate of 400 rpm, 12:1 MOMR, 2.5 wt% CTOR and temperature of 60 ˚C for 4

hours) ......................................................................................................................................................... 47

Figure 2.7: FAME yield obtained over Ca impregnated on Mg4Al2 support (Stirring rate of 400

rpm, 12:1 MOMR ratio, 2.5 wt% CTOR and temperature of 60 ˚C for 4 hours) .............................. 48

Figure 2.8: Effect of catalyst to oil ratio on FAME yield over the Ca600/Mg4Al2HT catalyst

(Stirring rate of 400 rpm, 12:1 MOMR and temperature of 60˚C for 4 hours) .................................. 49

Figure 2.9: Effect of alcohol to oil ratio on FAME yield over the Ca600/Mg4Al2HT catalyst (Stirring

rate of 400 rpm, 2.5 wt% CTOR, and temperature of 60 ˚C for 4 hours) ........................................... 50

Figure 2.10: Effect of reaction time on FAME yield over the Ca600/Mg4Al2HT catalyst (Stirring

rate of 400 rpm, 2.5 wt% CTOR, 12:1 MOMR and temperature of 60 ˚C) ........................................ 51

Figure 2.11: Combined effect on FAME yield over the Ca600/Mg4Al2HT catalyst (Stirring rate of

400 rpm, 2.5 wt% CTOR, 12:1 MOMR and temperature of 60 ˚C) .................................................... 51

Figure 3.1: XRD diffractograms of (a) supports alone and (b) Ni impregnated catalysts (c)

magnified 36-45o region for Ni impregnated catalysts ........................................................................... 56

Figure 3.2: Specific surface areas of supports alone and Ni impregnated catalysts ........................... 57

Figure 3.3: H2-TPR profiles of (a) CeO2, Ni/CeO2 (b) Y2O3, Ni/Y2O3 and (c) ZrO2, Ni/ZrO2 ............ 59

Figure 3.4: CO2-TPD profiles of supports alone and Ni impregnated catalysts.................................. 61

Figure 3.5: (a) H2 yields (b) Glycerol conversion to gaseous products (c) CO2 (d) CO (e) CH4

selectivities of GSR reaction over empty reactor (flow rate: 0.025 mL/min, steam/glycerol molar

ratio: 9) ....................................................................................................................................................... 63


selectivities of GSR reaction over Ni/CeO2, Ni/Y2O3 and Ni/ZrO2 (flow rate: 0.025 mL/min,

steam/glycerol molar ratio: 9) .................................................................................................................. 65

Figure 3.7: (a) TG and (b) DSC curves obtained for used Ni/CeO2, Ni/Y2O3 and Ni/ZrO2 ............... 68

Figure 3.8: Schematic of different zirconia phases ................................................................................ 71

Figure 3.9: XRD diffractograms of (a) Ni/ZrO2 (M) and Ni/ZrO2 (T) catalysts (b) magnified 36-45o

region .......................................................................................................................................................... 72

Figure 3.10: Specific surface areas of Ni/ZrO2 (M) and Ni/ZrO2 (T) ................................................... 73

Figure 3.11: H2-TPR profiles of Ni/ZrO2 (M) and Ni/ZrO2 (T) ............................................................ 73

Figure 3.12: CO2-TPD profiles of Ni/ZrO2 (M) and Ni/ZrO2 (T) ......................................................... 74


selectivities of GSR reaction over Ni/ZrO2 (M) and Ni/ZrO2 (T) (flow rate: 0.025 mL/min,

steam/glycerol molar ratio: 9) .................................................................................................................. 76

Figure 3.14: (a) TG and (b) DSC curves obtained for used Ni/ZrO2 (M) and Ni/ZrO2 (T)................ 78

Figure 3.15: (a) Selectivity to C products (CO2, CH4, CO) (b) H2 yields and glycerol conversion to

gaseous products (c) Liquid product concentrations (d) Total glycerol conversions of stability test

over Ni/ZrO2 (T) (flow rate: 0.025 mL/min, steam/glycerol molar ratio: 9) ........................................ 79



over Ni/ZrO2 (M) (flow rate: 0.025 mL/min, steam/glycerol molar ratio: 9)....................................... 80

Figure 3.17: a) TG and (b) DSC curves obtained for used Ni/ZrO2 (M) and Ni/ZrO2 (T) after

stability test ................................................................................................................................................ 81

Figure 4.1: XRD diffractograms of the uncalcined Ru-Mg-Al catalysts .............................................. 86

Figure 4.2: XRD diffractograms of the calcined Ru-Mg-Al catalysts .................................................. 87

Figure 4.3: Specific surface areas of the calcined Ru-Mg-Al catalysts ................................................ 88

Figure 4.4: H2-TPR profiles of the calcined Ru-Mg-Al catalysts .......................................................... 89

Figure 4.5: CO2-TPD profiles of calcined Ru-Mg-Al catalysts ............................................................. 90


selectivities of GSR reaction over Ru-Mg-Al catalysts (flow rate: 0.075 mL/min, steam/glycerol

molar ratio: 9) ........................................................................................................................................... 91

Figure 4.7: a) TG and b) DSC curves obtained for used Ru-Mg-Al catalysts ..................................... 93

Figure 4.8: (a) Total glycerol conversions and (b) liquid product concentrations obtained during

stability test over Ru-Mg-Al catalysts (flow rate: 0.075 mL/min, steam/glycerol molar ratio: 9) ..... 94

Figure 4.9: a) TG and b) DSC curves obtained during thermal treatment of the uncalcined Mg6Al2,

Mg6Al1.8La0.2, Mg6Al1.6La0.4 and Mg6Al1.2La0.8 supports ........................................................................ 98

Figure 4.10: XRD diffractograms of the uncalcined Mg6Al2, Mg6Al1.8La0.2, Mg6Al1.6La0.4 and

Mg6Al1.2La0.8 supports ............................................................................................................................. 100

Figure 4.11: XRD diffractograms of the calcined Mg6Al2, Mg6Al1.8La0.2, Mg6Al1.6La0.4 and

Mg6Al1.2La0.8 supports ............................................................................................................................. 102

Figure 4.12: Specific surface areas of the uncalcined and calcined Mg6Al2, Mg6Al1.8La0.2,

Mg6Al1.6La0.4, Mg6Al1.2La0.8 supports and calcined La2O3 ................................................................... 103

Figure 4.13: FTIR spectra of the uncalcined Mg6Al2, Mg6Al1.8La0.2, Mg6Al1.6La0.4 and Mg6Al1.2La0.8

supports .................................................................................................................................................... 104

Figure 4.14: FTIR spectra of the calcined Mg6Al2, Mg6Al1.8La0.2, Mg6Al1.6La0.4 and Mg6Al1.2La0.8

supports .................................................................................................................................................... 105

Figure 4.15: XRD diffractograms of the calcined Ni/Mg6Al2, Ni/Mg6Al1.8La0.2, Ni/Mg6Al1.6La0.4 and

Ni/Mg6Al1.2La0.8 solids ............................................................................................................................. 107

Figure 4.16: Comparison of specific surface areas of calcined supports (Mg6Al2, Mg6Al1.6La0.4 and

Mg6Al1.6Ce0.4) and Ni-impregnated solids after calcination (Ni/Mg6Al2, Ni/Mg6Al1.6La0.4 and

Ni/Mg6Al1.6Ce0.4) ...................................................................................................................................... 108

Figure 4.17: H2-TPR profiles of calcined Ni/Mg6Al2, Ni/Mg6Al1.8La0.2, Ni/Mg6Al1.6La0.4 and

Ni/Mg6Al1.2La0.8 solids ............................................................................................................................. 109

Figure 4.18: CO2-TPD profiles of calcined Ni/Mg6Al2, Ni/Mg6Al1.8La0.2, Ni/Mg6Al1.6La0.4 and

Ni/Mg6Al1.2La0.8 solids ............................................................................................................................. 111


selectivities of GSR reaction over Ni/Mg6Al2, Ni/Mg6Al1.8La0.2, Ni/Mg6Al1.6La0.4 and Ni/Mg6Al1.2La0.8

(flow rate: 0.025 mL/min, steam/glycerol molar ratio: 9) ................................................................... 113

Figure 4.20: a) TG and b) DSC curves obtained for used Ni/Mg6Al2, Ni/Mg6Al1.8La0.2,

Ni/Mg6Al1.6La0.4 and Ni/Mg6Al1.2La0.8 .................................................................................................... 116

Figure 4.21: XRD diffractograms of 5Ni/Mg6Al1.6La0.4, Ru-Ni/Mg6Al1.6La0.4 and 15Ni/Mg6Al1.6La0.4

.................................................................................................................................................................. 119

Figure 4.22: H2-TPR profiles of 5Ni/Mg6Al1.6La0.4, Ru-Ni/Mg6Al1.6La0.4 and 15Ni/Mg6Al1.6La0.4 .... 120

Figure 4.23: CO2-TPD profiles of 5Ni/Mg6Al1.6La0.4, Ru-Ni/Mg6Al1.6La0.4 and 15Ni/Mg6Al1.6La0.4 . 121


selectivities of GSR reaction over 5Ni/Mg6Al1.6La0.4, Ru-Ni/Mg6Al1.6La0.4 and 15Ni/Mg6Al1.6La0.4

(flow rate: 0.025 mL/min, steam/glycerol molar ratio: 9) ................................................................... 123

Figure 4.25: a) TG and b) DSC curves obtained for used 5Ni/Mg6Al1.6La0.4, Ru-Ni/Mg6Al1.6La0.4 and

15Ni/Mg6Al1.6La0.4 ................................................................................................................................... 125

Figure 4.26: (a) H2 yields (b) Glycerol conversion to gaseous products of GSR reaction over Ni/ZrO2

and Ni/Mg6Al1.6La0.4 (flow rate: 0.025ml/min, steam/glycerol molar ratio: 9) .................................. 127

Figure 4.27: (a) TG and (b) DSC curves obtained for used Ni/ZrO2 and Ni/Mg6Al1.6La0.4 .............. 128



over Ni/Mg6Al1.6La0.4 (flow rate: 0.025 mL/min, steam/glycerol molar ratio: 9) ............................... 129

Figure 4.29: TG and DSC curves obtained for used Ni/Mg6Al1.6La0.4 after stability test (flow rate:

0.025 mL/min) ......................................................................................................................................... 131



over Ni/Mg6Al1.6La0.4 (flow rate: 0.008 mL/min, steam/glycerol molar ratio: 9) ............................... 132

Figure 4.31: TG and DSC curves obtained for used Ni/Mg6Al1.6La0.4 after stability test (flow rate:

0.008 mL/min) ......................................................................................................................................... 134

LIST OF TABLES

Table 1.1: Physical properties of glycerol [27] ....................................................................................... 13

Table 1.2: Possible reactions in the GSR process ................................................................................... 20

Table 1.3: Pure metal oxide catalysts used in GSR ................................................................................ 27

Table 1.4: Hydrotalcite catalysts used in biodiesel production ............................................................. 31

Table 1.5: Hydrotalcites used in GSR ..................................................................................................... 31

Table 1.6: Types of carbon species formed by decomposition of CO on Ni [82] ................................. 34

Table 1.7: Carbon species formed in the steam reforming of hydrocarbons over Ni [82] ................. 35

Table 2.1: Physical and chemical properties of sunflower oil ............................................................... 38

Table 2.2: Summary of the prepared catalysts ....................................................................................... 40

Table 2.3: Specific surface areas of calcined catalysts ........................................................................... 43

Table 2.4: FAME profile produced from homogeneous catalysis ........................................................ 45

Table 2.5: Properties of produced biodiesel ........................................................................................... 52

Table 2.6: Properties of conventional petroleum diesel ......................................................................... 52

Table 3.1: Hydrogen consumption values of Ni impregnated catalysts ............................................... 61

Table 3.2: Basicity values of Ni impregnated catalysts .......................................................................... 62

Table 3.3: Hydrogen consumption values of Ni/ZrO2 (M) and Ni/ZrO2 (T) ........................................ 74

Table 3.4: Basicity values of Ni/ZrO2 (M) and Ni/ZrO2 (T) .................................................................. 75

Table 4.1: Lattice parameters and crystallite sizes of the uncalcined Ru-Mg-Al catalysts ................ 87

Table 4.2: Hydrogen consumption values of the calcined Ru-Mg-Al catalysts ................................... 89

Table 4.3: Basicity values of calcined Ru-Mg-Al catalysts .................................................................... 90

Table 4.4: Theoretical and experimental weight losses of the uncalcined Mg6Al2, Mg6Al1.8La0.2,

Mg6Al1.6La0.4 and Mg6Al1.2La0.8 supports .............................................................................................. 100

Table 4.5: Comparison of ionic radii and element electronegativities ............................................... 101

Table 4.6: Lattice parameters and crystallite sizes of the uncalcined Mg6Al2, Mg6Al1.8La0.2,

Mg6Al1.6La0.4 and Mg6Al1.2La0.8 supports .............................................................................................. 101

Table 4.7: Hydrogen consumption values of calcined Ni/Mg6Al2, Ni/Mg6Al1.8La0.2, Ni/Mg6Al1.6La0.4

and Ni/Mg6Al1.2La0.8 solids ..................................................................................................................... 110

Table 4.8: Basicity values of calcined Ni/Mg6Al2, Ni/Mg6Al1.8La0.2, Ni/Mg6Al1.6La0.4 and

Ni/Mg6Al1.2La0.8 solids ............................................................................................................................. 112

Table 4.9: Hydrogen consumption values of 5Ni/Mg6Al1.6La0.4, Ru-Ni/Mg6Al1.6La0.4 and

15Ni/Mg6Al1.6La0.4 ................................................................................................................................... 121

Table 4.10: Basicity values of 5Ni/Mg6Al1.6La0.4, Ru-Ni/Mg6Al1.6La0.4 and 15Ni/Mg6Al1.6La0.4 ........ 122

TABLE OF CONTENTS

General Introduction .................................................................................................................................. 1

Chapter 1: Literature Review .................................................................................................................... 4

1.1 Introduction ....................................................................................................................................... 5

1.2 Biodiesel production from vegetable oils ........................................................................................ 7

1.2.1 Chemistry of vegetable oils and fats ........................................................................................... 7

1.2.2 Transesterification ...................................................................................................................... 8

1.2.3 Undesired reactions in biodiesel production .............................................................................. 9

1.2.4 Parameters affecting biodiesel production ............................................................................... 10

1.2.5 Catalysts used in biodiesel production...................................................................................... 12

1.3 Glycerol as a by-product from biodiesel production ................................................................... 13

1.3.1 Properties of glycerol ................................................................................................................ 13

1.3.2 Uses and applications of glycerol ............................................................................................. 14

1.3.3 Glycerol steam reforming (GSR): a promising means of valorization .................................... 18

1.3.4 Thermodynamic studies ............................................................................................................ 19

1.3.5 Liquid Products ......................................................................................................................... 23

1.3.6 Catalysts used in GSR ............................................................................................................... 24

1.4 Choice of catalysts ........................................................................................................................... 28

1.5 Hydrotalcites or Anionic Clays ...................................................................................................... 29

1.5.1 Structure .................................................................................................................................... 29

1.5.2 Properties ................................................................................................................................... 30

1.5.3 Hydrotalcites used in biodiesel production .............................................................................. 30

1.5.4 Hydrotalcites used in GSR ........................................................................................................ 31

1.6 Catalyst Deactivation ...................................................................................................................... 32

1.7 Conclusion ....................................................................................................................................... 36

Chapter 2: Biodiesel production over Ca-based Mg-Al hydrotalcite derived catalysts ..................... 37

2.1 Characterization of vegetable oil ....................................................................................................... 38

2.2 Synthesis of catalysts ........................................................................................................................... 38

2.3 Characterization of catalysts.............................................................................................................. 40

2.3.1 X-ray diffractograms (XRD) of uncalcined and calcined Ca-Mg-Al catalysts ........................... 40

2.3.2 N2 sorption isotherms of calcined Ca-Mg-Al catalysts ................................................................ 42

2.3.3 CO2-Temperature Programmed Desorption (CO2-TPD) of calcined Ca-Mg-Al catalysts ......... 44

2.4 Biodiesel test results ............................................................................................................................ 45

2.4.1 Catalytic activity of homogeneous catalyst ................................................................................... 45

2.4.2 Catalytic activity of hydrotalcite derived mixed oxides ................................................................ 46

2.4.2.1 Influence of support ............................................................................................................... 46

2.4.2.2 Influence of impregnation ..................................................................................................... 47

2.5 Optimization ........................................................................................................................................ 49

2.5.1 Effect of catalyst to oil ratio .......................................................................................................... 49

2.5.2 Effect of alcohol to oil ratio .......................................................................................................... 50

2.5.3 Effect of reaction time ................................................................................................................... 50

2.5.4 Combined effect of alcohol to oil ratio and the reaction time ..................................................... 51

2.6 Characterization of produced biodiesel and comparison to petroleum diesel .............................. 52

2.7 Conclusion ........................................................................................................................................... 53

Chapter 3: Steam reforming of glycerol over Ni-based pure metal oxides .......................................... 54

3.1 Influence of the support on catalytic performance .......................................................................... 55

3.1.1 Synthesis of Ni-based pure metal oxides .................................................................................... 55

3.1.2 Characterization .......................................................................................................................... 55

3.1.2.1 X-ray diffractograms (XRD) of supports alone and Ni impregnated catalysts .................... 56

3.1.2.2 Specific surface area analysis (BET) of supports alone and Ni impregnated catalysts ...... 57

3.1.2.3 H2-Temperature Programmed Reduction (H2-TPR) of supports alone and Ni impregnated

catalysts............................................................................................................................................... 59

3.1.2.4 CO2-Temperature Programmed Desorption (CO2-TPD) of supports alone and Ni

impregnated catalysts ......................................................................................................................... 61

3.1.3 Steam reforming of glycerol results ........................................................................................... 62

3.1.3.1 The uncatalyzed reaction ....................................................................................................... 62

3.1.3.2 The catalyzed reaction: Steam reforming of glycerol over Ni/CeO2, Ni/Y2O3 and Ni/ZrO2 65

3.1.4 Liquid Product Analysis .............................................................................................................. 67

3.1.5 Characterization of used catalysts – Thermal analysis (TG-DSC) .......................................... 68

3.1.6 Discussion...................................................................................................................................... 70

3.2 Parameter Modification: Influence of zirconia phase on catalytic activity ................................... 70

3.2.1 Synthesis of monoclinic and tetragonal Ni/ZrO2 ....................................................................... 71

3.2.2 Characterization of calcined Ni/ZrO2 catalysts ......................................................................... 72

3.2.2.1 X-ray diffractograms (XRD) of calcined Ni/ZrO2 catalysts .................................................. 72

3.2.2.2 Specific surface area analysis (BET) of calcined Ni/ZrO2 catalysts .................................... 72

3.2.2.3 H2-Temperature Programmed Reduction (H2-TPR) of calcined Ni/ZrO2 catalysts ............ 73

3.2.2.4 CO2-Temperature Programmed Desorption (CO2-TPD) of calcined Ni/ZrO2 catalysts ...... 74


3.2.4 Liquid Product Analysis .............................................................................................................. 77


3.2.6 Stability Test ................................................................................................................................. 79

3.2.7 Discussion...................................................................................................................................... 82

3.3 Conclusion ........................................................................................................................................... 83

Chapter 4: Steam reforming of glycerol over Ru or Ni-based Mg-Al hydrotalcite derived catalysts

.................................................................................................................................................................... 84

4.1 Role of the preparation method in catalytic activity of Ru-Mg-Al catalysts ................................. 85

4.1.1 Synthesis of Ru-Mg-Al catalysts ................................................................................................. 85

4.1.1.1 Synthesis of Ru Grafted catalyst (RuG/Mg6Al2) ................................................................... 85

4.1.1.2 Synthesis of Ru Impregnated catalyst (RuI/Mg6Al2) ............................................................ 86

4.1.2 Characterization of Ru-Mg-Al catalysts .................................................................................... 86

4.1.2.1 X-ray diffractograms (XRD) of uncalcined and calcined Ru-Mg-Al catalysts .................... 86

4.1.2.2 Specific surface area analysis (BET) of calcined Ru-Mg-Al catalysts................................. 88

4.1.2.3 H2-Temperature Programmed Reduction (H2-TPR) of calcined Ru-Mg-Al catalysts ......... 88

4.1.2.4 CO2-Temperature Programmed Desorption (CO2-TPD) of calcined Ru-Mg-Al catalysts .. 90



4.1.5 Stability Test ................................................................................................................................. 94

4.1.6 Discussion...................................................................................................................................... 95

4.2 Influence of support modification on the catalytic activity of Ni-Mg-Al catalysts ........................ 97

4.2.1 Synthesis of supports ................................................................................................................... 97

4.2.2 Characterization of supports ...................................................................................................... 98

4.2.2.1 Thermal analysis (TG-DSC) of uncalcined supports ............................................................ 98

4.2.2.2 X-ray diffractograms (XRD) of uncalcined and calcined supports .................................... 100

4.2.2.3 Specific surface area analysis (BET) of uncalcined and calcined supports ...................... 103

4.2.2.4 Fourier Transform Infrared Spectroscopy (FTIR) study of uncalcined and calcined

supports............................................................................................................................................. 104

4.2.2.5 Discussion ............................................................................................................................. 106

4.2.3 Synthesis of Ni-based hydrotalcite derived mixed oxides ....................................................... 106

4.2.4 Characterization of Ni-impregnated supports ........................................................................ 107

4.2.4.1 X-ray diffractograms (XRD) of calcined Ni-impregnated supports ................................... 107

4.2.4.2 Specific surface area analysis (BET) of Ni-impregnated supports .................................... 108

4.2.4.3 H2-Temperature Programmed Reduction (H2-TPR) of Ni-impregnated supports ............ 109

4.2.4.4 CO2-Temperature Programmed Desorption (CO2-TPD) of Ni-impregnated supports ...... 111

4.2.5 Steam reforming of glycerol over Ni/Mg6Al2, Ni/Mg6Al1.8La0.2, Ni/Mg6Al1.6La0.4 and

Ni/Mg6Al1.2La0.8 ................................................................................................................................... 113

4.2.6 Liquid Product Analysis ............................................................................................................ 115

4.2.7 Characterization of used catalysts – Thermal analysis (TG-DSC) ........................................ 116

4.2.8 Discussion.................................................................................................................................... 117

4.3 Parameter Modification: Effect of a higher Ni loading (15%Ni) and bimetallic effect (1%Ru-

5%Ni) ....................................................................................................................................................... 119

4.3.1 Synthesis...................................................................................................................................... 119

4.3.2 Characterization ........................................................................................................................ 119

4.3.2.1 X-ray diffractograms (XRD) ................................................................................................ 119

4.3.2.2 H2-Temperature Programmed Reduction (H2-TPR) .......................................................... 120

4.3.2.3 CO2-Temperature Programmed Desorption (CO2-TPD) .................................................... 121

4.3.3 Steam reforming of glycerol results ......................................................................................... 123

4.3.4 Liquid Product Analysis ............................................................................................................ 125

4.3.5 Characterization of used catalysts – Thermal analysis (TG-DSC) ........................................ 125

4.3.6 Discussion.................................................................................................................................... 126

4.4 Comparison to Ni-based pure metal oxide...................................................................................... 127

4.5 Stability Test ...................................................................................................................................... 129

4.6 Optimization ...................................................................................................................................... 132

4.7 Conclusion ......................................................................................................................................... 134

Conclusion and Perspectives .................................................................................................................. 136

Appendix A: Biodiesel ............................................................................................................................ 142

Appendix B: Glycerol Steam Reforming .............................................................................................. 156

References ................................................................................................................................................ 172

Dissemination of Research Results ........................................................................................................ 183

1

General Introduction

2

Most of the current world energy demands are derived from fossil fuels. However, growing

concerns with issues such as global warming, population growth and limited fuel reserves have

prompted researchers to search for other alternative fuels. The use of renewable biofuels is a

promising solution as they have the potential to act as substitutes for fossil fuels and are

additionally carbon neutral. Biodiesel, in particular, can be produced by the transesterification of

vegetable oils with an alcohol in the presence of a catalyst. Glycerol is obtained as a by-product of

this reaction. With the rising demand for biodiesel production, glycerol production is also

increasing dramatically. Given the limited uses of glycerol, these excess quantities of glycerol

from biodiesel production could become a waste problem, as the international market for glycerol

is small compared to the produced quantities. Nevertheless, this glycerol can be valorized by using

it to produce hydrogen. Therefore, the excess of glycerol produced from the biodiesel industry can

be considered as a potential feedstock for renewable hydrogen production. Although various

glycerol reforming technologies are available, the steam reforming process is the most widely used

as well as the most competitive method for hydrogen production. The steam reforming of glycerol

combines steam with glycerol in the presence of a catalyst to produce hydrogen and carbon

dioxide. The presence of a catalyst is essential to both the biodiesel and steam reforming of

glycerol reactions, as it helps increase the reaction rate as well as control the selectivity towards

the formation of the desired products. An ideal catalyst should not only be active but should also

be stable and cheap. Since a catalyst is typically composed of both a support and an active phase,

the choice of both components is crucial when seeking the ideal catalyst. This work aims to design

adequate catalysts that will make the biodiesel production and glycerol steam reforming reactions

economically viable. Literature review showed that catalysts with basic properties are highly

preferred for both reactions. For biodiesel production, basic catalysts favor higher Fatty Acid

Methyl Ester (FAME) yields. For glycerol steam reforming, basic catalysts help limit coke

formation. Accordingly, Mg-Al hydrotalcites were chosen as the support materials for both

applications. Hydrotalcites are known for their low cost, high thermal stability and high basicity.

Active phases suitable to each reaction were also selected. Calcium was selected for the biodiesel

reaction whereas nickel and ruthenium were selected for the glycerol steam reforming reaction.

All prepared catalysts were characterized using various techniques and were then tested in their

respective reactions. Accordingly, this manuscript is divided into 4 chapters:

3

The first chapter is a literature review that covers both the production of biodiesel via

transesterification and the steam reforming of glycerol reactions. The catalysts used in each

reaction are studied and a description of hydrotalcites and their properties is also included.

The second chapter focuses on the use of Ca-based Mg-Al hydrotalcite derived mixed

oxides for biodiesel production. The effects of thermal treatment and the preparation

method for Ca incorporation (impregnation vs grafting) were studied. Optimization of the

reaction conditions was conducted on the most performing catalyst.

The third chapter focuses on the use of Ni-based pure metal oxides for the glycerol steam

reforming reaction as pure metal oxides are more commonly used as catalytic support

materials. The effect of utilizing different supports (CeO2, Y2O3 and ZrO2) was studied.

The effect of the crystalline phase of the most efficient Ni-based pure metal oxide was then

considered followed by stability tests. The most efficient catalyst in this series was selected

to later on be compared to the most efficient catalyst from the hydrotalcite series in Chapter

4.

The fourth chapter first focuses on the use of Ru-based Mg-Al hydrotalcite derived mixed

oxides for the glycerol steam reforming reaction. The effect of the preparation method was

studied (impregnation vs grafting). This is followed by the study of the use of Ni-based

Mg-Al hydrotalcite derived mixed oxides. In the case of the Ni-based catalysts, the Mg-Al

supports were also modified with La to additionally study the effect of promoter addition.

The bimetallic effect (1%Ru-5%Ni) and effect of a higher Ni loading (15 wt% Ni) were

also studied on the most efficient Ni-based hydrotalcite derived catalyst. The activity of

the most efficient hydrotalcite derived catalyst was then compared to the most efficient

pure metal oxide catalyst from Chapter 3. Stability tests and an optimization of the reaction

conditions were finally conducted on the most efficient hydrotalcite derived catalyst.

It is worth noting that for Chapters 2, 3 and 4, the synthesis, physico-chemical characterization

results, test results and optimization steps are all included in each chapter.

Finally, a general conclusion summarizes the main results obtained in this thesis as well as some

perspectives. Two appendices (Appendix A and Appendix B) detail the characterization

techniques, catalytic test descriptions and supplementary results for the biodiesel and glycerol

steam reforming chapters respectively.

4

Chapter 1: Literature Review

5

This chapter begins with an introduction that provides the background of the study. This is

followed by a detailed explanation of the production of biodiesel from vegetable oils along with

the parameters that affect biodiesel production as well as the catalysts commonly used. The

properties and uses of the by-product glycerol are then highlighted followed by the various

methods of hydrogen production from glycerol. Since the steam reforming reaction will be the

focus of this study, the parameters, thermodynamics and catalysts used in the steam reforming

reaction are detailed. Based on this literature review, the choice of catalysts for both reactions was

then explained. Finally, a description of hydrotalcites and their properties along with the different

types of catalyst deactivation is also provided.

1.1 Introduction

As a result of an ever-growing population, diminishing fossil fuel reserves and increasing

environmental concerns associated with fossil fuel utilization, the use of alternative energy

resources has become increasingly important [1],[2]. The search for alternative fuels that are

cheaper, less polluting, more efficient and most importantly that can act as a substitute for

conventional fossil fuels has recently led to an intense focus on bio-based fuels [2]. Biofuels or

bio-based fuels, by definition, are sources of energy obtained from biomass (crops, trees, animal

fats, etc.). In fact, these fuels are emerging as a long-term solution as they are renewable and

theoretically considered to be carbon dioxide neutral [3]. As shown in Figure 1.1, the burning of

conventional diesel originating from fossil fuels leads to a CO2 build up in the atmosphere, whereas

the atmospheric CO2 emitted from the burning of a biofuel is recycled via photosynthesis of the

crops used to produce the biofuel.

Figure 1.1: CO2 emission cycles of conventional diesel and a biofuel [3]

6

The demand and production of biodiesel in particular has increased dramatically over the last few

years [1]. Biodiesel is non-toxic, biodegradable and emits less pollutants (CO, SO2 and

hydrocarbons) compared to conventional diesel [4]. Biodiesel also has a higher cetane number,

lower sulfur content and a higher flash point [5]. However, in terms of price, biodiesel is not yet

competitive. The cost of biodiesel is 1.5-3 times the price of conventional diesel [6]. One way of

lowering the production cost of biodiesel would be the valorization of its main by-product glycerol

to provide an extra value [2]. This will be discussed in detail in the coming sections.

Based on the international energy statistics conducted by the U.S. Energy Information

Administration (EIA), Figure 1.2 shows the worldwide biodiesel production from the year 2010 to

2016 in terms of barrels per day. The increasing trend in recent years is clearly observed as more

focus is being placed on this renewable energy source. Figure 1.3 shows the leading biodiesel

producing countries in 2016. The United States of America, Brazil, Indonesia, Germany,

Argentina, France and Spain were the leading countries in 2016 [7].

2010 2011 2012 2013 2014 2015 20160

100

200

300

400

500

600

700

800

586

526

571529

467426

Bio

die

se

l P

rod

uctio

n (

Th

ou

sa

nd

Ba

rre

ls P

er

Da

y)

Others

North America

Europe

Central & South America

Asia

357

Figure 1.2: Worldwide biodiesel production from 2010-2016 [7]

7

0

10

20

30

40

50

60

70

80

90

100

110

Sp

ain

Fra

nc

e

Arg

en

tin

a

Ge

rma

ny

Ind

on

es

ia

Bra

zil

23

33

48

6063

66

Th

ou

sa

nd

Ba

rre

ls P

er

Da

y

104

Un

ite

d S

tate

s

Figure 1.3: Leading biodiesel producing countries of 2016 [7]

1.2 Biodiesel production from vegetable oils

Biodiesel can be produced via the transesterification of triglycerides with an alcohol in the

presence of a catalyst [8]. These triglycerides can be extracted from vegetable oils such as

sunflower oils, soybean and rapeseed (first generation biodiesel feedstock) [3]. Second generation

feedstock include waste cooking oil, grease and animal fats. In recent years, microalgae are being

used as a third generation biodiesel feedstock [9]. When the utilized alcohol is methanol, the

produced biodiesel is known as Fatty Acid Methyl Esters or FAME.

1.2.1 Chemistry of vegetable oils and fats

Fats and oils both belong to the same biological group of “lipids”. The main difference is that fats

are solid at room temperature whereas oils are liquid. Both have the same molecular structure

known as a triglyceride or triacylglycerol (TAG). Triglycerides are molecules in which each of the

three fatty acid molecules is linked to one of the three OH groups of a glycerol molecule by means

of an ester bond (Figure 1.4).

8

Figure 1.4: Triglyceride molecule [10]

Free fatty acids (FFA) can be defined as fatty acids that are no longer present inside the triglyceride

molecule due to exposure of the oil to heat, oxygen and/or moisture. Exposure to moisture or water

can cause hydrolysis and the cracking of oils to FFAs which will be discussed in section 1.2.3.

1.2.2 Transesterification

The main transesterification reaction is represented in Figure 1.5. As shown in this figure,

according to stoichiometry, three moles of methanol are required to convert one mole of

triglyceride into three moles of fatty acid methyl esters (FAME, biodiesel) and one mole of

glycerol as a by-product.

Figure 1.5: Overall transesterification reaction of triglycerides with methanol [11]

Figure 1.6 shows that the transesterification reaction actually occurs in three consecutive steps to

obtain the final required biodiesel. In the first step, the triglyceride (three fatty acids attached to

the glycerol skeleton via ester bonds) is converted to a diglyceride (two fatty acids attached to the

glycerol skeleton) and one FAME. In the second step, the diglyceride is converted to a

9

monoglyceride and another FAME is liberated. Finally, in the third step, the monoglyceride is

converted into the third FAME and glycerol is also obtained.

Figure 1.6: Transesterification reaction steps [12]

1.2.3 Undesired reactions in biodiesel production

Biodiesel synthesis ideally requires reactants free of water or at least within an acceptable range.

Moreover, the oil utilized should have a limited FFA content (0.1-0.5 wt%) [5], [13]. Triglycerides

in the presence of water undergo hydrolysis leading to the formation of FFA (Figure 1.7).

Figure 1.7: Hydrolysis of triglycerides [11]

10

A high quantity of FFAs could decrease the biodiesel yield in the presence of basic catalysts (e.g.

NaOH) due to the formation of undesired products such as soap and water as shown in Figure 1.8

[9]. The produced water could further react with the triglycerides in the reaction medium according

to Figure 1.7 giving more FFAs which in turn produce more soap and water and the cycle goes on.

The removal of these undesired products requires additional separation costs.

Figure 1.8: Saponification of FFA [14]

1.2.4 Parameters affecting biodiesel production

In order to increase the yield of biodiesel, there are various parameters in the transesterification

reaction:

Reaction Temperature

Several works have demonstrated that the temperature has a significant effect on the reaction rate

[15]. However, the choice of the exact reaction temperature actually depends on the alcohol used.

The reaction temperature should be kept below the boiling point of the reacting alcohol being used

to avoid its loss during the reaction. For example, transesterification using methanol should be

conducted at a temperature of around 60oC-65oC to avoid exceeding the boiling point of methanol

[16].

Choice of alcohol and alcohol to oil molar ratio

The type of alcohol used in the transesterification reaction affects the quality and cost of biodiesel

production. Higher chain alcohols generally result in a higher viscosity and a higher production

cost [13]. Methanol is mostly used as it is relatively cheap and readily available in addition to

having a short chain [4]. As shown in Figure 1.5, for the transesterification reaction to occur, an

alcohol to oil molar ratio of 3:1 is required by the stoichiometry. Since the main transesterification

reaction is reversible, a higher alcohol to oil molar ratio (alcohol in excess) will shift the

equilibrium to the right thus favoring the formation of FAME. However, excessive addition of

11

methanol could also have a negative effect as it could dilute the oil concentration and decrease the

availability of active sites thus reducing the reaction rate. Moreover, excess alcohol could enhance

emulsification which makes the separation of the produced biodiesel and glycerol difficult.

Therefore, the optimum alcohol to oil ratio should be determined experimentally as it depends on

the conditions and catalyst used.

Amount of catalyst

The quantity of catalyst used significantly affects the FAME yield. An insufficient amount of

catalyst could result in an incomplete reaction. Larger quantities of catalyst increase the rate of

ester formation. For example, when using a basic catalyst, as the catalyst weight increases, the

number of O2- anion sites increases which adsorb more H+ from methanol forming active centers.

The reactants will have more contact with these active centers thus increasing the yield. However,

in some cases an excess of catalyst could decrease the yield. Excess quantities of an alkali catalyst

could for example form more soap if the reactant oil has a high quantity of FFAs.

Amount of water

As mentioned previously, the presence of water in the reaction medium favors the formation of

FFAs which can lower the biodiesel yield. Canakci and Gerpen [17] demonstrated that a water

content as little as 0.5 wt% can affect the biodiesel yield and a content of 5 wt% can totally inhibit

the reaction.

Reaction time

Homogeneous catalysis for biodiesel production can be conducted in less than an hour. On the

other hand, heterogeneous catalysis for biodiesel production require a longer reaction time due to

mass transfer and diffusion problems between the oil and catalyst during the early stages of the

reaction [15].

Based on all the above and more in depth literature study, the following conditions were chosen

as a starting point for the current study: a reaction temperature of 60oC, 12:1 methanol to oil molar

ratio, 2.5 wt% catalyst to oil ratio and a reaction time of 4 hours. The optimum conditions for

biodiesel production will then be determined experimentally by varying these parameters.

12

1.2.5 Catalysts used in biodiesel production

The choice of an appropriate catalyst is necessary for the design of a competitive and sustainable

transesterification process [6]. Various types of catalysts have been investigated for biodiesel

production and have been classified as either homogeneous (liquid) or heterogeneous (solid).

Homogeneous catalysts [e.g. NaOH (base catalyst) and H2SO4 (acid catalyst)] are commonly used

in industries because of their high activity and low cost [13]. However, they are known to have

several drawbacks:

Contamination problems; homogeneous catalysts favor the occurrence of side reactions

including saponification with basic catalysts which later causes difficulty in the separation

and purification of the obtained biodiesel [18]. As discussed previously, FFA present in

the triglycerides can undergo saponification with homogeneous basic catalysts according

to Figure 1.8.

Formation of large quantities of wastewater; in order to remove the soap generated from

saponification reactions, thorough washing steps are required and this results in the

production of large volumes of acidic wastewater [19].

Lack of catalyst reusability; as these catalysts are soluble, their recovery from the obtained

products is time consuming and difficult which prevents their reuse.

Equipment corrosion [20].

Therefore, more focus is being placed on the use of heterogeneous solid catalysts as they are more

environmentally friendly and allow easier separation of the obtained products. They can also be

easily recovered, regenerated, reused and are non-corrosive [21].

Although the use of heterogeneous acid catalysts has the advantage of having no undesired side

reactions, the low catalytic activity of these catalysts necessitates using higher reaction

temperatures and reaction times. Heterogeneous base catalysts, on the other hand, when given a

suitable feedstock have a lower reaction temperature requirement as well as a lower reaction time

[9]. This makes them more interesting candidates for biodiesel production. Basic catalysts such as

resins, clays, metal oxides and lipases are mainly explored as basic properties favor a higher FAME

yield [22]. Recently, many studies have used CaO in particular as a heterogeneous base catalyst

for biodiesel production because of its low production cost and high basicity [23]–[25].

13

1.3 Glycerol as a by-product from biodiesel production

As shown in Figure 1.5, glycerol is obtained as a main by-product of the transesterification

reaction. In fact, 10% of the produced biodiesel quantity is glycerol, i.e. for every ton of biodiesel

produced, 100kg of glycerol are also produced [26].

1.3.1 Properties of glycerol

Glycerol, also known as propane 1,2,3 triol (IUPAC), is a trihydric alcohol as shown in Figure 1.9.

Commercially, it is also known as glycerin, glycerine, glyceritol or glycyl alcohol. Glycerol can

be classified as either natural or synthetic. Natural glycerol is produced by the transesterification

process in biodiesel synthesis (focus of this study), high-pressure splitting of water and fat, and

saponification in soap manufacturing. Synthetic glycerol is produced from propylene oxidation

and chlorination [27].

Glycerol is a colorless, odorless, viscous and sweet tasting liquid. It is non-toxic, hygroscopic and

classified as a polar protic solvent. Furthermore, it is completely soluble in water and short-chain

alcohols, partially soluble in organic solvents and completely insoluble in hydrocarbons [28]. The

three hydroxyl groups in its structure (Figure 1.9) are responsible for its solubility in water as well

as its hygroscopic nature. The physical properties of glycerol are summarized in Table 1.1.

Table 1.1: Physical properties of glycerol [29]

Properties

Chemical formula C3H8O3

Molecular Weight (g.mol-1) 92.1

Liquid density @293K (cm3.g-1) 1.26

Freezing Point (K) 291

Boiling Point (K) 563

Critical Temperature (K) 726

Critical Pressure (atm) 66

Critical Volume (cm3.mol-1) 255

Critical Compressibility (Zc) 0.28

Figure 1.9: Glycerol molecule

14

For industrial applications and this work, it is important to study the properties of binary glycerol-

water mixtures because these mixtures must be vaporized during the reaction. The boiling point of

pure glycerol at atmospheric pressure is 290oC. The addition of water reduces the boiling point of

the obtained solution. Figure 1.10 shows the variation of the boiling point (starting from T= 100oC)

and the variation of the vapor pressure (starting from P=760 mmHg) as the glycerol-water

composition is changed. As the weight percentage of glycerol increases in the solution, the vapor

pressure decreases while the boiling point increases. The reduction in vapor pressure becomes

more pronounced beyond 60 wt% where there is also a steep increase in the boiling point. The

addition of as little as 10 wt% of water causes a decrease in the boiling point from 290oC to 138oC.

Glycerol molecules form intermolecular hydrogen bonds with water molecules which could also

affect intramolecular hydrogen bonds in the glycerol molecules thereby explaining the lower

boiling points.

0 20 40 60 80 100

100

150

200

250

300

Boiling point

Vapor pressure

Wt% Glycerol

Boili

ng p

oin

t (o

C)

0

100

200

300

400

500

600

700

800

Vapor p

ressure

(mm

Hg)

Figure 1.10: Boiling points and vapor pressures of glycerol solutions [30]

1.3.2 Uses and applications of glycerol

Glycerol is a versatile chemical used in numerous applications due to its unique combination of

properties. The main uses are shown in Figure 1.11. Almost two thirds of glycerol industrial uses

are in food and beverage (23%), oral care (20%), cosmetics (13%) and tobacco (12%) [1].

15

Figure 1.11: Uses of glycerol

Due to the increased production of biodiesel in recent years, the quantity of glycerol produced has

also risen. This glut of glycerol in the world market necessitates the need for alternative uses [31].

One promising way is to use glycerol as a potential feedstock for renewable hydrogen production.

With advancements in the fuel cell industry, the demand for hydrogen continues to grow [32]. In

fact, hydrogen can be produced from glycerol by several ways [1]:

Pyrolysis

Partial oxidation gasification process

Auto-thermal reforming process

Aqueous-phase reforming process

Supercritical water reforming process

Steam reforming process

Pyrolysis

Glycerol pyrolysis is the thermal cracking of glycerol in an oxygen-free environment as shown in

Equation (1):

C3H8O3 3CO + 4H2 ΔH298K = 251 kJ/mol (1)

Uses of Glycerol

Foods (23%) Oral Care (20%) Cosmetics (13%) Tobacco (12%)

Drugs (9%) Miscellaneous (9%) Plastics (8%) Alkyds (3%)

Meat casing (2%) Explosive (1%)

16

As the reaction is highly endothermic, it requires a considerable input of heat. This leads to the

formation of various products via dehydration and dehydrogenation reactions. These species can

further be transformed to intermediates which then decompose to form syngas (H2, CO). Figure

1.12 shows the possible reactions that can occur in glycerol pyrolysis [27].

Figure 1.12: Possible reactions for glycerol pyrolysis [27]

Partial oxidation

Partial oxidation involves the breakdown of glycerol with oxygen. The reaction can be expressed

as shown in Equation (2):

C3H8O3 + 3/2O2 3CO2 + 4H2 ΔH298K = -598 kJ/mol (2)

By varying the stoichiometric number of oxygen inputted, syngas can also be produced via the

following general equation (3):

C3H8O3 + (3-x)/2O2 xCO + (3-x)CO2 + 4H2 ΔH298K = -598 + 283 * x kJ/mol (3)

17

Due to the exothermic nature of the reaction, the high temperatures and rapid oxygen consumption

rate provoke several side reactions [27].

Auto-thermal reforming

Auto-thermal reforming combines steam reforming with partial oxidation. In auto-thermal

reforming, water is co-fed with glycerol under partial oxidation conditions [Equation (4)]:

C3H8O3 + 3/4O2 + 3/2H2O 3CO2 + 11/2H2 ΔH298K = -240 kJ/mol (4)

Since it is a combined process, the endothermic steam reforming process absorbs the heat

generated by the exothermic partial oxidation process over the same catalyst bed [1]. Therefore, a

constant supply of expensive pure oxygen is required to maintain a temperature that is high enough

to drive the reaction [33].

Aqueous-phase reforming (APR)

Aqueous-phase reforming, also known as hydrothermal reforming, is when the reforming process

is conducted in an aqueous phase. This method usually operates in excess water content, lower

temperatures (200-300oC) and high pressures (up to 60 bars) [34].

Glycerol has a higher boiling point (290oC) than other hydrocarbon feedstock and undergoes

thermal decomposition near this temperature. Thus, glycerol vaporization not only requires

thermal input but also degrades the feedstock. To deal with this issue, Dumesic et al. [35] were the

first to use the method of aqueous phase reforming where the glycerol was used in aqueous phase

without the pre-vaporization step required in other methods. High purity hydrogen can be produced

with this process, however it is still relatively new [27].

Supercritical water reforming

Supercritical water can be defined as water that is heated and compressed at its critical temperature

(374oC) and pressure (22.1 MPa) [1]. Some authors have reported hydrogen production from the

reforming of glycerol with supercritical water [36], [37]. Nevertheless, it is still an emergent

technology like aqueous phase reforming (APR) and in the developing stage. In addition, the high

pressures involved make it less attractive for industrial application given the high costs it could

inflict.

18

Steam reforming of glycerol

The steam reforming of glycerol (Equation 5) is an equilibrium-limited process that involves the

decomposition or pyrolysis of glycerol (Equation 6) followed by the water gas shift reaction

(WGS) (Equation 7):

C3H8O3 + 3H2O 7H2 + 3CO2 ΔH298K = +128 kJ/mol (5)

C3H8O3 4H2 + 3CO ΔH298K = +251 kJ/mol (6)

CO + H2O CO2 + H2 ΔH298K = -41 kJ/mol (7)

The focus of this work will be on the steam reforming of glycerol.

1.3.3 Glycerol steam reforming (GSR): a promising means of valorization

The steam reforming of glycerol as a method of glycerol valorization is preferred over the other

methods for several reasons:

It can be conducted at atmospheric pressure [38].

The steam reforming process is the most widely used amongst all the existing reforming

processes for hydrogen production. Nearly 48% of worldwide produced hydrogen is

generated by the steam reforming of methane followed by the reforming of naptha/oil

(30%) [3]. Therefore, industries are already equipped to accomodate steam reforming

conditions. All that remains would be to change the feedstock from methane or naptha to

glycerol.

The steam reforming of methane, stoichiometrically speaking, yields up to 4 moles of

hydrogen whereas the steam reforming of glycerol yields up to 7 moles of hydrogen.

Therefore, the use of glycerol as a feedstock could be both promising and advantageous for

industrial applications. However, the GSR process, like any other process, faces several obstacles

that must be first overcome in order to establish effective commercialization on an industrial scale.

One of these obstacles is that the main reaction is often accompanied by various side reactions that

affect the production and purity of hydrogen [2]. Coke formation is also a major problem because

it leads to the deactivation of the catalyst used (discussed in detail in section 1.6). In order to solve

these problems, it is therefore crucial to control operating conditions and to design new catalysts

that can maximize hydrogen production [3]. These catalysts should not only be active but also

stable and cheap. In the upcoming sections, thermodynamic studies will help select suitable

19

operating conditions that will reduce the influence of side reactions as well as coke deposition. A

literature study on tested catalysts in the GSR reaction will also help select the active phase and

support material for the design of an appropriate catalyst.

1.3.4 Thermodynamic studies

Few thermodynamic studies are available for GSR as it is a relatively new process and still under

study. Nevertheless, the aim of these studies is to find the effect of operating conditions on

hydrogen production. Several parameters control the GSR reaction system:

Temperature

Reactant feed ratios (water / glycerol feed ratios)

Pressure

One of the most important parameters is temperature. The decomposition of glycerol is highly

endothermic (Equation 6) and requires more heat than that produced by the water gas shift reaction

(Equation 7) thereby making the overall GSR reaction endothermic (ΔH > 0). As a result of this

endothermic nature, higher temperatures are required to induce H2 production. Unfortunately, high

temperatures are not favorable economically as they result in high operating costs as well as the

need for more resistant reactors that can withstand such temperatures. This poses a challenge for

effective commercialization of the process. The water / glycerol feed ratio (WGFR) also highly

influences the production of hydrogen. Higher water to glycerol feed ratios are required to shift

the equilibrium to the right towards the consumption of excess water and production of more

hydrogen. This behavior has been studied and verified: hydrogen yield increases as the WGFR

increases from 4-15 whatever the temperature. It has also been particularly noted that for WGFR

greater than 9, H2 yield increases but at a slower rate especially at high temperatures. Nevertheless,

the amount of water used should not be too high due to the excessive vaporization costs it could

inflict for industrial application. Pressure also has an influence on the reaction system especially

because there is a variation in the number of moles of the GSR reaction which affects the

thermodynamic equilibrium. Higher pressures result in lower hydrogen production. Therefore,

lower pressures are more favorable as they lead to higher hydrogen production under equilibrium.

Additionally, lower pressures are more economically attractive as long as the pressure is not too

low. In summary, high temperatures, high WGFRs and low pressures favor H2 production [3].

Taking all these factors into consideration, the optimum conditions reported in literature [39]–[41]

are: temperatures between 580 and 700oC, WGFR between 9 and 12, and atmospheric pressure.

20

To reduce the influence of side reactions that could affect hydrogen production, it is important to

study all the possible reactions in the GSR process and at what temperature ranges these reactions

are favored. All compounds in glycerol steam reforming result from an atomic combination of

carbon, hydrogen and oxygen [42]. In fact, the list of possible species involved in the reaction

system is quite large. Therefore, it is necessary to determine only the most important ones in order

to facilitate both thermodynamic calculation and analysis. Compounds with more than three carbon

atoms are highly unlikely to exist in the glycerol steam reforming reaction as the Gibbs free energy

of formation increases with the number of carbon atoms [43]. Furthermore, previous studies that

have taken into account all possible methyl, ethyl and propyl compounds have found that under

typical reaction conditions, compounds such as methanal, methanol, ethylene, ethane, ethanal,

ethanol, propane, propene, propanal and propanone give negligible concentrations as demonstrated

by mole fractions below 10-6 [42]. In this work, seven main species were taken into consideration.

These include six gaseous species; glycerol (C3H8O3), water (H2O), hydrogen (H2), carbon dioxide

(CO2), carbon monoxide (CO), methane (CH4) and one solid; carbon (C). Based on these main

species, the following possible reactions were studied in detail in Table 1.2.

Table 1.2: Possible reactions in the GSR process

Reaction ΔH298

(kJ/mol)

C3H8O3 + 3H2O 7H2 + 3CO2 128 Glycerol steam reforming (5)

C3H8O3 4H2 + 3CO 251 Glycerol decomposition (6)

CO + H2O CO2 + H2 -41 Water gas shift (7)

CO2 + H2 CO + H2O 41 Reverse water gas shift (8)

CO + 3H2 CH4 + H2O -206 Methanation of CO (9)

CO2 + 4H2 CH4 + 2H2O -165 Methanation of CO2 (10)

C3H8O3 + 2H2 2CH4 + CO + 2H2O -160 Hydrogenolysis of glycerol (11)

CO2 + CH4 2H2 + 2CO 261 Methane dry reforming (12)

CH4 + H2O CO + 3H2 206 Methane steam reforming (13)

2CO CO2 + C -172 CO disproportionation reaction (14)

CH4 2H2 + C 75 Methane decomposition (15)

CO + H2 C + H2O -131 Carbon formation (16)

CO2 + 2H2 C + 2H2O -90 Carbon formation (17)

21

Figure 1.13 shows the calculated equilibrium constant Kp as a function of temperature for all the

possible reactions in the glycerol steam reforming system. If the Kp values of a reaction increase

with temperature, this reaction becomes more favored. Whereas if the Kp values decrease with

increasing temperature, the reaction becomes less favored.

200 300 400 500 600 700 800-20

-10

0

10

20

30

40Log

10K

p

Temperature (oC)

Figure 1.13: Equilibrium constants of possible reactions in GSR

The glycerol steam reforming and glycerol decomposition reactions (reactions 5 and 6) both have

the largest Kp values (1017-1026 and 1010-1026) in the studied temperature range (200-800oC). This

indicates that the equilibrium is shifted towards the formation of the products H2, CO2 and CO as

the reaction temperature increases.

Several side reactions also compete with the main reforming reaction. Methanation reactions

(reactions 9, 10 and 11), that lead to the production of methane, are thermodynamically feasible in

the studied temperature range. However, their Kp values decrease with increasing temperatures.

Methane steam and dry reforming reactions (reactions 12 and 13 respectively) are both

thermodynamically unfeasible, at lower temperatures, as demonstrated by their low Kp values.

However, at higher temperatures, these methane consuming reactions become favored.

This also applies to the carbon forming methane decomposition reaction (reaction 15), whereas

the Kp values of all the other carbon formation reactions (reactions 14, 16 and 17) decrease with

increasing temperatures.

22

Therefore, to inhibit coke formation and CH4 formation, the following conditions were chosen for

this study:

Temperature range of 400-700oC

Water to glycerol feed ratio of 9

Atmospheric pressure.

Wang et al. [44] generated two figures demonstrating the thermodynamic equilibrium of GSR

under the same conditions as those selected for this study. Figure 1.14(a) displays the glycerol and

steam conversions, while Figure 1.14(b) shows the variation of the number of moles of all products

(H2, CO2, CO and CH4) as a function of temperature.

Figure 1.14: Thermodynamic equilibrium calculations (T:400-700oC and WGFR 9:1) [44]

Figure 1.14(a) shows that thermodynamically, a total glycerol conversion of 100% is attained at

around 600oC. Figure 1.14(b) shows that at low temperatures, the formation of the products H2,

CO2 and CH4 is observed as a result of the water gas shift and methane formation reactions as they

23

are both exothermic reactions that are favored at low temperatures. Whereas at higher

temperatures, there is an increase in H2 and CO and a small decrease in CO2 due to the occurrence

of the endothermic glycerol steam reforming and reverse water gas shift reactions. It is important

to note that Wang et al. [44] as well as many other authors found that only a maximum of 6 moles

of hydrogen can actually be produced thermodynamically as opposed to the stoichiometric value

of 7.

1.3.5 Liquid Products

Although only gaseous products were taken into consideration when studying the possible

reactions in GSR, it should not be overlooked that liquid products can also be formed. In fact, there

are different possible reaction pathways for the GSR reaction. The wide distribution of possible

liquid products in the GSR reaction (Figure 1.15) is an indication of the complexity of the process.

Several reactions such as glycerol dehydrogenation, dehydration and hydrogenolysis take place

and a variety of chemical intermediates such as alcohols and ketones can be formed. It has been

suggested that since glycerol is not a thermally stable molecule, the extent of its pyrolysis can play

an important role under steam reforming conditions. It is likely that the intermediates formed from

glycerol cracking can also be reformed on the catalysts surface to produce hydrogen [45].

Figure 1.15: GSR liquid product distribution [45]

In this work, only acetaldehyde, acetone, hydroxyacetone and ethylene glycol will be quantified

and studied. As shown in Figure 1.15, acetaldehyde and hydroxyacetone are both produced via

24

glycerol dehydration. Hydrogenation of hydroxyacetone followed by the dehydration of the

resulting propylene glycol leads to the formation of acetone. Dehydrogenation and

decarboxylation of glycerol lead to the production of ethylene glycol (1,2 ethane diol).

1.3.6 Catalysts used in GSR

In the last few years, extensive research has been conducted on the use of heterogeneous catalysts

in GSR for the production of hydrogen and/or syngas. Several reviews showing GSR catalyst

developments have been published [3], [27], [46]. An ideal catalyst for GSR should be able to

actively cleave the C-C, O-H, and C-H bonds in the glycerol molecule whilst maintaining C-O

bonds to limit the formation of by-products. C-O bond cleavage is generally not desired as

subsequent hydrogenation leads to the formation of smaller alcohols and alkanes.

Noble metal based catalysts (Ru, Rh, Ir, Pd, Pt) are known to be active and stable during GSR.

However, their high cost and low availability limits their use and application on an industrial scale

[45]. This is why more focus is placed on developing catalytic systems based on transition metals

(Ni, Co, Cu). Hirai et al. [47] investigated the activity of catalysts made of group 8–10 metals

supported on La2O3 and the following order Ru ≥ Rh > Ni > Ir > Co > Pt > Pd > Fe was observed

where the Ru-based catalyst exhibited better catalytic properties. Compared to other noble metals,

ruthenium is relatively the least costly. Nevertheless, for economic reasons, only small weight

loadings are usually utilized (maximum of 2.5 wt%) [48]. Studies on ruthenium-based catalysts in

GSR are quite limited, even though ruthenium is known to reduce deactivation caused by coke

deposition in reforming reactions [49]–[51].

Amongst transition metals, Ni-based systems in particular have received considerable attention for

several reasons. Their lower cost and higher availability compared to noble metal catalysts makes

them attractive for industrial applications [3]. Nickel exhibits hydrogenation/dehydrogenation

capabilities and causes bond breakage in the following order: O-H, -CH2-, C-C and –CH3 [52]. It

also catalyzes the water gas shift reaction which removes adsorbed CO from the metallic surface

[53]. However, under typical steam reforming conditions, Ni-based catalysts suffer from

deactivation as a result of sintering of the metallic phase and/or coke deposition [45]. At higher

temperatures, Ni clusters tend to form, upon which unreactive carbon can easily form, leading to

the eventual deactivation of the catalyst due to the blocking of the pores and confinement of the

25

active sites [54]. Consequently, many efforts have been made to improve the performance of Ni-

based catalysts in GSR and reduce coking.

Several parameters of Ni-based catalysts can be modified such as the characteristics of the support,

metal content, introduction of promoters to the catalyst composition and the preparation method.

The use of additional metals in the support such as Mg and Al that form basic oxides can reduce

coking by improving the ability of Ni catalysts to chemisorb oxidants such as H2O and CO2 [55].

The introduction of promoters is also of key importance as it helps increase the catalytic activity.

Promoters are additionally capable of modifying the acid-base and redox properties of a catalyst

such as to reduce coke formation [54]. Lanthanides have been found to be good promoters as they

favor metal dispersion as well as strengthen CO2 adsorption on the support which prevents the

deposition of carbon through reverse disproportionation: [56],[57]

C + CO2 ↔ 2CO (18)

Lanthanum, in particular, has gained a lot of interest as several studies have reported its ability to

enhance catalytic stability [56], [58]. It has been suggested that the presence of oxycarbonates on

an La2O3 support can act as a dynamic oxygen pool thereby facilitating the removal of coke

[56],[57]. Rare earth oxides also influence nickel dispersion. Studies on lanthanum modified nickel

catalysts have shown an improved dispersion of Ni [54]. This high dispersion also delays metal

sintering thus improving coke resistance [57].

The preparation method affects the properties of catalytic materials as it influences the metal-

support interaction which in turn can define several properties including reducibility, resistance to

sintering and active metal dispersion [59]. Surface and structural properties also influence the

catalytic activity and coke-resistant behavior of the catalyst. Despite the prevalence of various

methods of active metal incorporation, impregnation techniques (wet, dry, incipient…) are the

most common and conventional ones where the active phase resides on the surface of the support

[60]. The technique of grafting or intercalation of an active phase has been used in recent years as

a novel approach in several works in order to stabilize the active phase in the support matrix [61]–

[63].

As mentioned previously, the choice of support influences the catalytic performance. The activity

of nickel supported on various pure metal oxides has been studied. Adhikari et al. [64] studied the

catalytic activity of Ni supported on MgO, CeO2 and TiO2. At 600oC and a WGFR of 9:1, a

26

hydrogen selectivity in the following order was observed: CeO2 (70%) > MgO (40%) > TiO2

(15%). Both Ni/CeO2 and Ni/MgO experienced a 100% total glycerol conversion at this

temperature. The CeO2 supported catalyst experienced the least coke formation whereas the TiO2

supported catalyst experienced the most. It has been suggested that the dual oxidation state (+4/+3)

of ceria leads to a release of oxygen which reacts with any deposited carbon thereby reducing coke

formation. Due to the acidic character of TiO2, more coke was formed.

Nichele et al. [65] compared the activities of Ni/ZrO2, Ni/SBA-15 and Ni/TiO2 at 650oC using a

WGFR of 46:1. The Ni/TiO2 catalyst exhibited negligible catalytic activity with a glycerol

conversion lower than 10%. Ni/ZrO2 gave a stable glycerol conversion of 72% even after 20 hours

while the Ni/SBA-15 catalyst gave an initial conversion of 49% but deactivated with time on

stream. The deactivation of Ni/SBA-15 was found to be due to its low hydrothermal stability in

steam conditions at high temperatures.

Papageridis et al. [45] studied the activity of Ni/Al2O3 at 600oC using a WGFR of 20:1. Initially,

the total glycerol conversion was 90.7% but it experienced a drastic deactivation in the first 7

hours. This behavior is typical for Ni/Al2O3 in GSR and has been reported by several authors [66],

[67]. Highly acidic supports in general tend to dehydrate glycerol and this yields undesired coke

precursors. Subsequent condensation of these coke precursors on the metal/support can clog the

system and result in deactivation. To deal with this issue, researchers have either used more basic

supports or introduced alkaline promoters to alumina to neutralize its acidic sites [68].

Dieuzeide et al. [53] showed that the addition of basic Mg(II) to Ni/Al2O3 helped minimize coke

formation. With increasing Mg (II) loading, there was a significant decrease in carbon formation.

However, a minimized coke formation did not necessarily mean an improved catalytic activity.

Iriondo et al. [69]–[71] studied the effect of the addition of different promoters (La2O3, CeO2 and

ZrO2) to the catalytic activity of Ni/Al2O3. The addition of these promoters led to an improvement

in catalytic activity as well as a decrease in coke formation. The addition of La2O3 in particular

reduced the acidic nature of alumina and improved the catalyst ability to transform oxygenated

hydrocarbons (OHCs) to carbon dioxide and hydrogen.

Table 1.3 presents a summary of the conditions and performance of the previously described Ni-

based pure metal oxides used in GSR. Some works used H2 yield (%) as a reaction metric for the

27

production of hydrogen whereas others used H2 selectivity (%). Based on the results of these

studies, the use of acidic supports is clearly not preferable in the GSR process.

Table 1.3: Pure metal oxide catalysts used in GSR

Catalyst Mass of

catalyst

Operating

Conditions

Total Glycerol

Conversion

H2 yield/H2

selectivity

Reference

11.6 wt%

Ni/CeO2

750 mg P= 1 atm

T= 600oC

WGFR = 9:1

Feed flow rate =

0.5 mL/min

100% H2 selectivity:

70%

[64]

9.62 wt%

Ni/MgO

750 mg P= 1 atm

T= 600oC

WGFR = 9:1

Feed flow rate =

0.5 mL/min

100% H2 selectivity:

40%

[64]

12.7 wt%

Ni/TiO2

750 mg P= 1 atm

T= 600oC

WGFR = 9:1

Feed flow rate =

0.5 mL/min

63% H2 selectivity:

15%

[64]

10 wt%

Ni/TiO2

200 mg

P=1 atm

T= 650oC

WGFR = 46:1

Feed flow rate =

0.06 mL/min

<10% H2 yield:

< 5%

[65]

10 wt%

Ni/ZrO2

200 mg

P=1 atm

T= 650oC

WGFR = 46:1

Feed flow rate =

0.06 mL/min

72% H2 yield:

65%

[65]

10 wt%

Ni/SBA-15

200 mg P=1 atm

T= 650oC

WGFR = 46:1

Feed flow rate =

0.06 mL/min

49% H2 yield:

37%

[65]

8 wt%

Ni/Al2O3

200 mg P=1 atm

T= 600oC

WGFR = 20:1

Feed flow rate =

0.12 mL/min

90.7% H2 selectivity:

61%

[45]

3 wt% Mg(II)

modified

Ni/Al2O3

45.5mg P=1 atm

T= 600oC

WGFR = 9:1

Contact time =

3.09 gcat

h/molglycerol

90.3% H2 yield:

77%

[53]

28

1.4 Choice of catalysts

Based on the literature study, catalysts with basic properties appear to be promising candidates for

both the biodiesel and steam reforming of glycerol reactions. Mixed oxides derived from

hydrotalcites have recently gained interest in the development of new environmentally friendly

catalysts. Mg–Al hydrotalcites and their calcination products in particular are very relevant as

basic catalysts [72]. The main focus of this work will therefore be the use of Mg-Al hydrotalcite

derived mixed oxides as support materials for both the biodiesel and the steam reforming of

glycerol reactions. The active metal phases utilized for each reaction will be different. For the

biodiesel reaction, calcium will be used as the active phase since CaO is cheap and known for its

high basicity. Whereas for the steam reforming of glycerol, ruthenium and nickel will be used as

the active phases given their known activities. For the Ni-based hydrotalcite catalysts, the supports

will also be modified with lanthanum to study the effect of promoter addition. Ni on some pure

metal oxide supports (CeO2, Y2O3 and ZrO2) will also be studied and compared to the Ni on

hydrotalcite derived mixed oxide supports in the steam reforming reaction. Different preparation

methods (impregnation vs grafting) will also be compared to study their effect on the catalytic

properties and activity.

8.4 wt% Ni/γ-

Al2O3/6 wt%

La2O3

200 mg

T= 500oC

P= 4 atm

WHSV= 7.7h-1

WGFR = 46:1

100% H2 yield:

99.7%

[69]

11.9 wt% Ni/

γ- Al2O3/4.3

wt% CeO2

200mg T= 500oC

P= 4 atm

WHSV= 7.7h-1

WGFR = 46:1

100% H2 yield:

80.3%

[70]

13.4 wt% Ni/

γ- Al2O3-7

wt% ZrO2

200mg T= 500oC

P= 4 atm

WHSV= 7.7h-1

WGFR = 46:1

100% H2 yield:

62%

[71]

29

1.5 Hydrotalcites or Anionic Clays

1.5.1 Structure

Hydrotalcites or hydrotalcite-like compounds are double lamellar hydroxides having the general

formula: [M2+1-xM

3+x(OH)2]

x+(An-) x/n.mH2O where M2+ is a divalent metal ion, M3+ is a trivalent

metal ion, and An- is an anion.

In order to understand the structure, one must look at the unit crystallographic structure from which

all hydrotalcites are derived. In the case of the mineral “hydrotalcite” (Mg6Al2(OH)16CO3.4H2O),

this structure is called brucite and has a formula of [Mg(OH)2]. In brucite, the Mg2+ ions are 6-fold

coordinated to OH- ions to form octahedra. These octahedra share edges to form infinite sheets

that are stacked on top of each other and held together by hydrogen bonds. When these Mg2+ ions

are substituted by a trivalent cation having a similar radius such as Al3+, an excess of positive

charge is created in the hydroxyl sheet which is compensated by the CO32- anions that lie in the

interlayer space between the hydroxyl sheets. This interlayer space also contains the water of

crystallization.

Figure 1.16: Schematic representation of the structure of hydrotalcites

For other hydrotalcites with different compositions, the sheets are built just as in brucite but

containing other cations. These cations are also 6-fold coordinated to OH- ions and randomly

occupy the octahedral holes in between the packed OH- ions. Anions and water are also located

randomly in the interlayer between the sheets and are free to move by breaking their bonds and

forming new ones [73].

30

1.5.2 Properties

Hydrotalcites are used mostly after calcination or thermal decomposition which leads to the

formation of metal oxides. These metal oxides exhibit several interesting properties making them

suitable for various catalytic applications:

High surface area: Thermal decomposition leads to the escape of CO2 from the interlayer

and the removal of water molecules from the structure. This results in the creation of more

pores and craters on the surface of the material [74].

Basic properties: Mg–Al hydrotalcites and their calcination products are particularly well-

known for their basic properties. One of the most unique properties of hydrotalcites is the

possibility of tuning their basic strength. This can be achieved by changing the nature of

the metal cations (M2+, M3+), varying the flexible M2+/M3+ ratio of the structure,

introducing a suitable anion in the interlayer, controlled thermal treatment or doping with

various elements or promoters. This makes hydrotalcites attractive for use as catalyst

supports [72].

Memory effect: the hydrotalcite structure can reconstruct itself and retain its original

structure when treated with aqueous solutions containing anions [73].

High anionic exchange capacity: hydrotalcites allow the exchange of their original anions

with those present in an aqueous solution. Changing the nature of the anion (CO32-, NO3

-,

Cl-, SO42-) can help tune the properties of the material [72].

Formation of homogeneous mixtures of oxides with high specific surface area, stable to

thermal treatments, which upon reduction form small and thermally stable crystallites with

the absence of chemical segregation [75].

Finally, hydrotalcites are stable against sintering, provide high metal dispersion and are

inexpensive [73].

1.5.3 Hydrotalcites used in biodiesel production

As mentioned previously, research efforts are currently focused on the search for an efficient,

heterogeneous, basic catalyst for the transesterification reaction. Table 1.4 presents the results of

some selected hydrotalcite-based catalysts used for biodiesel production under different

conditions. The use of Mg-Al hydrotalcites for biodiesel production appears promising as many of

these studies obtained very high yields.

31

Table 1.4: Hydrotalcite catalysts used in biodiesel production

Hydrotalcite Methanol

/Oil ratio

Catalyst

Wt %

Mg/

Al

Temp

Rxn

˚C

Rxn

Time

hrs

Calcination

Temp

˚C

Calcination

Time

hrs

Oil

used

Yield

%

Reference

Mg-Al 6:1 1.5 3 65 4 500 12 Rape oil 90.5 [76]

HT2B(Mg-

Al)

9:1 or

12:1

2.5 2 60-65 4 700 5 - 97.1 [77]

Mg-Al 15:1 7.5 3 - 9 500 8 Soybean

oil

67 [15]

Mg-Al 6:1

+methanol

/ethanol

ratio =4.2

1.5 3 80 2.5 500 - Pretreated

waste

cooking

oil

95.2 [78]

KF/Ca-Al

(80%)

12:1 5 - 65 5 550 5 Palm oil 97.98 [79]

KF/Ca-Al

(100%)

12:1 5 - 65 3 550 5 Palm oil 99.74 [79]

Ca-Al 12:1 5 - 65 5 Non-

calcined

- Palm oil 4.35 [79]

Ca-Al 12:1 5 - 65 5 Calcined 5 Palm oil 8.24 [79]

KF/Mg-Al

(80%)

12:1 5 - 65 5 Calcined 5 Palm oil 88.7 [79]

KF/HT(Ca-

Mg-Al)

KF load

92.1%

12:1 5 - 65 3 600 3 Palm oil 95.3 [80]

1.5.4 Hydrotalcites used in GSR

Some works have also studied the use of hydrotalcite-like compounds in the steam reforming of

glycerol. Their results are summarized in Table 1.5. High hydrogen selectivities as well as total

glycerol conversions were obtained over some of these catalysts.

Table 1.5: Hydrotalcites used in GSR

Hydrotalcite Mass of catalyst Operating

Conditions

Total Glycerol

Conversion

H2 yield/H2

selectivity

Reference

Ni-Mg-Al 1000 mg P= 1 atm

T= 650oC

WGFR = 9:1

88% H2 selectivity:

78.5%

[44]

Ni-Cu-Al 1000 mg P= 1 atm

T= 650oC

WGFR = 9:1

96% H2 selectivity:

92.9%

[81]

Pt/Mg-Al 40 mg P= 1 atm

T= 600oC

WGFR = 9:1

85% (4hrs) H2 selectivity:

48.5%

[82]

32

1.6 Catalyst Deactivation

An ideal catalyst should not only exhibit a high activity and high product selectivity; it should also

be stable over a long period of time [83]. Catalyst deactivation is a phenomenon characterized by

the loss of catalytic activity as well as product selectivity over the reaction time. All catalysts

experience a decline in performance at some point and must eventually be replaced. However, the

time scale for this deactivation can vary from seconds to years depending on the catalyst. To deal

with catalyst deactivation, industries often have to undergo catalyst replacements, regeneration

and plant shut-downs, all of which involve increased energy consumption and a loss of production

time. The additional costs required for these processes can sometimes amount to billions of dollars

annually for an industry [84]. Therefore, it is necessary to understand the mechanisms behind

catalyst deactivation as more preventive measures can be taken to prolong the time before a

catalyst begins to undergo deactivation. The mechanisms behind catalyst deactivation can be

categorized as either chemical and/or physical. Three main catalyst deactivation mechanisms will

be discussed:

Poisoning

Poisoning is the loss of catalytic activity as a result of the chemisorption of impurities on active

sites of the catalyst. These impurities are often found in small quantities in the feed stream. Most

poisons act by blocking active reaction sites or modifying their ability to adsorb and/or dissociate

reactant molecules. For steam reforming catalysts, in general, sulfur is the most severe poison [85].

Under reforming conditions, sulfur is present in the form of H2S which is chemisorbed on active

metal sites:

H2S + M M-S + H2 (19)

A sulfur-poisoned catalyst can be regenerated by removing the retained sulfur by oxidation or

controlled re-reduction.

Sintering

Sintering is an irreversible physical process and a thermally-induced form of catalyst deactivation.

It involves the loss of active catalyst surface as a result of changes in the crystalline structure.

Operation at high temperatures for extended periods of time often results in the growth of crystals

33

either from the support material or the active phase. Sintering agents like steam can also accelerate

the sintering process.

Fouling

Fouling is associated with the physical/mechanical deposition of species originating from the

reacting medium onto the catalyst surface and pores. The loss of activity results from active site

coverage and pore blockage. As opposed to poisoning which is a chemical phenomenon, fouling

is entirely physical. Carbon and coke deposition/formation are the most prominent types of

catalytic fouling. In steam reforming processes, carbon and coke deposition/formation are the main

causes of catalyst deactivation as these deposits block the access of the hydrocarbon and steam

reactants.

It is important to make a distinction between the terms “carbon” and “coke”. Although both terms

are sometimes used interchangeably, by default their definition is related to their origin. Carbon is

considered the product of CO disproportionation whereas the term coke refers to the material that

originates from the decomposition or condensation of hydrocarbons. Coke usually consists of

polymerized heavy hydrocarbons. However, the composition of coke could vary from high

molecular weight hydrocarbons to primary carbons such as graphite depending on reaction

conditions.

The mechanisms of carbon deposition and coke formation on active metal sites from CO and

hydrocarbons are shown in Figure 1.17 and 1.18 respectively.

34

Figure 1.17 shows the formation of carbon from CO on Ni-based catalysts.

Figure 1.17: Formation, transformation and gasification of carbon on Ni [86]

CO dissociates on metal sites to form an adsorbed atomic type of carbon (Cα) and atomic oxygen.

This Cα could further react to form a polymeric type of carbon (Cβ). Both types of carbon are

formed at lower temperatures, amorphous in structure and highly reactive. At higher temperatures

and over extended periods of time, they can be converted to a less reactive, graphitic form of

carbon (Cc). Atomic carbon (Cα) can also form a metal carbide (Cγ) or yield whisker-like,

vermicular carbon (Cν). Table 1.6 summarizes the properties of each type of carbon including the

temperature of formation and gasification temperature with H2.

Table 1.6: Types of carbon species formed by decomposition of CO on Ni [86]

Structure Designation Temperature of

formation (oC)

Temperature of

gasification with

H2 (oC)

Adsorbed, atomic (surface carbide) Cα 200-400 200

Polymeric, amorphous films or filaments Cβ 250-500 400

Vermicular filaments, fibers and/or

whiskers

Cν 300-1000 400-600

Nickel carbide (bulk) Cγ 150-250 275

Graphitic (crystalline) platelets or film Cc 500-550 550-850

Figure 1.18 shows the formation of coke from hydrocarbons. It is obviously more complicated

than CO dissociation as various segmented hydrocarbons species are produced (CHx, C2Hy, CnHz

etc.) which may condense to form higher molecular weight coke.

35

Figure 1.18: Formation and transformation of coke on metal surfaces [86]

Table 1.7 provides a summary of the properties of the different coke/carbon species formed in

steam reforming of hydrocarbons on nickel catalysts.

Table 1.7: Carbon species formed in the steam reforming of hydrocarbons over Ni [86]

Encapsulating Film Whisker-like Pyrolytic carbon

Formation Slow polymerization of

CnHm radicals on Ni-

surface, into

encapsulating film

Diffusion of C through

Ni-crystal, nucleation

and whisker growth

with Ni-crystal at top

Thermal cracking of

hydrocarbon;

deposition of C

precursors on catalyst

Effects Progressive

deactivation

No deactivation on Ni-

surface: breakdown of

catalyst and increasing

ΔP

Encapsulation of

catalyst particle;

deactivation and

increasing ΔP

Temperature range

(oC)

<500 >450 >600

Critical Parameters Low temperature

Low H2O/CnHm

Low H2/CnHm

Aromatic feed

High temperature

Low H2O/CnHm

No enhanced H2O

adsorption

Low activity

Aromatic feed

High temperature

High void fraction

Low H2O/CnHm

High pressure

Acidic catalyst

As shown in Table 1.7, some types of coke result in a loss of catalytic activity and some do not.

As discussed previously, to minimize carbon and coke deposition, there are two possible strategies.

36

The first strategy involves the addition of a dopant (alkaline earth metal or lanthanide groups) into

the catalyst composition. The second strategy involves the control of reaction conditions

(temperature, feed ratios, etc.) such as to reduce carbon and coke deposition whilst still maintaining

a reasonable catalytic activity.

1.7 Conclusion

The production of biodiesel via the transesterification of vegetable oils results in the formation of

glycerol as a by-product. This obtained glycerol can be valorized via steam reforming to produce

hydrogen. Literature studies demonstrate that both of these reactions require basic catalysts to

function efficiently. Thus, a common support material was selected to accommodate both

reactions. In this work, Mg-Al hydrotalcites will be investigated as the support materials as they

are well-known for their basic character. Different active phases suitable to each reaction will be

utilized.

In Chapter 2, Ca-based Mg-Al catalysts will be tested for the production of biodiesel. In Chapter

3, Ni-based pure metal oxides will first be studied in the glycerol steam reforming reaction. In

Chapter 4, both Ru-based Mg-Al catalysts and Ni-based Mg-Al catalysts will be tested for glycerol

steam reforming reaction. In the case of the Ni-based Mg-Al catalysts, the Mg-Al supports will

also be modified with the promoter (La) to study the effect of promoter addition. The activity of

the most efficient catalyst from this series will be compared to the activity of the most efficient

Ni-based pure metal oxide from Chapter 3.

37

Chapter 2: Biodiesel production over Ca-

based Mg-Al hydrotalcite derived catalysts

38

This chapter focuses on the use of Mg-Al type hydrotalcite catalysts in the transesterification

reaction of sunflower oil for the production of biodiesel. Calcium was used as the active phase.

The feedstock utilized (sunflower vegetable oil) was first characterized to ensure its suitability.

Following synthesis and characterization, all prepared catalysts were tested in the

transesterification reaction. The effects of thermal treatment and the preparation method for Ca

incorporation (grafting vs impregnation) were compared. Optimization of reaction parameters was

then conducted on the most efficient catalyst. Finally, the obtained biodiesel was also characterized

and its properties were compared to those of conventional diesel.

2.1 Characterization of vegetable oil

It is important to evaluate the properties of the feedstock prior to its use in the transesterification

process. Therefore, the physical and chemical properties (density, kinematic viscosity, FFA

content, saponification value and water content) of sunflower vegetable oil were determined

according to standard methods (detailed in Appendix A). The results listed in Table 2.1 are in good

accordance with the specified ASTM (American Society for Testing and Materials) and AOCS

(American Oil Chemists’ Society) requirements. Therefore, sunflower vegetable oil can be used

as a feedstock for the transesterification process.

Table 2.1: Physical and chemical properties of sunflower oil

2.2 Synthesis of catalysts

The preferred method to use when synthesizing hydrotalcites is the co-precipitation technique at a

constant pH in order to obtain a homogenous phase when combining several metals. Generally

speaking, this coprecipitation is conducted at low supersaturation where the pH is controlled by

Property Value ASTM Requirement Test Method

Density at 25 ºC (Kg/m3) 920.50 918.00 – 921.00 ASTM D4052

Kinematic Viscosity at 40

ºC (mm2/s)

53.87 - ASTM D445

FFA Content

(wt%)

0.22 0.5 max AOCS Ca-5a-40

Saponification Value (mg

KOH/g)

194 188 – 194 AOCS Cd-3-25

Water Content

(wt%)

0.09 0.2 max AOCS Ca-2c-25

39

the slow addition of two solutions, one containing the metals and the second containing the base

[73].

Appropriate amounts of Mg(NO3)2.6H2O (Sigma-Aldrich, 98% purity) and Al(NO3)3.9H2O

(Sigma-Aldrich, 98% purity) were dissolved in distilled, deionized water to prepare an aqueous

solution (250 mL). This aqueous solution was precipitated using a basic solution of 1 M sodium

carbonate Na2CO3 (Uni-Chem, 99.8% purity) and 2M sodium hydroxide NaOH (SureChem

Products, 96% purity) at a temperature of 60˚C whilst maintaining the pH between 9.5 and 10

during stirring. The metal nitrate solution was added drop by drop, the pH was adjusted to 9.5-10

using the basic solution, and the formation of a white precipitate was observed. The resulting

suspension was stirred for 1 hour at 60˚C. It was then placed for 18 hours in an oven at 60˚C to

allow crystallization of the hydrotalcite phase. The mixture was then filtered and continuously

washed with hot distilled water (60˚C) until a neutral pH of the filtrate was obtained. This washing

step is necessary to remove all soluble ions such as nitrates and sodium ions. The solid was

allowed to dry for 48 hours in an oven at 60˚C and then grinded to a powder form. The prepared

support was named Mg4Al2 where the subscripted numbers represent the respective molar ratios

of the metals.

The prepared support Mg4Al2, in both uncalcined and calcined form, was impregnated with 40

wt% Ca via a wet impregnation technique. An appropriate quantity of Ca(NO3)2.4H2O (Sigma-

Aldrich, 98% purity) was dissolved in 50 ml of distilled water and added to 1 gram of the support

and stirred for 2 hours at room temperature. The excess of water was slowly eliminated in a rotary

evaporator (60°C, under vacuum). The obtained residue is kept in an oven at 60°C during 24 h.

The catalysts were then calcined at 600 ˚C for 4 hours at a rate of 1 ˚C/min.

For comparison, two Ca containing hydrotalcites were additionally synthesized by co-precipitation

where appropriate amounts of Ca(NO3)2.4H2O, Mg(NO3)2.6H2O and Al(NO3)3.9H2O were

precipitated together with a basic solution following the same protocol as above. This technique is

known as grafting as it involves the stabilization of the active phase Ca in the catalyst matrix [87].

The prepared solids were directly calcined at 600˚C under a flow of air for 4 hours (rate: 1˚C/min)

and named Ca2Mg2Al2 and Ca4Al2.

The calcination temperature greatly influences the total basicity as well as the surface area of

catalysts and 600oC was found to be the optimum calcination temperature when using CaO as an

active phase [16],[80],[88]. A summary of the prepared samples and their characteristics is

40

presented in Table 2.2. Regarding catalyst naming, the abbreviation HT represents catalysts in

their uncalcined form, whereas the numbers indicate the corresponding calcination temperatures.

Table 2.2: Summary of the prepared catalysts

Catalyst

prepared

Calcination

Temperature

(˚C)

Impregnation Calcination

after

impregnation

Sample name

Mg4Al2 - - - Mg4Al2 HT

Mg4Al2 600oC - - Mg4Al2 600

Mg4Al2 - 40 wt% Ca 600˚C Ca600/Mg4Al2HT

Mg4Al2 600oC 40 wt% Ca 600˚C Ca600/Mg4Al2 600

Ca2Mg2Al2 - - - Ca2Mg2Al2HT

Ca2Mg2Al2 600˚C - - Ca2Mg2Al2600

Ca4Al2 - - - Ca4Al2 HT

Ca4Al2 600˚C - - Ca4Al2 600

2.3 Characterization of catalysts

The prepared catalysts were characterized using various techniques. The characterization

techniques used, their description and conditions are presented in Appendix A.

2.3.1 X-ray diffractograms (XRD) of uncalcined and calcined Ca-Mg-Al catalysts

The XRD diffractograms of the uncalcined solids are shown in Figure 2.1.

10 20 30 40 50 60 70 80 90

** **

**

Ca4Al

2 HT

Ca2Mg

2Al

2 HT

Mg4Al

2 HT

Inte

nsity (

a.u

.)

2()

***

*

* ***

* * * * * **

CaCO3

Ca(OH)2

* Hydrotalcite phase

Figure 2.1: XRD diffractograms of uncalcined hydrotalcite supports

41

The diffractogram of Mg4Al2HT shows typical diffraction peaks for the hydrotalcite structure

(JCPDS 22-0700) [19]. Less intense hydrotalcite peaks are observed for Ca2Mg2Al2HT indicating

a lower degree of crystallinity as a result of Ca2+ incorporation into the brucite layers [18]. This

could be explained by the larger ionic radius of Ca2+ (0.1nm) compared to Mg2+ (0.072nm). No

hydrotalcite peaks are observed for the Ca4Al2HT. Peaks corresponding to the CaCO3 phase

(JCPDS 33-0268) are observed in both calcium containing solids indicating that Ca2+ and CO32-

combined directly. This results from the huge difference of ionic radii sizes between Ca2+ (0.1nm)

and Al3+ (0.053nm) as well as the high solubility product constant (Ksp) of CaCO3 [89]. The

Ca(OH)2 (JCPDS 84-1263) phase is also observed in both Ca2Mg2Al2HT and Ca4Al2HT catalysts.

Figure 2.2 shows X-ray diffractograms of the calcined catalysts.

10 20 30 40 50 60 70 80 90

# MgO

MgAl2O

4

CaO

Ca(OH)2

CaCO3

##

###

Ca4Al

2 600

Ca2Mg

2Al

2 600

Ca 600/Mg4Al

2 600

Ca 600/Mg4Al

2HT

Inte

nsity (

a.u

.)

2()

Mg4Al

2 600

### #

#

##

## ##

Figure 2.2: XRD diffractograms of calcined supports (Mg4Al2600, Ca2Mg2Al2600, Ca4Al2600) and Ca impregnated Mg4Al2

support (Ca600/Mg4Al2HT, Ca600/Mg4Al2600)

Following thermal treatment, the Mg-Al hydrotalcite is transformed to an Mg(Al)O mixed oxide.

Peaks corresponding to the periclase MgO (JCPDS 45-0946) are observed for all magnesium

containing catalysts except Ca600/Mg4Al2600. For all catalysts, the absence of peaks

corresponding to AlOx indicates that Al3+ ions are well dispersed in the MgO lattice. Ca600/

Mg4Al2HT exhibits peaks corresponding to the spinel phase MgAl2O4 (JCPDS 075-1798) and CaO

phase (JCPDS 37-1497). No mixed oxide composed of Ca is observed. This could also be due to

the difference in ionic radii between Ca2+ and Mg2+ [88]. Ca(OH)2 (JCPDS 84-1263) and CaCO3

(JCPDS 33-0268) were still observed for the Ca600/Mg4Al2600, Ca2Mg2Al2600 and Ca4Al2600

42

catalysts. This is due to the hydration and carbonation of basic CaO in contact with air [79]. The

double thermal treatment for Ca600/Mg4Al2600 resulted in less intense peaks compared to

Ca2Mg2Al2600 and Ca4Al2600.

2.3.2 N2 sorption isotherms of calcined Ca-Mg-Al catalysts

N2 adsorption-desorption isotherms for the calcined catalysts are displayed in Figure 2.3.

0.0 0.2 0.4 0.6 0.8 1.0

Ad

so

rbe

d N

2 (

cm

3.g

-1)

Relative Pressure (P/P0)

Mg4Al

2 600

Ca 600/Mg4Al

2HT

Ca 600/Mg4Al

2 600

Ca2Mg

2Al

2 600

Ca4Al

2 600

Figure 2.3: N2 adsorption/desorption isotherms of calcined catalysts

The obtained results reveal that all the prepared catalysts are mesoporous. The isotherms are

classified as type IV with a clear hysteresis loop. This is characteristic of hydrotalcite derived

materials [90]. Following thermal treatment, surface areas and porosities of hydrotalcite derived

materials increase due to the decomposition of the structure and creation of craters throughout the

layers following the departure of H2O and CO2 [20]. It is worth noting that both impregnated

samples (Ca600/Mg4Al2HT and Ca600/Mg4Al2600) exhibited smaller loops as a result of pore

clogging due to Ca incorporation. These hysteresis loops were also shifted towards a higher

relative pressure indicating an increase in the average pore size. For the grafted catalysts

(Ca2Mg2Al2600 and Ca4Al2600), as Ca content increased in the hydrotalcite composition, the

hysteresis loop became more compact. Between all catalysts, the isotherm of Ca600/Mg4Al2HT

was the most parallel to the x-axis indicating that the pores were almost entirely filled with CaO

[25].

43

The good textural properties of these samples are deduced from the isotherms and are shown in

Figure 2.4 and Table 2.3. Average pore diameters of all catalysts are shown in Figure 2.4 and vary

from 10-20nm.

0 20 40 60 800.0

0.5

1.0

1.5

2.0

2.5

3.0

3.5

dV

/dD

(cm

3.g

-1.n

m-1)

Pore diameter (nm)

Mg4Al

2 600

Ca 600/Mg4Al

2HT

Ca 600/Mg4Al

2 600

Ca2Mg

2Al

2 600

Ca4Al

2 600

Figure 2.4: Pore size distribution of calcined catalysts

As speculated from the isotherms of Figure 2.3, Ca600/Mg4Al2HT and Ca600/Mg4Al2600 both

exhibit an increased average pore diameter compared to the other catalysts.

Table 2.3 displays specific surface areas obtained for all calcined catalysts.

Table 2.3: Specific surface areas of calcined catalysts

Following calcination, Mg-Al mixed oxides possess high surface areas. This is evident for

Mg4Al2600 (213 m2/g). However, after impregnation with 40 wt% Ca, a drastic decrease is

observed as a result of pore clogging [88]. The Ca600/Mg4Al2HT possessed the lowest surface

area between all catalysts (12 m2/g). This verifies that the pores were entirely filled with CaO as

Catalyst Surface area

(m2/g)

Mg4Al2600 213

Ca600/Mg4Al2HT 12

Ca600/Mg4Al2600 28

Ca2Mg2Al2600 109

Ca4Al2600 41

44

speculated from the isotherms (Figure 2.3). As Ca content increased in the hydrotalcite

composition of the grafted catalysts, surface areas decreased.

2.3.3 CO2-Temperature Programmed Desorption (CO2-TPD) of calcined Ca-Mg-Al

catalysts

Basic properties of catalysts were studied by CO2-temperature programmed desorption and

obtained profiles are shown in Figure 2.5.

100 200 300 400 500 600 700

Inte

nsity (

a.u

.)

Temperature (oC)

Ca4Al

2 600

Ca2Mg

2Al

2 600

Ca 600/Mg4Al

2 600

Ca 600/Mg4Al

2HT

Mg4Al

2 600

Figure 2.5: CO2-TPD profiles of calcined catalysts

The amount of CO2 adsorbed corresponds to the amount of basic sites and the temperature

indicates the strength of the basic site. The stronger the basic site, the higher the CO2 desorption

temperature. Desorption peaks in the temperature range of 50-200oC corresponds to weak basic

sites resulting from the interaction of CO2 with surface hydroxyl groups [22]. In this temperature

range, the most intense peak was observed for the Mg4Al2600 catalyst. Impregnation with calcium

led to a decrease in the intensity of this band since CaO covered the Mg-Al surface and reduced

CO2 adsorption on the surface. Medium strength basic sites between 450-600oC are due to the

interaction of CO2 with metal-oxygen pairs more specifically Ca2+-O2- pairs. It is evident that the

addition of Ca to the catalyst composition increases the presence of strong basic sites (T>600oC)

[88]. This is particularly evident for Ca600/Mg4Al2HT and Ca4Al2600. The Ca600/Mg4Al2HT

appears to be the only catalyst with basic sites in the 300-450oC range.

45

2.4 Biodiesel test results

The biodiesel test set-up and experimental procedure are detailed in Appendix A. The catalytic

activity of KOH as a homogeneous catalyst was first tested (Section 2.4.1) followed by the activity

of the prepared heterogeneous catalysts (Section 2.4.2).

2.4.1 Catalytic activity of homogeneous catalyst

KOH was first tested for the transesterification of sunflower oil. KOH was used as a representative

homogeneous catalyst since it is able to produce a very high Fatty Acid Methyl Ester (FAME)

yield as the reaction goes essentially to completion. The conditions used were a stirring rate of 400

rpm, 6:1 methanol to oil molar ratio, 1 wt% catalyst to oil ratio, reaction time of 1 hour and a

temperature of 60 ˚C [91]. A FAME yield of 98.3% was obtained. Calculations of the FAME

content for the homogeneous KOH and determination of the fatty acid methyl ester profile are

detailed in Appendix A. The fatty acid methyl ester profile of the biodiesel produced via a

homogeneous catalyst is shown in Table 2.4. Similar results for the relative percentages have been

reported by other authors that also used sunflower vegetable oil as a feedstock [24], [92].

Table 2.4: FAME profile produced from homogeneous catalysis

Corresponding Methyl Ester Relative Percentage (%)

Methyl Dodecanoate (C12:0) 0.06

Methyl Myristate (C14:0) 0.09

Methyl Palmitoleate (C16:0) 0.12

Methyl Palmitate (C16:1) 6.15

Methyl Heptadecanoate (IS) (C17:0) -

Methyl Linoleate (C18:2) 58.70

Methyl Oleate (C18:1) 30.50

Methyl Stearate (C18:0) 3.40

Methyl Arachidate (C20:0) 0.25

Methyl Lignocerate (C24:0) 0.72

The names of each methyl ester are followed by their lipid numbers in brackets. Lipid numbers

take the form of Cx:y where x is the number of carbon atoms and y is the number of double bonds

in the fatty acid chain.

46

The methyl ester peaks produced were identified by comparing the peaks in the GC chromatogram

obtained from homogeneous catalysis (KOH) to the peaks obtained from a standard FAME mix

C4-C24 (Appendix A). After labelling the peaks corresponding to each methyl ester, their relative

percentages were calculated based on the peak areas. As shown in Table 2.4, nine main fatty acid

methyl esters were detected in the obtained chromatogram excluding the internal standard (IS).

The calculation of the FAME content of the heterogeneous catalysts was based on the peak area

of each methyl ester peak identified in this section using Equation A.6 (Appendix A).

2.4.2 Catalytic activity of hydrotalcite derived mixed oxides

All prepared catalysts were tested for the transesterification of sunflower oil. Experiments were

performed at a stirring rate of 400 rpm, 12:1 methanol to oil molar ratio, 2.5 wt% catalysts to oil

ratio and a temperature of 60 ˚C.

2.4.2.1 Influence of support

The hydrotalcite supports with different compositions (Mg4Al2, Ca2Mg2Al2 and Ca4Al2) and in

their uncalcined (HT) and calcined form were tested in the transesterification reaction for biodiesel

synthesis to study the effect of composition and calcination on FAME yield. The obtained results

are shown in Figure 2.6.

At the same reaction conditions (Stirring rate of 400 rpm, 12:1 methanol to oil molar ratio, 2.5

wt% catalyst to oil and temperature of 60 ˚C for 4 hours), the Mg4Al2 support was tested both in

its uncalcined (Mg4Al2HT) and calcined form (Mg4Al2600). The obtained results showed that the

two catalysts have a low activity (FAME yield less than 2 %). In an attempt to increase the basicity

of the support, Ca was integrated into the hydrotalcite structure with different compositions. Two

catalysts were prepared, Ca2Mg2Al2 and Ca4Al2. The transesterification reactions performed using

the uncalcined and calcined Ca2Mg2Al2 showed no significant enhanced activity. The yield

obtained for uncalcined Ca2Mg2Al2 was 3.75 % and that of the calcined was 4.12 %. Even though

the calcium fraction was even higher in the Ca4Al2 solid, the FAME % remained insignificant and

was less than 1% for both the uncalcined and calcined form. It has been reported that hydrotalcites

in the absence of calcinations provide no catalytic activity for the transesterification reaction [93].

The lack of activity for the calcined Ca-containing catalysts could be attributed to the

inaccessibility of the calcium active phase as a result of the preparation method. CaO is known for

its good catalytic activity in the transesterification reaction, however it appears that in these

47

catalysts CaO was trapped in the bulk of the material [61]. Another possible explanation for the

lack of activity could be the formation of the CaCO3 and Ca(OH)2 phases (shown in XRD

diffractograms, Figure 2.2) resulting from contact between basic CaO and air (moisture and CO2)

[79],[13].

1.22

3.75

0.57 0.24

4.12

0.48

0

5

10

15

20

Ca 4Al 2

Ca 2Mg 2

Al 2

Calcined (600)F

AM

E y

ield

(%

)Uncalcined (HT)

Mg 4Al 2

Mg 4Al 2

Ca 2Mg 2

Al 2

Ca 4Al 2

Figure 2.6: FAME yield obtained over uncalcined and calcined Mg4Al2, Ca2Mg2Al2 and Ca4Al2 catalysts (Stirring rate of

400 rpm, 12:1 MOMR, 2.5 wt% CTOR and temperature of 60 ˚C for 4 hours)

2.4.2.2 Influence of impregnation

The hydrotalcite supports prepared were inactive for transesterification under the reaction

conditions used in this work. Therefore, both uncalcined and calcined Mg4Al2 supports were

impregnated with 40 wt% Ca, calcined at 600˚C and then evaluated in the biodiesel synthesis

reaction. No significant effects were detected for the Ca600/Mg4Al2600 catalyst. This may be due

to the double thermal treatment (600oC and then 600oC). Calcination results in the loss of interlayer

and crystallization water as well as CO2 escape and this increases the availability of basic sites as

well as catalytic activity [77]. However, some thermal treatments can damage the crystal structure

and cause the formation of phases which could clog the pore structure and cause its collapse [78].

This leads to a decrease in biodiesel yield. This could be the case for the Ca600/Mg4Al2600

catalyst. CaCO3 and Ca(OH)2 phases, known for their low activity in the biodiesel reaction [94],

were also observed for this catalyst.

48

On the other hand, calcium impregnation greatly improved the catalytic activity as it was observed

for Ca600/Mg4Al2HT catalyst by increasing the FAME yield to 70% as shown in Figure 2.7. The

Ca impregnated on the surface clearly created active sites for the reaction to occur. CaO is

responsible for the high activity of the catalyst due to the formation of O2- anions that serve as

active centers [91]. XRD diffractograms (Figure 2.2) verified the presence of the CaO phase and

spinel MgAl2O4. The addition of Ca enhanced the basic properties and increased basic site strength

as verified in CO2-TPD profiles (Figure 2.5).

1.22 0.24

70

0.910

20

40

60

80

100 Mg

4Al

2

Ca

60

0/M

g4A

l 2 6

00

Ca

lcin

ed

FA

ME

yie

ld (

%)

Un

ca

lcin

ed

Ca

60

0/M

g4A

l 2H

T

Ca impregnated Mg4Al

2

Figure 2.7: FAME yield obtained over Ca impregnated on Mg4Al2 support (Stirring rate of 400 rpm, 12:1 MOMR ratio,

2.5 wt% CTOR and temperature of 60 ˚C for 4 hours)

49

2.5 Optimization

To determine the optimum reaction conditions, the effect of catalyst to oil ratio (CTOR), methanol

to oil molar ratio (MOMR) and reaction time on sunflower oil transesterification were studied on

the most active catalyst: Ca600/Mg4Al2HT. The mixing intensity was kept at 400 rpm and the

reaction temperature at 60 ˚C.

2.5.1 Effect of catalyst to oil ratio

Catalyst loading significantly affects the FAME (Fatty Acid Methyl Ester) yield. An insufficient

amount of catalyst results in an incomplete reaction [15]. When the catalyst percentage is

increased, the yield is enhanced. As the catalyst weight increases, the number of O2- anion sites

increases which adsorb more H+ from methanol forming active centers [91]. The reactants will

therefore have more contact with these active centers thus increasing the yield. In this work, as the

catalyst amount increased from 2.5 wt% to 10 wt%, the yield increased from 70% to 73% as shown

in Figure 2.8. It appears that the additional amount of catalyst added had no significant effect on

the yield. In some cases, a larger quantity of catalyst in the same amount of oil and methanol could

result in a poorer mixing given the higher viscosity of the mixture of reactants and catalyst. The

resulting increased mass transfer resistance could limit the FAME yield [21]. The optimum catalyst

amount used in the following experiments for biodiesel synthesis was therefore chosen to be 2.5

wt%.

70 73

0

20

40

60

80

100Effect of CTOR: Ca 600/Mg

4Al

2HT

10 wt%

FA

ME

yie

ld (

%)

2.5 wt%

Figure 2.8: Effect of catalyst to oil ratio on FAME yield over the Ca600/Mg4Al2HT catalyst (Stirring rate of 400 rpm, 12:1

MOMR and temperature of 60˚C for 4 hours)

50

2.5.2 Effect of alcohol to oil ratio

The alcohol to oil ratio plays a significant role in the production of biodiesel. For the

transesterification reaction to occur, a ratio of 3:1 of alcohol to oil should be present [79]. Excess

alcohol or methanol can be added to shift the equilibrium forward, driving the reaction near

completion and maximizing the production of the needed product in a short period of time. It has

been reported that the biodiesel yield increases as the amount of methanol increases in the reaction

medium [95]. This was clearly demonstrated in our results (Figure 2.9) where an increase in the

methanol to oil ratio from 12:1 to 15:1 under the same conditions (stirring rate of 400 rpm, 2.5

wt% catalysts to oil and temperature of 60 ˚C), improved the yield considerably from 70% to 92%.

The optimum molar ratio of methanol to sunflower oil was therefore found to be 15:1.

70

92

0

20

40

60

80

100

Methanol:Oil=15:1

Effect of Alcohol to Oil ratio:

Ca 600/Mg4Al

2HT

FA

ME

yie

ld (

%)

Methanol:Oil=12:1

Figure 2.9: Effect of alcohol to oil ratio on FAME yield over the Ca600/Mg4Al2HT catalyst (Stirring rate of 400 rpm, 2.5

wt% CTOR, and temperature of 60 ˚C for 4 hours)

2.5.3 Effect of reaction time

The effect of reaction time on the yield was also investigated (Figure 2.10). Conversion of the ester

in transesterification reactions proceeds near completion if it is given enough time [9].The reaction

starts slow when dispersing the alcohol into the oil but speeds up after a certain time. As the

reaction time increases, the conversion increases significantly. An experiment was conducted

where the time was increased from 4 to 6 hours under the same conditions (stirring rate of 400

rpm, 2.5 wt% catalysts to oil and temperature of 60 ˚C). Results showed that the conversion

increased steadily from 70% to 83 %. Molecules were given enough time to collide and react with

each other causing an increase in the reaction rate.

51

70

83

0

20

40

60

80

100Effect of Reaction Time:

Ca 600/Mg4Al

2HT

6 hours

FA

ME

yie

ld (

%)

4 hours

Figure 2.10: Effect of reaction time on FAME yield over the Ca600/Mg4Al2HT catalyst (Stirring rate of 400 rpm, 2.5 wt%

CTOR, 12:1 MOMR and temperature of 60 ˚C)

2.5.4 Combined effect of alcohol to oil ratio and the reaction time

Transesterification was also performed over the most active Ca600/Mg4Al2HT catalyst to study

the combined effect of increasing the methanol:oil ratio from 12:1 to 15:1 along with the time of

reaction from 4 hours to 6 hours (Figure 2.11). The reaction was left for 6 hours with 2.5 wt%

catalyst, stirring rate of 400 rpm, 15:1 methanol to oil ratio, and a temperature of 60 ˚C. The

obtained results showed that increasing methanol to oil ratio decreased the mixture viscosity and

enhanced mass transfer [9]. The highest yield 95% was obtained by controlling both of these

parameters.

95

0

20

40

60

80

100

Ca 600/Mg4Al

2HT

FA

ME

yie

ld (

%)

2.5

wt%

Meth

anol:O

il =

15:1

6 h

ours

Figure 2.11: Combined effect on FAME yield over the Ca600/Mg4Al2HT catalyst (Stirring rate of 400 rpm, 2.5 wt%

CTOR, 12:1 MOMR and temperature of 60 ˚C)

52

2.6 Characterization of produced biodiesel and comparison to petroleum diesel

The equipment used and test methods used to determine fuel properties are found in Appendix A.

The properties of the produced biodiesel were analyzed and results are presented in Table 2.5. The

properties of conventional petroleum diesel (purchased from an IPT gas station in Lebanon) were

also analyzed for comparison and are shown in Table 2.6. The results obtained for both fuels are

in agreement with the ASTM requirements. The density of biodiesel is higher than that of diesel.

In addition, the flash point of biodiesel is also higher than that of diesel which makes the handling

and transportation of biodiesel much easier. The corrosion index is the same for both fuels

implying that neither of these fuels will cause corrosion to metal parts when used in diesel engines.

Moreover, the kinematic viscosity of biodiesel is higher than that of petroleum diesel in agreement

with literature.

Table 2.5: Properties of produced biodiesel

Property Value ASTM

requirement

Test method

Density at 25 ˚C

(Kg/m3)

885 860-900 ASTM D4052

Kinematic viscosity

at 40˚C (mm2/s)

4.7 1.9-6.0 ASTM D445

Flash point (˚C) 93 93 min ASTM D39

Copper strip

corrosion

3 3 ASTM D130

Acid number

(mgKOH/g)

0.03 0.05 max AOCS Ca-5a-40

Table 2.6: Properties of conventional petroleum diesel

Property Value ASTM

requirement

Test method

Density at 25 ˚C

(Kg/m3)

831.5 820-845 ASTM D4052

Kinematic viscosity

at 40˚C (mm2/s)

3.3 1.3-4.1 ASTM D445

Flash point (˚C) 65 55 min ASTM D39

Copper strip

corrosion

3 3 ASTM D130

53

2.7 Conclusion

Hydrotalcite derived mixed oxides were studied in the transesterification reaction for biodiesel

production. The Ca600/Mg4Al2HT catalyst showed the best catalytic performance and verified that

Ca addition is essential for the creation of active sites. Moreover, the findings of this chapter

demonstrated that the thermal treatment of the support/catalyst and the preparation method for Ca

incorporation both play a major role. Impregnation of the Ca active phase on the Mg4Al2 support

followed by a thermal treatment was clearly the best preparation method as the highest yield was

obtained over this catalyst. This could be attributed to an enhancement of basic properties as shown

from characterization results. Following a series of experiments, the optimum conditions for

biodiesel synthesis were: a methanol to oil molar ratio of 15:1, a catalyst weight of 2.5 wt% and a

reaction time of 6 hours which gave a FAME yield of 95%. The properties of the produced

biodiesel were also analyzed and are in agreement with the ASTM requirements.

In the next two chapters, the focus will be on the valorization of glycerol, the by-product obtained

during biodiesel production, by the steam reforming reaction to produce hydrogen.

54

Chapter 3: Steam reforming of glycerol over

Ni-based pure metal oxides

55

In this chapter, the activity of Ni-based pure metal oxides was first studied in the steam reforming

of glycerol reaction. The supports utilized were all commercial and impregnated with 5 wt% Ni.

In the first section, the effect of using different supports (CeO2, Y2O3 and ZrO2) was studied. The

second section focused on parameter modification of the most efficient Ni-based pure metal oxide

followed by stability tests. This study was ultimately done to help select the most efficient Ni-

based pure metal oxide to later on compare its activity to the most efficient Ni-based hydrotalcite

derived mixed oxide in Chapter 4.

3.1 Influence of the support on catalytic performance

In this section, different Ni-based pure metal oxides will be studied. The physico-chemical

properties of all catalysts will be compared as well as their catalytic activities in the glycerol steam

reforming (GSR) reaction.

3.1.1 Synthesis of Ni-based pure metal oxides

Commercial supports: CeO2, Y2O3 and ZrO2 were purchased from Sigma Aldrich. The supported

Ni catalysts were prepared by wet impregnation on the different supports. An adequate amount of

nickel nitrate Ni(NO3)2.6H2O (Sigma-Aldrich, 98% purity) was dissolved in 50 mL of deionized

water in order to obtain 5 wt% of Ni in the final solid. This solution was then added to 1 g of the

support and mixed during two hours. The excess of water was slowly eliminated in a rotary

evaporator (60°C, under vacuum). The obtained residue was kept in an oven at 60°C during 24 h.

The dried solids were then thermally stabilized by calcination at 600°C (1°C/min) under an air

flow for four hours. The calcination temperature of 600oC was chosen based on literature study.

For active phases such as Ni, the calcination temperature is a key parameter as it governs the

concentration of surface Ni as well as the interaction between the metal and support which could

later influence catalytic activity. The temperature of 600oC was found to be optimum and was

chosen on this basis [96]. The prepared solids were labelled: Ni/CeO2, Ni/Y2O3 and Ni/ZrO2.

3.1.2 Characterization

The commercial supports alone were characterized as well as the calcined Ni impregnated

catalysts. Characterization results were compared to study the influence of Ni addition on the

support in addition to the influence of each support in the GSR reaction. The various

characterization techniques used, their description and conditions are presented in Appendix B.

56

3.1.2.1 X-ray diffractograms (XRD) of supports alone and Ni impregnated catalysts

X-ray diffractograms of the supports alone and the Ni impregnated catalysts are shown in Figure

3.1(a) and (b) respectively.

X-ray diffractograms of the supports alone verified the presence of the phases corresponding to

each support:

CeO2 (C) (JCPDS 34-0394) with fluorite structure [97]

Y2O3 (Y) (JCPDS 41-1105) of cubic phase [98]

ZrO2 (Z) (JCPDS 37-1484) with the monoclinic phase [99]

Figure 3.1: XRD diffractograms of (a) supports alone and (b) Ni impregnated catalysts (c) magnified 36-45o region for Ni

impregnated catalysts

20 30 40 50 60 70 80

Y

NiO

Ni/ZrO2

Ni/CeO2

Inte

nsity (

a.u

.)

2(o)

Ni/Y2O

3

(b)

C

CC C

C C C C

Y

YY Y Y

Y

Y

YY

Z

Z

Z

ZZ

ZZ Z

ZZ Z Z Z Z

36 37 38 39 40 41 42 43 44 45

Z

YY

Z NiO

Inte

nsity (

a.u

.)

2(o)

Ni/CeO2

Ni/Y2O

3

Ni/ZrO2

(c)

Z

20 30 40 50 60 70 80

Y YYY

YY YY

Y

Y

CCCCCC

C

Inte

nsity (

a.u

.)

2(o)

ZrO2

Y2O

3

CeO2

(a)

C

Z

Z

Z

ZZ

ZZ Z

ZZ Z Z Z Z

57

Upon comparison with the diffractograms of the calcined Ni-impregnated catalysts, it is noted that

the peaks corresponding to the CeO2 phase and ZrO2 phase became less intense indicating a

decreased crystallinity of the supports. As for the Y2O3 support, Ni impregnation resulted in more

intense Y2O3 peaks as well as a shift in 2 angles to higher values. This shift can be attributed to

the contraction of the Y2O3 lattice possibly due to Ni2+ diffusion. Since the Ni2+ (0.78Å) is smaller

than Y3+ (0.89Å) and both of them have a cubic structure, it is possible that its diffusion into the

Y2O3 lattice led to a contraction [4].

To take a closer look at NiO (JCPDS 44-1159) crystallites over the Ni impregnated catalysts, the

36-45o region was magnified and is shown in Figure 3.1(c). Ni/CeO2 and Ni/ZrO2 both display

peaks corresponding to NiO (JCPDS 44-1159) whereas no such peaks are observed on the Ni/Y2O3

catalyst. This indicates that Ni on the Y2O3 support is more dispersed [101]. This further verifies

that some Ni2+ ions could have diffused into the Y2O3 lattice and caused its contraction.

Diffraction peaks of NiO are weaker for Ni/CeO2 indicating the presence of smaller NiO

crystallites than in Ni/ZrO2 [68]. In upcoming sections, the Ni particle size of each catalyst is

deduced on account of the intensity of the peaks assigned to NiO [68].

3.1.2.2 Specific surface area analysis (BET) of supports alone and Ni impregnated catalysts

Specific surface areas of the supports alone and their corresponding calcined Ni-impregnated

solids are shown in Figure 3.2.

40

31 30

38

5 6

0

10

20

30

40

50

Specific

Surf

ace A

rea (

m2/g

)

CeO2

Ni/CeO2

Y2O

3Ni/Y

2O

3ZrO

2Ni/ZrO

2

Figure 3.2: Specific surface areas of supports alone and Ni impregnated catalysts

Generally speaking, surface areas of all catalysts are low due to the high calcination temperature

of the commercial supports. CeO2 and Y2O3 exhibit higher surface areas than ZrO2. Monoclinic

58

zirconia calcined at high temperature is known to present low surface areas [99]. Variations in the

specific surface areas of the Ni impregnated catalysts compared to the supports alone are due to

the modification of the surface texture following impregnation.

59

3.1.2.3 H2-Temperature Programmed Reduction (H2-TPR) of supports alone and Ni


H2-TPR profiles of each support and its corresponding Ni impregnated catalyst are shown in Figure

3.3.

Figure 3.3: H2-TPR profiles of (a) CeO2, Ni/CeO2 (b) Y2O3, Ni/Y2O3 and (c) ZrO2, Ni/ZrO2

100 200 300 400 500 600 700

Inte

nsity (

a.u

.)

Temperature (oC)

Ni/CeO2

CeO2

213oC

292oC

170oC

(a)

100 200 300 400 500 600 700

(b)

Inte

nsity (

a.u

.)

Temperature (oC)

Ni/Y2O

3

Y2O

3

214oC

330oC

395oC

100 200 300 400 500 600 700

(c)

Ni/ZrO2In

tensity (

a.u

.)

Temperature (oC)

ZrO2

365oC

60

TPR profiles of all the supports alone show weak reduction peaks revealing that the supports are

barely reduced in the studied temperature range. The Ni impregnated catalysts exhibit different

reduction peaks all corresponding to the reduction of NiO. The reduction temperatures vary

depending on the nature of the support.

For the CeO2 support, the lower temperature reduction peaks observed from 300-500oC are due to

the reduction of surface ceria. A higher temperature reduction peak generally reported for CeO2

beyond our studied temperature range ascribed to the reduction of bulk ceria [102], [103]. For the

Ni/CeO2 catalyst, 3 reduction peaks are observed. The first two peaks at 170oC and 213oC

correspond to the reduction of NiO weakly interacting with surface ceria. The peak at 292oC

corresponds to the reduction of NiO in better interaction with the CeO2 support.

The reduction peaks observed for Y2O3 are an indication of the presence of surface oxygen species.

Hydrogen molecules from the reducing gas mixture consume this oxygen present on the surface

of the Y2O3 support [104], [105]. Ni/Y2O3 exhibits 3 reduction peaks. These peaks can be attributed

to the reduction of NiO promoted by oxygen vacancies (214oC), reduction of free NiO (330oC)

and reduction of NiO interacting with Y2O3 (395oC) [106]. In fact, the high temperature reduction

peak for Ni/Y2O3 could be more specifically due to the reduction of NiO that has diffused into the

Y2O3 lattice (as discussed in the XRD) [100].

The small reduction peaks present for ZrO2 can also be attributed to surface oxygen species. For

Ni/ZrO2, a single reduction peak (365oC) is observed corresponding to the reduction of NiO

interacting with ZrO2 [107].

The major reduction peak corresponding to NiO interacting with the support occurs at higher

temperatures for the Ni/ZrO2 (365oC) and Ni/Y2O3 catalyst (395oC) indicating a stronger metal-

support interaction compared to Ni/CeO2 (292oC).

The TPR analysis showed that the nature of the support affects the stability of the active phase. In

order to achieve a stable active phase, there should be strong interactions between the support and

metal.

The experimental hydrogen consumption values are shown in Table 3.1.

61

Table 3.1: Hydrogen consumption values of Ni impregnated catalysts

For Ni/CeO2 and Ni/ZrO2, experimental hydrogen consumption values are close to the theoretical

value ((896 µmol/g) whereas Ni/Y2O3 shows a higher value. This might be due to the partial

reduction of Y2O3 facilitated by the hydrogen spillover on reduced Ni particles [108].

3.1.2.4 CO2-Temperature Programmed Desorption (CO2-TPD) of supports alone and Ni


Figure 3.4 shows the CO2-TPD profiles of the supports alone and Ni impregnated catalysts.

100 200 300 400 500 600 700

Ni/ZrO2

Ni/Y2O

3

Ni/CeO2

ZrO2

Y2O

3

CeO2

Inte

nsity (

a.u

.)

Temperature (oC)

I

Figure 3.4: CO2-TPD profiles of supports alone and Ni impregnated catalysts

The profiles of the Ni impregnated calcined supports show little to no difference compared to the

profiles of the supports alone. Weak basic sites (I) corresponding to surface hydroxyl groups are

observed only for the Ni/Y2O3 and Ni/CeO2 catalysts. Basicity values in Table 3.2 show that the

Ni/CeO2 and Ni/Y2O3 have a similar amount of basic sites.

Catalyst H2 consumption

(μmol H2/gcatalyst)

Ni/CeO2 789

Ni/Y2O3 1328

Ni/ZrO2 781

62

Table 3.2: Basicity values of Ni impregnated catalysts

It is important to note that during sample handling Y2O2CO3 can be formed by adsorption of CO2

from the atmosphere on the basic sites of Y2O3. At higher temperatures, this carbonate decomposes

which could give misleading results especially during CO2-TPD analysis. Therefore, all samples

were pretreated at 600oC before the CO2-TPD analysis (detailed in Appendix B).

3.1.3 Steam reforming of glycerol results

Details of the catalytic test set-up, test conditions and reaction parameters of the steam reforming

of glycerol are in Appendix B.

3.1.3.1 The uncatalyzed reaction

Prior to catalytic testing, a preliminary test without catalyst (blank run) was conducted at the

studied temperature range (400-700oC) to study the extent of glycerol pyrolysis in our operating

conditions (WGFR 9:1, flow rate: 0.025 mL/min). Glycerol is a thermally unstable molecule and

it is therefore crucial to understand the contribution of its pyrolysis in the reforming reaction.

Figure 3.5 shows H2 yields, glycerol conversion to gaseous products, and CO2, CO, CH4

selectivities obtained with an empty reactor. The total glycerol conversions and the concentrations

of some obtained liquid by-products are shown in Appendix B.

Catalyst Basicity

(μmol CO2/ gcatalyst)

Ni/CeO2 99

Ni/Y2O3 105

Ni/ZrO2 -

400 450 500 550 600 650 7000

10

20

30

40

50

60

70

80

90

100(b) Empty Reactor

Gly

cero

l C

onvers

ion G

aseous (

%)

Temperature (oC)

400 450 500 550 600 650 7000

1

2

3

4

5

6

7 Empty Reactor

Hydro

gen y

ield

(m

ol H

2/m

ol gly

cero

l)

Temperature (oC)

(a)

63

Results confirm that glycerol decomposes in a mixture of both gaseous and liquid products [109].

The gaseous phase is essentially composed of H2, CO2, CO and CH4. In the absence of a catalyst,

glycerol conversion to gaseous products and hydrogen yields are notably very low. Nevertheless,

higher temperatures are enough to lead to the cracking of glycerol based on the following

endothermic reaction:

C3H8O3 3CO + 4H2 ΔH298K = 251 kJ/mol

Decreasing CO2 selectivities and increasing CO and CH4 selectivities are observed verifying the

occurrence of the glycerol cracking reaction as well as side methanation reactions such as glycerol

hydrogenolysis where glycerol decomposes to methane:

Figure 3.5: (a) H2 yields (b) Glycerol conversion to gaseous products (c) CO2 (d) CO (e) CH4 selectivities of GSR

reaction over empty reactor (flow rate: 0.025 mL/min, steam/glycerol molar ratio: 9)

400 450 500 550 600 650 7000

10

20

30

40

50

60

70

80

90

100(c) Empty Reactor

CO

2 S

ele

ctivity (

%)

Temperature (oC)

400 450 500 550 600 650 7000

10

20

30

40

50

60

70

80

90

100(d) Empty Reactor

CO

Sele

ctivity (

%)

Temperature (oC)

400 450 500 550 600 650 7000

10

20

30

40

50

60

70

80

90

100 Empty Reactor

CH

4 S

ele

ctivity (

%)

Temperature (oC)

(e)

64

C3H8O3 + 2H2 2CH4 + CO + 2H2O ΔH298K = -160 kJ/mol

The occurrence of the reverse water gas-shift reaction also explains the CO2 and CO selectivity

trends:

CO2 + H2 CO + H2O ΔH298K = 41 kJ/mol

Total glycerol conversions were obtained from liquid product analysis and the figures are in

Appendix B. There is a significant amount of total glycerol conversion which increases as

temperature increases. Since values of glycerol conversion to gaseous products were quite low,

this indicates the formation of a significant quantity of liquid by-products. As discussed in Chapter

1, the complexity of the GSR process leads to the formation of a wide distribution of liquid

products. Concentration values of all studied by-products (also in Appendix B) exhibited an

increasing trend for the blank test.

The results of this blank test verify that in the absence of a catalyst, the glycerol cracking reaction

occurs and this also explains the high concentration of liquid products observed. As will be

demonstrated in the upcoming sections, in the presence of a catalyst, the gaseous phase is still

composed of mainly CO, H2, CO2 and CH4 but the product distribution (both gaseous and liquid)

will vary depending on the catalyst. In fact, according to literature, the use of a catalyst shifts

product distribution in favor of more hydrogen production and less formation of 1C-2C

hydrocarbons [110].

65

3.1.3.2 The catalyzed reaction: Steam reforming of glycerol over Ni/CeO2, Ni/Y2O3 and Ni/ZrO2

Figure 3.6 shows the H2 yields, glycerol conversion to gaseous products, and CO2, CO, CH4

selectivities over the Ni/CeO2, Ni/Y2O3 and Ni/ZrO2 catalysts.


reaction over Ni/CeO2, Ni/Y2O3 and Ni/ZrO2 (flow rate: 0.025 mL/min, steam/glycerol molar ratio: 9)

400 450 500 550 600 650 7000

10

20

30

40

50

60

70

80

90

100(c) Ni/CeO

2

Ni/Y2O

3

Ni/ZrO2

CO

2 S

ele

ctivity (

%)

Temperature (oC)

400 450 500 550 600 650 7000

10

20(e) Ni/CeO

2

Ni/Y2O

3

Ni/ZrO2

CH

4 S

ele

ctivity (

%)

Temperature (oC)

400 450 500 550 600 650 7000

10

20

30

40

50

60

70

80

90

100(d)

Ni/CeO2

Ni/Y2O

3

Ni/ZrO2

CO

Sele

ctivity (

%)

Temperature (oC)

400 450 500 550 600 650 7000.0

0.2

0.4

0.6

0.8

1.0

1.2

1.4

1.6

1.8

2.0(a)

Hydro

gen y

ield

(m

ol H

2/m

ol gly

cero

l)

Temperature (oC)

Ni/CeO2

Ni/Y2O

3

Ni/ZrO2

400 450 500 550 600 650 7000

10

20

30

40

50(b)

Ni/CeO2

Ni/Y2O

3

Ni/ZrO2

Gly

cero

l C

onvers

ion G

aseous (

%)

Temperature (oC)

66

The catalytic activities of the supports alone are not presented as they were similar to the blank

test results due to the absence of an active phase. The displayed results indicate that the catalytic

activity varies in the following order: Ni/ZrO2 > Ni/Y2O3 > Ni/CeO2 in terms of glycerol

conversions to gaseous products and hydrogen yields. The increase in glycerol conversion to

gaseous products can be directly correlated to the increase in hydrogen yield.

Ni/Y2O3 and Ni/CeO2 exhibited similar trends in terms of CO2, CO and CH4 selectivities. From

400-500oC, the dominance of the water gas-shift reaction is clear over Ni/Y2O3 and Ni/CeO2 as

evidenced by increasing CO2 and decreasing CO selectivities:

CO + H2O CO2 + H2 ΔH298K = -41 kJ/mol

The CeO2 support facilitates the water-gas shift reaction. This is due to the redox properties and

high oxygen storage capacity of CeO2 which influences the activation of H2O. The OH groups

formed from the dissociation of H2O remove absorbed CO and produce CO2 and H2 at the same

time [111]. This also appears to be the case for the Ni/Y2O3 catalyst where the Y2O3 support

facilitated the water gas shift reaction [50].

Beyond 500oC, dominance of the endothermic reverse water gas shift reaction over both catalysts

causes CO2 selectivities to decrease and CO selectivities to increase. This is facilitated by the basic

sites of Ni/CeO2 and Ni/Y2O3 (shown in CO2-TPD, Figure 3.4) which favor CO2 adsorption:


In terms of CH4 selectivities, both experience increasing trends as a result of methane-producing

(methanation) reactions that are feasible over the entire temperature range:

CO + 3H2 CH4 + H2O ΔH298K = -206 kJ/mol

CO2 + 4H2 CH4 + 2H2O ΔH298K = -165 kJ/mol


At higher temperatures, both catalysts did not have sufficient capacity to reform this produced CH4

into H2 and CO via the dry or steam reforming of methane [112].

67

For the Ni/ZrO2 catalyst, a different mechanism explains the different selectivity trends observed

over this catalyst. Most importantly, the decomposition of glycerol reaction is clearly more favored

over Ni/ZrO2:

C3H8O3 4H2 + 3CO ΔH298K = 251 kJ/mol

At lower temperatures, the hydrogenolysis of glycerol reaction is also particularly dominant:


This explains the notably high CO selectivities observed at lower temperatures as well as the high

CH4 selectivity over Ni/ZrO2. The produced CO and H2O prompt the water gas shift reaction at

higher temperatures resulting in increasing CO2 selectivities. The produced CO could also be

converted to CO2 via the Boudouard reaction:

2CO CO2 + C ΔH298K = -172 kJ/mol

The relatively higher quantity of CH4 produced by methanation reactions over Ni/ZrO2 could also

be reformed by the endothermic methane decomposition reaction explaining the relatively stable

CH4 selectivity as well as the increasing hydrogen yields:

CH4 2H2 + C ΔH298K = 75 kJ/mol

It is important to note that the higher production of H2 and CO over the Ni/ZrO2 makes it attractive

for syngas production.

3.1.4 Liquid Product Analysis

The figures showing the results of the liquid product analysis are in Appendix B. In terms of total

glycerol conversions, the Ni/ZrO2 catalyst exhibited the highest total glycerol conversion and was

the closest to thermodynamic predictions by almost reaching 100% total conversion at 700oC. In

terms of liquid product concentrations, different trends were observed over each catalyst. For

Ni/CeO2, as temperature increased, the concentration of liquid products continued to increase. This

increasing trend was similar to that observed in the blank test which indicates the occurrence of

glycerol pyrolysis and not much contribution from the Ni/CeO2 catalyst. This correlates with its

low catalytic activity. The Ni/Y2O3 catalyst experienced the highest production of liquid by-

products. Rare earth metal oxides are known to favor the dehydrogenation of alcohols. In fact,

68

Y2O3 is more active for dehydrogenation than oxides of other lanthanides [113]. The acidic sites

on the Y2O3 support also participate in the dehydration of the glycerol molecule [101]. This

explains the higher concentration of liquid products obtained over the Y2O3 support especially

acetaldehyde and hydroxyacetone (both dehydration products). The Ni/ZrO2 exhibited a

decreasing trend of liquid product concentrations with increasing temperature verifying its ability

to convert liquid by-products to gaseous products (H2, CO, CO2, CH4). In fact, the higher

acetaldehyde concentrations formed could also explain the higher CH4 and CO selectivities. The

decomposition of acetaldehyde leads to the formation of methane and carbon monoxide [114].

3.1.5 Characterization of used catalysts – Thermal analysis (TG-DSC)

To study coke deposition over the catalysts, characterization of the used catalysts was also

conducted. The TG and DSC curves recorded for the used catalysts are presented in Figure 3.7(a)

and (b) respectively.

TG curves of the spent catalysts indicate weight losses due to the oxidation of coke. More coke

was formed on the Ni/ZrO2 catalyst due to its higher activity. The Ni/CeO2 produced the least

amount of coke given its low activity. Coke formation over Ni/Y2O3 and Ni/CeO2 was significantly

low compared to that of Ni/ZrO2 as the basic sites from the Y2O3 and CeO2 supports (shown in

CO2-TPD, Figure 3.4) could have also selectively oxidized carbon species [104].

The DSC profiles in Figure 3.7(b) show exothermic peaks that correspond to the different types of

coke formed. As discussed previously in Chapter 1, coke oxidized at a lower temperature (<500oC)

Figure 3.7: (a) TG and (b) DSC curves obtained for used Ni/CeO2, Ni/Y2O3 and Ni/ZrO2

100 200 300 400 500 600 700 800

(b) Ni/ZrO2

Ni/Y2O

3

Ni/CeO2

555oC

570oC

D

SC

sig

nal (a

.u.)

Temperature (oC)

586oC

100 200 300 400 500 600 700 800

III

II I

I

DS

C s

ign

al (a

.u.)

Temperature (oC)

II

100 200 300 400 500 600 700 80080

85

90

95

100

105(a)

0.3%

1.7%

Ni/ZrO2

Ni/Y2O

3

Ni/CeO2

We

igh

t L

oss (

%)

Temperature (oC)

16%

69

is generally referred to as encapsulating or non-filamentous coke whereas coke oxidized at a higher

temperature (>500oC) is whisker-like or filamentous coke [84]. Comparing DSC profiles of

Ni/CeO2 and Ni/Y2O3, it is observed that the Ni/CeO2 catalyst produced three types of coke

whereas Ni/Y2O3 produced two. The Ni/ZrO2 catalyst produced only one type of coke

(filamentous). This filamentous coke (I) is also produced by Ni/Y2O3 and Ni/CeO2 but is oxidized

at lower temperatures. Additionally, non-filamentous coke was also formed on these catalysts: II

for Ni/Y2O3; II and III for Ni/CeO2.

According to literature, there are several routes for coke formation from different side or secondary

reactions [115]:

Condensation and polymerization of oxygenates formed as by-products

Boudouard reaction (2CO CO2 + C)

Methane decomposition reaction (CH4 2H2 + C)

Furthermore, it has been reported by several authors that encapsulating coke is formed by the

condensation of oxygenates. Dehydration products in particular formed on acidic sites evolve into

coke precursors for encapsulating coke. Whereas filamentous coke is formed from CO via the

Boudouard reaction and CH4 via methane decomposition [116].

It is important to note that the large quantity of filamentous coke produced over the Ni/ZrO2

catalyst verifies the occurrence of the Boudouard reaction and methane decomposition reaction as

stated previously in the catalytic test results since CO and CH4 are known precursors of

filamentous coke. Furthermore, the amorphous or encapsulating coke formed on the Ni/Y2O3 and

Ni/CeO2 catalysts could be due to the higher liquid by-product concentrations as some of these by-

products are precursors of encapsulating coke [116].

Ni particle size also influences the coke formation mechanism [117]. It has been reported in

literature that larger Ni particles favor the growth of filamentous coke whereas smaller ones favor

the growth of non-filamentous coke [116]. X-ray diffractograms showed that Ni/ZrO2 had larger

NiO crystallites than Ni/Y2O3 and Ni/CeO2 and coke analysis verified that filamentous coke was

mostly produced over the Ni/ZrO2 catalyst whereas non-filamentous coke was formed over the

Ni/Y2O3 and Ni/CeO2 catalysts.

70

3.1.6 Discussion

In this section, the influence of the support on catalytic performance was studied for Ni-based pure

metal oxides. X-ray diffractograms revealed different NiO sizes depending on the support. The

specific surface areas of Ni/CeO2 and Ni/Y2O3 were larger than Ni/ZrO2. Although higher surface

areas are often an indication of a more active catalyst, in this case the specific surface area played

no significant role in catalytic activity. CO2-TPD results showed higher basicities for Ni/CeO2 and

Ni/Y2O3 compared to Ni/ZrO2 as the CeO2 and Y2O3 supports both possessed weak basic sites.

H2-TPR profiles revealed varying stabilities of the active phase depending on the nature of the

support. Test results demonstrated that the type of support has a significant impact on the reaction

pathway and distribution of obtained products [113]. The Ni/ZrO2 catalyst was the most active in

terms of hydrogen yields and glycerol conversions to gaseous products. Liquid product analysis

revealed a decreasing trend of liquid product concentrations for the Ni/ZrO2 catalyst with

increasing temperatures further validating its ability to reform oxygenated hydrocarbons to

gaseous products. Thermal analysis of the used catalysts helped study the coke deposition over

each catalyst. Filamentous coke was only produced over the Ni/ZrO2 catalyst whereas both

filamentous and non-filamentous coke were produced over the Ni/Y2O3 and Ni/CeO2 catalysts.

The acidity of the support plays a role in the formation of by-products which are precursors of

encapsulating coke [27]. In addition, the basic properties of the support also affect the ability of

Ni-based catalysts to resist coke deposition [53].

3.2 Parameter Modification: Influence of zirconia phase on catalytic

activity

Based on the results of the previous section, the Ni/ZrO2 exhibited the highest hydrogen yields and

glycerol conversion to gaseous products. Zirconia has been used as a catalyst both as a single oxide

and mixed with other metal oxides. It is very stable at high temperatures and this makes it suitable

to use as a catalyst support especially in endothermic reactions which are performed at high

temperatures [118].

Zirconia has three crystalline phases: monoclinic, tetragonal and cubic (Figure 3.8). The

monoclinic phase is stable up to 1200oC, the tetragonal is stable up to 1900oC and the cubic is

stable above 1900oC. The properties of ZrO2 based materials depend on the crystalline phase [99].

71

Figure 3.8: Schematic of different zirconia phases

Many authors have found different ways of preparing the tetragonal phase without the need for

thermal transformation beyond 1200oC. The aim of this section is to prepare the tetragonal zirconia

phase via variation of the synthesis method and compare its activity to the monoclinic zirconia

phase in GSR.

3.2.1 Synthesis of monoclinic and tetragonal Ni/ZrO2

The same commercial monoclinic zirconia support studied in the previous section was used.

Hydrolysis is the most common method used for synthesizing tetragonal zirconia polycrystals.

These crystals are produced by the calcination of hydrous zirconia particles. For synthesis of the

tetragonal zirconia, zirconyl chloride octahydrate was used as the zirconium precursor. Ethylene

diamine was used as a precipitating agent and colloidal protecting agent (bifunctional). As a

precipitating agent, it precipitates ZrOCl2 into ZrO(OH)2 and acts as a colloidal protecting agent

for ZrO(OH)2 [119].

Based on the method reported by D’souza et al. [120], 500 mL of an aqueous solution of 45 mM

ZrOCl2.8H2O was hydrolyzed with 6.2 mL of ethylene diamine (H2NCH2CH2NH2) in a 1 L round

bottomed flask under vigorous stirring. After precipitation, the obtained solution was stirred for an

additional 30 minutes then placed in an oven at 110oC for 48 hours of digestion (precipitate

ageing). After digestion, the hydroxide precipitate was filtered on a Buchner funnel and washed

twice with 500 mL of water to ensure complete removal of all chloride ions. The absence of

chloride ions was verified by the addition of a few drops of silver nitrate (AgNO3) to the filtrate

liquid to make sure no AgCl is formed. The obtained solid was then dried in the oven at 110oC for

12 hours resulting in a white sample. The hydroxide was then calcined at 700oC in a furnace under

a flow of air with a heating rate of 1oC/min.

72

The Ni-based catalysts were prepared by wet impregnation of the monoclinic and tetragonal

supports with an aqueous solution of nickel nitrate to obtain a 5 wt% loading of Ni. The obtained

solutions were stirred for 2 hours, slowly evaporated under vacuum in a rotary evaporator and then

dried at 60oC to ensure complete liquid removal. After impregnation, the catalysts were calcined

at 600oC for 4 hours (heating rate 1oC/min) under a flow of air. The Ni/ZrO2 catalysts were labelled

as Ni/ZrO2 (M) (monoclinic zirconia) and Ni/ZrO2 (T) (tetragonal zirconia).

3.2.2 Characterization of calcined Ni/ZrO2 catalysts

3.2.2.1 X-ray diffractograms (XRD) of calcined Ni/ZrO2 catalysts

X-ray diffractograms of the calcined Ni/ZrO2 catalysts are shown in Figure 3.9.

X-ray diffractograms verified the presence of the monoclinic phase (M) (JCPDS 37-1484) for the

monoclinic catalyst and the tetragonal phase (T) (JCPDS 50-1089) for the tetragonal catalyst [121].

Diffraction peaks of NiO are weaker for the tetragonal catalyst indicating the presence of smaller

NiO crystallites than the monoclinic catalyst. They could also be an indication of an improved

dispersion of Ni on the tetragonal catalyst [122].

3.2.2.2 Specific surface area analysis (BET) of calcined Ni/ZrO2 catalysts

Specific surface areas of both catalysts are compared in Figure 3.10.

Figure 3.9: XRD diffractograms of (a) Ni/ZrO2 (M) and Ni/ZrO2 (T) catalysts (b) magnified 36-45o region

20 30 40 50 60 70 80

NiO

Ni/ZrO2 (T)

Ni/ZrO2 (M)

M MMMM

M

MM

M

M MM

M

M

T

TT

T

Inte

nsity (

a.u

.)

2(o)

(a)

36 37 38 39 40 41 42 43 44 45

(b)

Ni/ZrO2 (M)

Ni/ZrO2 (T)

NiO

Inte

nsity (

a.u

.)

2(o)

73

0

20

40

60

80

100

69

Ni/ZrO2 (T)

Spe

cific

Surf

ace A

rea

(m

2/g

)

Ni/ZrO2 (M)

6

Figure 3.10: Specific surface areas of Ni/ZrO2 (M) and Ni/ZrO2 (T)

The monoclinic catalyst exhibits a very low specific surface area compared to the tetragonal

catalyst. Once again, this is due to the high calcination temperature of the monoclinic zirconia as

it is a commercial support. The tetragonal phase in Ni/ZrO2 (T) increases the surface area of the

catalyst [99].

3.2.2.3 H2-Temperature Programmed Reduction (H2-TPR) of calcined Ni/ZrO2 catalysts

H2-TPR profiles of both catalysts are shown in Figure 3.11.

100 200 300 400 500 600 700

Ni/ZrO2 (T)

Ni/ZrO2 (M)

Inte

nsity (

a.u

.)

Temperature (oC)

364oC

376oC 574

oC

Figure 3.11: H2-TPR profiles of Ni/ZrO2 (M) and Ni/ZrO2 (T)

The monoclinic catalyst exhibits a single reduction peak at 364oC corresponding to the reduction

of NiO interacting with the support. The tetragonal catalyst exhibits two reduction peaks (376 and

574oC), both at a higher temperature indicating a stronger metal-support interaction. This also

74

verifies a higher dispersion over the tetragonal catalyst as speculated in the XRD analysis. Smaller

Ni particle size enhances its interaction with the zirconia support [123]. Hydrogen consumption

values, shown in Table 3.3, reveal that more NiO is reduced for the monoclinic catalyst in the

studied temperature range compared to the tetragonal. The theoretical hydrogen consumption

value is 896 μmol H2/gcatalyst.

Table 3.3: Hydrogen consumption values of Ni/ZrO2 (M) and Ni/ZrO2 (T)

3.2.2.4 CO2-Temperature Programmed Desorption (CO2-TPD) of calcined Ni/ZrO2 catalysts

CO2-TPD profiles of both catalysts are shown in Figure 3.12.

100 200 300 400 500 600 700

Ni/ZrO2 (M)

Ni/ZrO2 (T)

CO

2 d

eso

rbe

d (

a.u

.)

Temperature (oC)

I

Figure 3.12: CO2-TPD profiles of Ni/ZrO2 (M) and Ni/ZrO2 (T)

The tetragonal catalyst exhibits a higher basicity than the monoclinic as evidenced by a desorption

peak at 100oC verifying the presence of weak basic sites (I). Yamasaki et al. [124] studied the

influence of the zirconia phase on the catalytic activity for CO2 methanation. The tetragonal

zirconia-supported nickel catalysts also showed a greater CO2 adsorption than the monoclinic

zirconia-supported nickel catalysts. Basicity values are shown in Table 3.4.


(μmol H2/gcatalyst)

Ni/ZrO2 (M) 781

Ni/ZrO2 (T) 305

75

Table 3.4: Basicity values of Ni/ZrO2 (M) and Ni/ZrO2 (T)

Catalyst Basicity


Ni/ZrO2 (M) -

Ni/ZrO2 (T) 59

76


Catalytic test results of Ni/ZrO2 (M) and Ni/ZrO2 (T) are shown in Figure 3.13.


reaction over Ni/ZrO2 (M) and Ni/ZrO2 (T) (flow rate: 0.025 mL/min, steam/glycerol molar ratio: 9)

400 450 500 550 600 650 7000

10

20

30

40

50

60

70

80

90

100(c) Ni/ZrO

2 (M)

Ni/ZrO2 (T)

CO

2 S

ele

ctivity (

%)

Temperature (oC)

400 450 500 550 600 650 7000

10

20

30

40

50

60

70

80

90

100(d)

Ni/ZrO2 (M)

Ni/ZrO2 (T)

CO

Se

lectivity (

%)

Temperature (oC)

400 450 500 550 600 650 7000

10

20(e) Ni/ZrO

2 (M)

Ni/ZrO2 (T)

CH

4 S

ele

ctivity (

%)

Temperature (oC)

400 450 500 550 600 650 7000.0

0.5

1.0

1.5

2.0 Ni/ZrO

2 (M)

Ni/ZrO2 (T)

Hyd

rog

en

yie

ld (

mo

l H

2/m

ol g

lyce

rol)

Temperature (oC)

(a)

400 450 500 550 600 650 7000

10

20

30

40

50(b) Ni/ZrO

2 (M)

Ni/ZrO2 (T)

Gly

ce

rol co

nve

rsio

n to

ga

se

ou

s (

%)

Temperature (oC)

77

In terms of glycerol conversion to gaseous products and hydrogen yields, the monoclinic catalyst

was more active in the studied temperature range. This implies that there was more extensive C-C

bond cleavage over the monoclinic catalyst. Differing CO2, CO and CH4 selectivity trends indicate

different mechanisms occurring on both catalysts as a result of the different zirconia phases.

The tetragonal catalyst displayed selectivity trends similar to the Ni/CeO2 and Ni/Y2O3 catalysts

in the previous section. Below 500oC the water gas shift reaction was favored but beyond 500oC,

the endothermic reverse water gas shift reaction became more favored as evidenced by decreasing

CO2 and increasing CO selectivities. This could be due to the presence of the basic sites shown by

CO2-TPD profiles (Figure 3.12) which favored CO2 adsorption [125]:


CH4 selectivity over the tetragonal catalyst both increased and then decreased over the studied

temperature range indicating the occurrence of both methane producing and methane consuming

reactions. Beyond 550oC, CH4 selectivity decreased due to the dominance of the methane

decomposition reaction.


The lower catalytic activity of the tetragonal catalyst can be explained by H2-TPR analysis (Figure

3.11). Due to the increased metal-support interaction resulting from a change in the Ni particle

size, less NiO was reduced in the studied temperature range for the tetragonal catalyst. Thus, there

was less reduced NiO available on this catalyst to cleave the C-C bonds in the glycerol molecule.


The figures showing the results of the liquid product analysis are in Appendix B. Total glycerol

conversion values for the monoclinic catalyst are higher than the tetragonal catalyst and continued

to increase with increasing temperatures whereas those of the tetragonal remained somewhat

constant. Liquid product concentrations showed that there is clearly a higher formation of liquid

by-products over the tetragonal catalyst. As discussed previously, liquid by-products can be

produced by oligomerization reactions that occur mainly on acid sites [45]. Dehydration in

particular is also known to occur on acid sites [68]. Higher concentrations of hydroxyacetone (a

dehydration by-product) obtained over the tetragonal catalyst could be an indication of the

presence of more acid sites. Even at higher temperatures, the tetragonal catalyst had significant

78

quantities of hydroxyacetone and acetaldehyde. Whereas for the monoclinic catalyst,

concentrations of all liquid products decreased with increasing temperatures. This indicates that

the monoclinic catalyst was more efficient at converting oxygenated hydrocarbons (OHCs) into

CH4, CO, CO2 and H2 especially at higher temperatures. This is also verified by the higher total

glycerol conversions and glycerol conversion to gaseous products obtained for the monoclinic

catalyst.


Thermal analysis results of the used catalysts are shown in Figure 3.14.

TG results (Figure 3.14(a)) after test reveal a single weight loss step (16%) for the monoclinic

catalyst and two weight loss steps (7%) for the tetragonal catalyst. The larger quantity of coke

formed over the monoclinic catalyst is due to its higher activity in the GSR reaction. DSC profiles

(Figure 3.14(b)) showed that the monoclinic catalyst produced only filamentous coke (oxidation

T>500oC). The tetragonal catalyst produced both filamentous coke (oxidation T>500oC) as well

as encapsulating coke (T<500oC) which is easier to oxidize. The formation of the encapsulating

coke only for the tetragonal catalyst is clearly linked to the higher production of liquid by-products.

These by-products are the precursors of nonfilamentous/encapsulating coke. The presence of

filamentous coke in both catalysts is attributed to both the methane decomposition and Boudouard

reactions.

Figure 3.14: (a) TG and (b) DSC curves obtained for used Ni/ZrO2 (M) and Ni/ZrO2 (T)

100 200 300 400 500 600 700 80084

86

88

90

92

94

96

98

100

102

104(a)

Ni/ZrO2 (M)

Ni/ZrO2 (T)

Weig

ht Lo

ss (

%)

Temperature (oC)

7%

16%

100 200 300 400 500 600 700 800

(b) Ni/ZrO2 (M)

Ni/ZrO2 (T)

586oC

DS

C s

ign

al (a

.u.)

Temperature (oC)

426oC

586oC

79

3.2.6 Stability Test

To study the activity of both catalysts over longer periods of time on stream, stability tests were

conducted at 600oC using the same test conditions (WGFR 9:1, flow rate: 0.025 mL/min). Glycerol

conversions to gaseous products, hydrogen yields and CO2, CO, CH4 selectivities were monitored

and the results for the tetragonal and monoclinic catalysts are presented in Figures 3.15 and 3.16

respectively. Liquid products were collected at intervals and total glycerol conversions and liquid

product concentration results are also presented for each catalyst in these figures.

The tetragonal catalyst was deactivated during the first six hours as evidenced by a sharp drop in

glycerol conversions to gaseous products and hydrogen yields especially dominant in the first two

Figure 3.15: (a) Selectivity to C products (CO2, CH4, CO) (b) H2 yields and glycerol conversion to gaseous products (c)

Liquid product concentrations (d) Total glycerol conversions of stability test over Ni/ZrO2 (T) (flow rate: 0.025 mL/min,

steam/glycerol molar ratio: 9)

0 60 120 180 240 300 3600

10

20

30

40

50

60

70

80

90

100Ni/ZrO

2 (T)

Sele

ctivity to C

pro

du

cts

(%

)

Time (min)

CO2

CH4

CO

(a)

0 60 120 180 240 300 3600

10

20

30

40

50

60

Time (min)Gly

ce

rol co

nve

rsio

n to

ga

se

ou

s (

%)

0

1

2

3

4

5

6

7(b) Ni/ZrO

2 (T)

H2 y

ield

(mo

l H2 /m

ol o

f gly

ce

rol)

120 240 3600.00

0.05

0.10

0.15

0.20

0.25

0.30

0.35

0.40

0.45

0.50(c)

Concentr

ation (

mol/L)

Time (min)

Acetaldehyde

Acetone

Hydroxyacetone

Ethylene Glycol

Ni/ZrO2 (T)

120 240 3600

10

20

30

40

50

60

70

80

90

100(d) Ni/ZrO

2 (T)

Tota

l G

lycero

l C

onvers

ion (

%)

Time (min)

80

hours on stream. The decline continued for the remaining four hours but at a less pronounced pace.

CO2 and CO selectivities also varied in the first two hours and then remained constant. The

intervals for the liquid product collection of the tetragonal catalyst were every 2 hours. Total

glycerol conversions further verified the drop in activity discussed previously. In just the first four

hours, total conversions dropped from 68% to 35%. Liquid product concentrations over the 6 hours

kept increasing particularly hydroxyacetone demonstrating that the tetragonal catalyst was unable

to convert these by-products to gaseous products over the stability test.

On the other hand, the monoclinic catalyst was found to be relatively stable for 24 hours in terms

of glycerol conversions to gaseous products and hydrogen yields. CO2 selectivities increased


Liquid product concentrations (d) Total glycerol conversions of stability test over Ni/ZrO2 (M) (flow rate: 0.025 mL/min,


0 500 1000 15000

20

40

60

80

100(a) Ni/ZrO

2 (M)

Sele

ctivity to C

pro

du

cts

(%

)

Time (min)

CO2

CH4

CO

0 500 1000 15000

5

10

15

20

25

30

35

40

45

50Ni/ZrO

2 (M)

Time (min)Gly

cero

l con

vers

ion

to g

aseo

us (

%)

0

1

2

3

4

5

6

7(b)

H2 y

ield

(mol H

2 /mol o

f gly

cero

l)

1260 14400.00

0.05

0.10

0.15

0.20

0.25

0.30

0.35

0.40

0.45

0.50(c) Ni/ZrO

2 (M)

Co

ncentr

ation (

mol/L)

Time (min)

Acetaldehyde

Acetone

Hydroxyacetone

Ethylene Glycol

0 500 1000 15000

10

20

30

40

50

60

70

80

90

100(d) Ni/ZrO

2 (M)

Tota

l G

lycero

l C

onvers

ion (

%)

Time (min)

81

during the 24 hours while CO selectivities decreased. This could be an indication of the occurrence

of the water gas shift reaction:


And/or Boudouard reaction:


For the monoclinic catalyst, liquid products were collected in the first 3 hours of the test, left

overnight and collected at the last 3 hours. Total glycerol conversions showed that over the first

three hours of the stability test, the total glycerol conversion was 73% and after 18 hours on stream

it only slightly decreased to 61%. In the last three hours of the stability test, conversions further

decreased to 45%. As for liquid product concentrations in the last 3 hours, the concentrations of

hydroxyacetone and ethylene glycol appeared to increase whereas those of acetaldehyde and

acetone decreased. This indicates that towards the end of the 24-hour stability test, the monoclinic

catalyst began to lose its ability to convert by-products to gaseous products.

Thermal analysis results of both used catalysts after stability test are shown in Figure 3.17.

TG curves of the catalysts after stability test (Figure 3.17(a)) also revealed higher coke formation

on the monoclinic catalyst (26%) compared to the tetragonal catalyst (7%). Once again DSC

profiles (Figure 4.17(b)) showed that only one type of coke was formed on the monoclinic catalyst

Figure 3.17: a) TG and (b) DSC curves obtained for used Ni/ZrO2 (M) and Ni/ZrO2 (T) after stability test

100 200 300 400 500 600 700 80070

75

80

85

90

95

100 (a)

Ni/ZrO2 (M)

Ni/ZrO2 (T)

Weig

ht Lo

ss (

%)

Temperature (oC)

26%

7%

100 200 300 400 500 600 700 800

(b) Ni/ZrO2 (M)

Ni/ZrO2 (T)

DS

C s

ign

al (a

.u.)

Temperature (oC)

615oC

577oC

424oC

82

(filamentous) and two types were formed on the tetragonal (filamentous and encapsulating). This

indicates that the deactivation of the tetragonal catalyst was related to coke deposition. As

explained in Chapter 1, some types of coke result in a loss of catalytic whereas others do not.

Filamentous coke, although formed in the vicinity of active metal sites on the surface of the

catalyst, does not lead to catalytic deactivation. Whereas encapsulating coke interacts with surface

metal sites and leads to progressive deactivation. Therefore, deactivation depends on the nature of

the coke and not the coke quantity. Although the monoclinic catalyst had a higher coke content,

because it was of the filamentous nature which does not block active metal sites, it remained

relatively stable. Whereas the tetragonal catalyst produced both encapsulating and filamentous

coke. The encapsulating coke blocked active Ni metal sites which led to its deactivation [116].

3.2.7 Discussion

In this section, the activity of a monoclinic Ni/ZrO2 catalyst was compared to the activity of a

tetragonal Ni/ZrO2 catalyst in the GSR reaction to study the influence of the zirconia phase on

catalytic activity. X-ray diffractograms showed that the monoclinic catalyst had larger NiO

crystallites than the tetragonal. The specific surface area of the monoclinic catalyst was much

lower than the tetragonal given the low surface area of the monoclinic phase as well as the high

calcination temperature of the support since it was commercial. H2-TPR profiles revealed an

increased metal-support interaction for the tetragonal catalyst evidenced by higher temperature

reduction peaks. Nevertheless, a lower hydrogen consumption was obtained in the studied

temperature range. CO2-TPD profiles showed the higher basicity of the tetragonal catalyst. Test

results revealed different mechanisms occurring over each catalyst as a result of the crystalline

phase. The monoclinic catalyst was more active and gave higher hydrogen yields and glycerol

conversion to gaseous products. The enhanced reducibility of Ni over the monoclinic catalyst

(shown in H2-TPR results) could be responsible for the better performance. Liquid product analysis

showed higher total glycerol conversion over the monoclinic catalyst verifying its higher activity.

It also revealed a higher production of liquid by-products over the tetragonal catalyst. Thermal

analysis after test showed more coke formation over the monoclinic catalyst given its higher

activity. It also showed that the monoclinic catalyst experienced the formation of strictly

filamentous coke whereas the tetragonal catalyst experienced the formation of both non-

filamentous (encapsulating) and filamentous coke. Larger Ni particle sizes favor the formation of

filamentous coke whereas smaller Ni particle sizes favor the formation of encapsulating coke. This

83

was clearly observed in the results of this section. The higher concentration of liquid by-products

obtained over the tetragonal catalyst could also be the precursors of the encapsulating coke.

Stability tests at 600oC showed that the tetragonal catalyst deactivated during only 6 hours on

stream whereas the monoclinic catalyst remained relatively stable for 24 hours. Thermal analysis

after stability test verified that the deactivation of the tetragonal catalyst was due to the presence

of encapsulating coke which is known to block active Ni sites and leads to progressive

deactivation. Although the monoclinic catalyst produced a larger amount of coke, because it was

of a filamentous nature, it did not affect the catalytic activity or stability. Nevertheless, the decrease

of activity towards the end of the stability test for the monoclinic catalyst cannot be overlooked.

As TG-DSC results verified only the formation of filamentous coke over the monoclinic catalyst,

in this case it appeared that the filamentous coke played a role in the loss of activity towards the

end. The only plausible explanation for this is that even though filamentous coke does not block

active metal sites, if the formation of filaments is very high it could eventually block the access of

the glycerol and steam molecules to the catalyst thereby reducing the activity. Indeed, upon

termination of the stability test, the reactor was filled with a significant quantity of visible coke

(filamentous) and this could have prevented access of the reactant molecules to the pores of the

catalyst towards the end of the stability test thereby explaining the gradual loss of activity [126].

3.3 Conclusion

In this chapter, the activity of Ni-based pure metal oxides was studied in the glycerol steam

reforming reaction. The effect of the support on catalytic activity was first studied where Ni/ZrO2

was found to be the most active catalyst compared to Ni/CeO2 and Ni/Y2O3 catalysts. The effect

of the zirconia phase was studied next and the monoclinic zirconia catalyst was found to be more

active and stable compared to the tetragonal zirconia catalyst. The findings of this chapter

emphasize the importance of the choice of the support material as it determines the physico-

chemical properties of the catalyst which in turn define its performance in the glycerol steam

reforming reaction. Since the supports utilized in this chapter are quite expensive, it is important

to study the use of cheaper supports that are more attractive for industrial application. Therefore,

in the next chapter, Mg-Al based hydrotalcite derived mixed oxides will be utilized as the supports.

Their use as supports for biodiesel production has already been shown to be promising in Chapter

2.

84

Chapter 4: Steam reforming of glycerol over

Ru or Ni-based Mg-Al hydrotalcite derived

catalysts

85

This chapter focuses on the use of Mg-Al type hydrotalcite catalysts for the steam reforming of

glycerol. In the first section, ruthenium is used as the active phase. Two different preparation

methods are utilized for the synthesis of Ru-Mg-Al catalysts and their catalytic activities are

compared. The activity of the Mg-Al support alone without an active phase is also studied. In the

second section, nickel is used as the active phase and the Mg-Al support is modified with the

promoter La in different quantities. The influence of La quantity on support characteristics is

studied and the Ni impregnated catalysts are tested in the GSR. In the third section, parameter

modifications on the most efficient Ni-based catalyst are conducted. In the fourth section, the

activity of the most efficient Ni-based hydrotalcite catalyst will be compared to the most efficient

Ni-based pure metal oxide catalyst from Chapter 3. In the fifth and sixth sections, a stability test

and optimization of reaction conditions are also conducted on the most efficient Ni-based

hydrotalcite catalyst.

4.1 Role of the preparation method in catalytic activity of Ru-Mg-Al

catalysts

In this section, two Ru-Mg-Al catalysts will be synthesized using two different preparation

methods: impregnation and grafting. The physico-chemical properties of both catalysts will be

compared as well as their catalytic activities in the glycerol steam reforming (GSR) reaction.

4.1.1 Synthesis of Ru-Mg-Al catalysts

4.1.1.1 Synthesis of Ru Grafted catalyst (RuG/Mg6Al2)

The ruthenium containing or ruthenium grafted hydrotalcite was synthesized by co-precipitation

at a constant pH [127]. A solution containing the metal nitrates of the cations [Ru(NO)(NO3)3 (1.5

wt% Ru), Mg(NO3)2.6H2O and Al(NO3)3.9H2O] in the appropriate ratios such as to obtain a Ru

loading of 1 wt% was precipitated with a basic solution of Na2CO3 and NaOH. The pH was

constantly maintained between 9.5 and 10. The obtained slurry was heated at 60oC under stirring

for 1 hour. After washing, the precipitate was dried overnight at 60oC. The catalyst was then

calcined at 600oC for 4 hours (heating rate 1oC/min). The catalyst was labeled RuG/Mg6Al2 where

“G” indicates that the catalyst is grafted.

86

4.1.1.2 Synthesis of Ru Impregnated catalyst (RuI/Mg6Al2)

The ruthenium impregnated catalyst was prepared by wet impregnation of a Mg6Al2 support

prepared according to the method reported above. The support was first calcined at 700oC and then

impregnated with a solution of Ru(NO)(NO3)3 (1.5 wt% Ru) such as to obtain a metal loading of

1 wt% Ru. The obtained solution was stirred for 2 hours, slowly evaporated under vacuum in a

rotary evaporator and then dried at 60oC to ensure complete liquid removal. After impregnation,

the catalyst was calcined at 600oC for 4 hours (heating rate 1oC/min). The catalyst was labeled

RuI/Mg6Al2 where “I” indicates that the catalyst is impregnated.

To distinguish between samples before and after calcination, the abbreviation HT was used when

the catalyst is in its uncalcined form.

4.1.2 Characterization of Ru-Mg-Al catalysts

The prepared catalysts were characterized to compare their physico-chemical properties. The

various characterization techniques used, their description and conditions are presented in

Appendix B.

4.1.2.1 X-ray diffractograms (XRD) of uncalcined and calcined Ru-Mg-Al catalysts

X-ray diffractograms of the uncalcined Ru-Mg-Al catalysts are shown in Figure 4.1.

10 20 30 40 50 60 70 80

**

Inte

nsity (

a.u

.)

2(o)

*

**

* ***

* Hydrotalcite phase

*

* ** * **

RuG/Mg6Al

2 HT

RuI/Mg6Al

2 HT

Figure 4.1: XRD diffractograms of the uncalcined Ru-Mg-Al catalysts

Both diffraction patterns show the characteristic peaks of a hydrotalcite structure (JCPDS 22-0700)

confirming that the hydrotalcite structure is retained even after ruthenium incorporation [88].

87

However, diffraction peaks of the uncalcined grafted catalyst (RuG/Mg6Al2HT) are less intense

and broader than those of the impregnated catalyst (RuI/Mg6Al2HT). This indicates a decrease in

crystallinity as a result of the incorporation of Ru into the hydrotalcite structure [61]. Ru3+ ions

(0.068nm) have a larger ionic radius than Al3+ (0.054nm) leading to a distortion of the brucite

layers explaining the poorer crystallinity. This is also demonstrated by the crystallite sizes (Table

4.1) calculated at the (003) plane. The lattice parameter (a) for the grafted catalyst was also slightly

higher than that of the impregnated due to the larger ionic radius of ruthenium compared to

aluminum. Whereas, the value of lattice parameter (c) was lower for the grafted catalyst indicating

an increase in electrostatic attraction between negative and positive charges in the layers resulting

from Ru3+ introduction [57]. The formulas used to calculate these values are detailed in Appendix

B.

Table 4.1: Lattice parameters and crystallite sizes of the uncalcined Ru-Mg-Al catalysts

Catalyst 2 (o) d003(Å) c (Å) 2 (o) d110(Å) a (Å)

Crystallite

size (nm)

RuG/Mg6Al2HT 11.3270 7.8055 23.4164 60.3770 1.5319 3.0638 7.4

RuI/Mg6Al2HT 11.3250 7.8069 23.4206 60.4050 1.5312 3.0625 10.5

X-ray diffractograms of both calcined catalysts, shown in Figure 4.2, show peaks of the MgO

phase (JCPDS 45-0946). No peaks corresponding to ruthenium are observed indicating a good

dispersion of the active phase [128].

20 30 40 50 60 70 80

# MgO

RuG/Mg6Al

2

RuI/Mg6Al

2

Inte

nsity (

a.u

.)

2(o)

#

#

#

#

#

#

Figure 4.2: XRD diffractograms of the calcined Ru-Mg-Al catalysts

88

4.1.2.2 Specific surface area analysis (BET) of calcined Ru-Mg-Al catalysts

Figure 4.3 compares the specific surface areas of the calcined Ru-Mg-Al catalysts.

0

50

100

150

200

250

RuI/Mg6Al

2

Sp

ecific

su

rfa

ce

are

a (

m2/g

)

RuG/Mg6Al

2

218

112

Figure 4.3: Specific surface areas of the calcined Ru-Mg-Al catalysts

The grafted catalyst shows a larger specific surface area than the impregnated catalyst after

calcination. This could be due to pore clogging and deposition on the surface of the impregnated

catalyst as a result of the impregnation technique of the active metal [88]. Another possible

explanation could be the two-step thermal treatment (700oC of the support followed by 600oC after

impregnation) which led to a lower surface area [129].

4.1.2.3 H2-Temperature Programmed Reduction (H2-TPR) of calcined Ru-Mg-Al catalysts

H2-TPR profiles of the calcined Ru-Mg-Al catalysts are shown in Figure 4.4.

89

100 200 300 400 500 600 700

192oC

426oC

446oC

Inte

nsity (

a.u

.)

Temperature (oC)

RuG/Mg6Al

2

RuI/Mg6Al

2

213oC

Figure 4.4: H2-TPR profiles of the calcined Ru-Mg-Al catalysts

Both catalysts exhibit two main reduction peaks due to the presence of different Ru oxidic species.

The low temperature reduction peaks can be attributed to the reduction of well dispersed RuOx

species and bulk RuO2 species. The high temperature reduction peaks can be assigned to the

reduction of oxidized Ru species interacting strongly with the Mg-Al support [49]. The high

temperature peak indicates a good stabilization of oxidized ruthenium species in the catalyst

matrix. Both peaks are shifted to higher temperatures for the grafted catalyst indicating a stronger

interaction of the active phase with the support resulting from the preparation method. This

suggests a better Ru dispersion on the grafted catalyst than on the impregnated catalyst [54]. Based

on H2-TPR profiles, hydrogen consumptions were calculated and are shown in Table 4.2.

Table 4.2: Hydrogen consumption values of the calcined Ru-Mg-Al catalysts

Both catalysts exhibited higher consumption values than the stoichiometric value required for the

reduction of RuO2 to metallic Ru (200μmol H2/ gcatalyst). This could be due to the partial reduction

of oxides present in the support facilitated by the presence of ruthenium species [130]. Lower

hydrogen consumption values for the grafted catalyst indicate a lower accessibility of the active

phase compared to the impregnated. It appears that for the grafted catalyst, stabilization of the Ru


(μmol H2/ gcatalyst)

RuG/Mg6Al2 283

RuI/Mg6Al2 341

90

in the catalyst matrix led to an increased metal-support interaction but also resulted in a lower

accessibility [61].

4.1.2.4 CO2-Temperature Programmed Desorption (CO2-TPD) of calcined Ru-Mg-Al catalysts

Basic properties of both catalysts were studied using CO2-TPD. The amount of CO2 adsorbed is

related to the total amount of base sites and the CO2 desorption temperature indicates the strength

of the base site. The higher the CO2 desorption temperature, the stronger the basic site [54]. The

profiles for both catalysts are shown in Figure 4.5.

100 200 300 400 500 600 700

Inte

nsity (

a.u

)

Temperature (oC)

RuG/Mg6Al

2

RuI/Mg6Al

2

Figure 4.5: CO2-TPD profiles of calcined Ru-Mg-Al catalysts

No desorption peaks were observed at temperatures above 700oC. A broad desorption peak was

observed for both catalysts in the temperature range of 50-400oC. This peak is due to interaction

of CO2 with low and medium strength basic sites resulting from surface hydroxyl groups and

metal-oxygen pairs (Mg2+-O2-or Al3+-O2-) respectively [88]. Lower basicity values for the grafted

catalyst (Table 4.3) could be explained by the replacement of aluminum ions in the brucite layers

by ruthenium ions.

Table 4.3: Basicity values of calcined Ru-Mg-Al catalysts


Figure 4.6 (a) and (b) show the hydrogen yields and glycerol conversions to gaseous products

respectively. Selectivities of CO2, CO and CH4 are presented in Figure 4.6 (c), (d), and (e)

Catalyst Basicity


RuG/Mg6Al2 128

RuI/Mg6Al2 158

91

respectively. To confirm the role of ruthenium in the catalysts, the catalytic activity of the support

alone (Mg6Al2) was also studied and the results are presented for comparison.

Figure 4.6: (a) H2 yields (b) Glycerol conversion to gaseous products (c) CO2 (d) CO (e) CH4 selectivities of GSR reaction over

Ru-Mg-Al catalysts (flow rate: 0.075 mL/min, steam/glycerol molar ratio: 9)

400 450 500 550 600 650 7000.0

0.2

0.4

0.6

0.8

1.0

1.2

1.4

1.6

1.8

2.0

H2 Y

ield

(m

ole

/mo

le o

f g

lyce

rol)

Temperature (oC)

RuG/Mg6Al

2

RuI/Mg6Al

2

Mg6Al

2

(a)

400 450 500 550 600 650 7000

10

20

30

40

50

Gly

ce

rol co

nve

rsio

n g

ase

ou

s (

%)

Temperature (oC)

RuG/Mg6Al

2

RuI/Mg6Al

2

Mg6Al

2

(b)

400 450 500 550 600 650 7000

10

20

30

40

50

60

70

80

90

100

CO

2 S

ele

ctivity (

%)

Temperature (oC)

RuG/Mg6Al

2

RuI/Mg6Al

2

Mg6Al

2

(c)

400 450 500 550 600 650 7000

10

20

30

40

50

60

70

80

90

100C

O S

ele

ctivity (

%)

Temperature (oC)

RuG/Mg6Al

2

RuI/Mg6Al

2

Mg6Al

2

(d)

400 450 500 550 600 650 7000

10

20 RuG/Mg

6Al

2

RuI/Mg6Al

2

Mg6Al

2

CH

4 S

ele

ctivity (

%)

Temperature (oC)

(e)

92

Hydrogen yield over the Mg6Al2 support was notably very low. The metal oxides in the Mg6Al2

support can promote the main steam reforming reaction to some extent and increase glycerol

conversion to gaseous products due to their surface basicity. Nevertheless, the H2 yield obtained

is negligible. This is commonly observed for pristine metal oxides [131]. The increasing trend of

glycerol conversion to gaseous products observed for all catalysts including the support alone can

also be explained by the endothermic nature of the GSR reaction [67]. At lower temperatures, the

steam reforming reaction is limited as other reactions such as water gas shift and methanation are

more dominant. Nevertheless, at higher temperatures, the metal loaded catalysts provide higher

conversion values compared to the support alone. This is consistent with findings in literature

which report that active metal sites enhance the cleavage of C-C or C-O bonds in glycerol

molecules through dehydrogenation on active metal sites [45]. A synergistic effect between Ru

and the basic sites of the Mg-Al support could also be responsible for the improved catalytic

performance of both Ru containing catalysts [132].

Hydrogen yields over the support alone were also significantly lower than the impregnated and

grafted catalysts due to the dominance of methanation reactions at lower temperatures and the

reverse water gas shift (RWGS) reaction at higher temperatures all of which lead to the

consumption of H2 [133]. The decreasing trend for CO2 selectivity and increasing trend for CO

selectivity with increasing temperatures over the support also verifies the dominance of the

endothermic reverse water gas shift reaction [67]:

CO2 + H2 CO + H2O ΔH298 = 41 kJ/mol

Increasing CH4 selectivities further verify the dominance of methanation reactions over the support

alone. Both Ru-containing catalysts exhibited similar CO2 and CO selectivities until 600oC where

the grafted catalyst demonstrated lower CO2 selectivities and higher CO selectivities as a result of

the reverse water gas shift reaction. Methane selectivity remained relatively constant over the

impregnated catalyst while it increased for the grafted catalyst verifying the occurrence of

methanation reactions. In terms of hydrogen yield and glycerol conversion to gaseous products,

the grafted demonstrated lower values for both parameters compared to the impregnated catalyst.

The similarity of trends observed beyond 600oC for the grafted in comparison to the support alone

leads to the conclusion that beyond 600oC, the accessibility to the active phase was limited as a

result of the preparation method. Studies on the influence of active metal particle size on catalytic

93

activity in the steam reforming reaction are scarce [134]. Nevertheless, for noble metals, a

correlation between surface-specific activity and varying particle size has been observed by several

authors [135], [136]. This correlation varies depending on the active metal and support utilized.

Wang et al. [137] studied the influence of ruthenium nanoparticles supported on carbon nanotubes

for selective hydrogenolysis of glycerol to glycols and found that the glycerol conversion and

product selectivity strongly depended on the size of the Ru nanoparticles. Thus, the difference of

activity observed between the grafted and impregnated catalysts could also be attributed in part to

a difference in Ru particle size as a result of the preparation method. The double thermal treatment

of the impregnated catalyst as well as the reconstruction of the hydrotalcite structure during the

impregnation step could have influenced the Ru particle size [61].


Carbon deposition over RuI/Mg6Al2 and RuG/Mg6Al2 after the reaction was investigated by means

of TG-DSC as shown in Figure 4.7.

TG curves (Figure 4.7(a)) of the spent catalysts indicate weight losses due to the removal of carbon

[138]. RuI/Mg6Al2 demonstrated a weight loss (34%) larger than RuG/Mg6Al2 (17%). More coke

was formed on the impregnated catalyst due to its higher activity beyond 600oC. Looking at the

DSC profiles in Figure 4.7(b), both catalysts exhibit three exothermic peaks corresponding to the

different types of coke formed. Oxidation peaks (T<500oC) correspond to the oxidation of

adsorbed carbon species interacting with surface metal sites also referred to as non-filamentous

Figure 4.7: a) TG and b) DSC curves obtained for used Ru-Mg-Al catalysts

100 200 300 400 500 600 700 800

17%

We

igh

t L

oss (

%)

Temperature (oC)

RuG/Mg6Al

2

RuI/Mg6Al

2

34%

(a)

100 200 300 400 500 600 700 800

RuG/Mg6Al

2

RuI/Mg6Al

2

DS

C s

ign

al (a

.u.)

Temperature (oC)

377

447

550

323

423

540

(b)

94

coke. Higher temperature peaks (T>500oC) correspond to the oxidation of whisker-like,

filamentous coke also found in the vicinity of active metal sites [139]. The lower temperatures of

the peaks of RuG/Mg6Al2 compared to RuI/Mg6Al2 indicates that the carbon formed on the grafted

catalyst is easier to oxidize [66].

4.1.5 Stability Test

Given that the test conditions used in this study are severe enough to provoke coke deposition,

analysis of the condensate during a short-term stability test was conducted in order to gain more

insight into the difference of activity observed between both catalysts. The total glycerol

conversion was monitored over 6 hours on stream at 600oC and the results are presented in Figure

4.8 (a). Liquid product concentrations during the short term stability test are also presented in

Figure 4.8 (b).

The impregnated catalyst showed an overall higher glycerol conversion than the grafted as

expected. However, both catalysts experienced a dramatic decrease in glycerol conversion possibly

due to the accumulation of coke on the catalyst surface [31]. This trend has been observed with

other authors [66], [140]. Adsorbed carbon species, also referred to as encapsulating coke, can

block active metal sites leading to a decrease in conversion [126]. Conversions became more stable

after the first two hours as the drop became less expressed. It is important to note that the use of a

basic material (Mg-Al) does not necessarily assure the stability of the catalyst for glycerol steam

Figure 4.8: (a) Total glycerol conversions and (b) liquid product concentrations obtained during stability test over Ru-

Mg-Al catalysts (flow rate: 0.075 mL/min, steam/glycerol molar ratio: 9)

60 120 180 240 300 3600.0

0.2

0.4

0.6

0.8

1.0

1.2

1.4

1.6

1.8

2.0

Hydroxyacetone

Acetone

Acetaldehyde

C

on

ce

ntr

atio

n (

mo

l/L

)

Time (min)

RuI/Mg6Al

2

RuG/Mg6Al

2 (b)

60 120 180 240 300 36020

25

30

35

40

45

50

To

tal G

lyce

rol C

on

ve

rsio

n (

%)

Time (min)

RuG/Mg6Al

2

RuI/Mg6Al

2

(a)

95

reforming. This is because dehydration can still occur on dual sites (basic oxygen centers and

coordinatively unsaturated cations acting as Lewis acid sites) on the surface of the support which

can still provoke the formation of undesired coke precursors [82].

The complexity of the GSR process leads to the formation of a wide distribution of liquid products

[82]. Concentrations of only three major liquid products over the given time range are studied:

hydroxyacetone, acetone and acetaldehyde. The concentrations of all studied products are notably

low at 600oC as liquid product formation is more favored at lower temperatures. At higher

temperatures, products form intermediates that lead to the production of carbon oxides and

hydrogen [141]. Variations of these liquid product concentrations with time on stream verifies that

coke deposition is associated with the dehydration, dehydrogenation and hydrogenolysis reactions

that take place on catalyst active sites in GSR [142].

Acetone concentrations over both catalysts decrease with time on stream whereas hydroxyacetone

and acetaldehyde concentrations slightly increase over the studied time range. The grafted catalyst

produces lower acetone and hydroxyacetone concentrations than the impregnated catalyst. This

can be correlated with its lower catalytic activity and lower glycerol conversions at this

temperature. However, acetaldehyde concentrations over the grafted are notably higher than the

impregnated. Acetaldehyde undergoes C-C bond breakage leading to the production of CH4 and

CO [143]. This could also explain the higher CH4 and CO selectivities obtained over the grafted

catalyst above 600oC [144].

4.1.6 Discussion

XRD analysis verified the incorporation of Ru in the hydrotalcite structure of the grafted catalyst.

However, the grafted catalyst displayed a lower catalytic activity above 600oC in the GSR reaction.

This can be attributed to the lower accessibility of the active phase in the grafted catalyst as a result

of the preparation method. It appears that some of the active phase were “trapped” in the bulk of

the solid and thus became less accessible. This was verified by H2-TPR analysis which revealed

higher metal-support interactions in the grafted evidenced by higher temperature reduction peaks

but overall lower hydrogen consumptions. Both grafted and impregnated catalyst exhibited

deactivation during a short-term stability test (6 hours) as a result of carbon formation associated

with liquid product formation. Differences in catalytic activity, distribution of liquid by-products,

and coke formation between both catalysts are a strong indication of differing Ru particle sizes

96

also resulting from the different preparation method. This section also demonstrated how the

addition of the active phase Ru to Mg-Al supports significantly improved their catalytic activity

in the glycerol steam reforming reaction as a result of a possible synergistic effect. The results of

this section have been published in the International Journal of Hydrogen Energy and the article is

presented at the end of this manuscript.

97

4.2 Influence of support modification on the catalytic activity of Ni-

Mg-Al catalysts

In this section, nickel will be used as the active phase. The Mg6Al2 will still be used as the support

material but it will also be modified with La in different quantities for comparison. The La will

also be co-precipitated during the synthesis of the supports. The influence of La addition and La

quantity on the physico-chemical properties and catalytic activities will be studied. Based on the

results of the previous section, the preparation method of impregnation will be used as the means

of adding the active phase (Ni) to the supports.

4.2.1 Synthesis of supports

Supports were synthesized by the precipitation of an aqueous solution (250 mL) containing the

appropriate dissolved amounts of magnesium (II) nitrate hexahydrate Mg(NO3)2.6H2O (Sigma-

Aldrich, 98% purity), aluminum (III) nitrate nonahydrate Al(NO3)3.9H2O (Sigma-Aldrich, 98%

purity) and lanthanum (III) nitrate hexahydrate La(NO3)3.6H2O (Sigma-Aldrich, 99.995%). This

aqueous solution was precipitated with a 1M solution of sodium carbonate Na2CO3 (Uni-Chem,

99.8% purity) and 2M solution of sodium hydroxide NaOH (SureChem Products, 96% purity) at

a temperature of 60oC and a pH between 9.5 and 10 while stirring. The metal nitrates solution was

added drop by drop from a buret to a beaker containing 200 mL of distilled water and the pH was

adjusted to the range of 9.5-10 using the basic solution of Na2CO3 and NaOH. The obtained

solution was stirred for 1 hour at 60oC. It was then immediately placed for 18 hours in an oven at

60oC for the hydrotalcite phase to crystallize. The mixture was then filtered and continuously

washed with hot, deionized water (60oC) until a neutral pH of the filtrate was obtained. This

washing step was carried out to remove all soluble ions such as nitrates and sodium ions. The solid

was placed in an oven at 60oC for 48 hours then grinded to a powder form.

The solids prepared were named: Mg6Al2HT, Mg6Al1.8La0.2HT, Mg6Al1.6La0.4HT and

Mg6Al1.2La0.8HT where the subscripted numbers represent the relative amounts of the metals and

the abbreviation HT signifies that the solids were uncalcined (in hydrotalcite form).

These solids were then calcined for 4 hours at a temperature of 700oC under a flow of dry air. The

heating rate of the calcination was 1oC/min. The solids calcined at 700oC were named: Mg6Al2,

Mg6Al1.8La0.2, Mg6Al1.6La0.4 and Mg6Al1.2La0.8. The temperatures of calcination were selected

98

following thermo-gravimetric analysis. Additionally, since the upper temperature limit of all the

catalytic tests is 700oC, it is better to thermally stabilize all the supports at this temperature.

4.2.2 Characterization of supports

Supports in both their uncalcined and calcined forms were characterized by various techniques to

verify the presence of the hydrotalcite structure prior to calcination as well as its destruction and

formation of mixed oxides after calcination. These physico-chemical techniques are detailed in

Appendix B. The effect of lanthanum content on support characteristics was also studied.

4.2.2.1 Thermal analysis (TG-DSC) of uncalcined supports

Figure 4.9 (a) and (b) represent the TG (thermogravimetry) and DSC (differential scanning

calorimetry) curves obtained during the calcination of the hydrotalcite supports from room

temperature up to 800°C.

100 200 300 400 500 600 700 800

III

III

III

II

I

(a)

Mg6Al

1.2La

0.8HT

Mg6Al

1.6La

0.4HT

Mg6Al

1.8La

0.2HT

Mg6Al

2HT

We

igh

t L

oss (

%)

Temperature (oC)

I

I

I

II

II

II

100 200 300 400 500 600 700 800

4

4

4

(b)

DS

C s

ign

al (a

.u.)

Temperature (oC)

Mg6Al

1.2La

0.8HT

Mg6Al

1.6La

0.4HT

Mg6Al

1.8La

0.2HT

Mg6Al

2HT

3

21

3

21

3

21

3

21

3

21

Figure 4.9: a) TG and b) DSC curves obtained during thermal treatment of the uncalcined Mg6Al2, Mg6Al1.8La0.2,

Mg6Al1.6La0.4 and Mg6Al1.2La0.8 supports

Destruction of the hydrotalcite structure is characterized by two main weight losses (I and II). The

first weight loss which occurs at a temperature lower than 250oC corresponds to the loss of

physisorbed water and the elimination of interlayer water. This is associated with the first

endothermic peak (1) observed in the DSC curves. The second weight loss which occurs between

300-450oC corresponds to the dehydroxylation of the brucite layer and loss of interlayer anions

(carbonates). It is during this step that the hydrotalcite structure completely collapses and the oxide

phases are formed. The dehydroxylation and loss of interlayer anions are associated with the

99

second (2) and third (3) endothermic peaks respectively [145]. It is often difficult to differentiate

between these two peaks as they occur at very close temperatures. In addition, a fourth endothermic

peak (peak 4) is observed for the supports containing lanthanum. This peak corresponds to the

decomposition of lanthanum carbonates (later observed in XRD) and is accompanied by a very

small loss of mass [146].

The thermal stability of the hydrotalcite depends on the nature of the cations present in the brucite

layers [72]. As the quantity of La in the support composition increased, the temperature of peak

(1) decreased indicating a decrease in the thermal stability. This can be attributed to the

incorporation of La3+ ions in the sheets which leads to a deformation in the octahedral structure

leading to weaker electrostatic forces between the hydroxide layers and interlayer [54]. Another

reason could be that the presence of uncoordinated lanthanum species affected the co-precipitation

of the metallic salts during the synthesis process making the structure less stable [54].

It is worth noting that at 700°C, the hydrotalcite structure of all supports was completely destroyed

to form stable oxides and no further weight losses were observed. Hence, this temperature was

chosen for the calcination of the synthesized supports.

The destruction or collapse of the hydrotalcite structure leads to the formation of mixed metal

oxides and results in weight loss due to the departure of CO2 and H2O. This destruction can be

represented by the following equations for each support:

Mg6Al2(OH)16CO3.4H2O 6MgO + Al2O3 + CO2 + 12H2O

Mg6Al1.8La0.2(OH)16CO3.4H2O 6MgO + 9/10Al2O3 + 1/10La2O3 + CO2 + 12H2O



The theoretical weight losses were calculated from these equations and compared to obtained

experimental values as shown in Table 4.4. For each support, both theoretical and experimental

values were relatively close.

100

Table 4.4: Theoretical and experimental weight losses of the uncalcined Mg6Al2, Mg6Al1.8La0.2, Mg6Al1.6La0.4 and

Mg6Al1.2La0.8 supports

4.2.2.2 X-ray diffractograms (XRD) of uncalcined and calcined supports

The X-ray diffractograms (XRD) of the uncalcined supports are presented in Figure 4.10.

10 20 30 40 50 60 70 80

• •

••••

•••••••

•

•

•

•

•

•

•••••

•

•

•

•

•

••

•

•••••*

Mg6Al

1.2La

0.8HT

Mg6Al

1.6La

0.4HT

Mg6Al

1.8La

0.2HT

Inte

nsity (

a.u

.)

2(o)

Mg6Al

2HT

*

**

* ** *

** * **

*

** * * * * *

* * * * * * *

•

••• •

• • ••

* Hydrotalcite phase LaCO3OH

La2(CO

3)

2(OH)

2

LaAl(OH)2(CO

3)

2

Figure 4.10: XRD diffractograms of the uncalcined Mg6Al2, Mg6Al1.8La0.2, Mg6Al1.6La0.4 and Mg6Al1.2La0.8 supports

These diffractograms reveal a typical hydrotalcite structure (*) for all prepared supports in

compliance with the JCPDS reference number 22-0700. The small peaks evident in all the

lanthanum containing supports are attributed to the LaCO3OH (JCPDS 49-0981), La2(CO3)2(OH)2

(JCPDS 70-1774) and LaAl(OH)2(CO3)2 (JCPDS 52-1059) phases [54]. The strong anionic

character of lanthanum favors the formation of these carbonate and oxyhydroxide species during

the coprecipitation stage [54]. Different intensities are observed for the hydrotalcite peaks at 2 =

11.3o, 22.9o and 34.6o indicating a different stacking of the hydrotalcite layers in each support

[56],[57]. In fact, as lanthanum content increased, these hydrotalcite peaks become lower in

intensity indicating a lower crystallization of the hydrotalcite phase. This poorer crystallization

Catalyst Theoretical Value (%) Experimental Value (%)

Mg6Al2HT 43.1 44.7

Mg6Al1.8La0.2HT 41.5 42.2

Mg6Al1.6La0.4HT 40.1 41.2

Mg6Al1.2La0.8HT 37.5 39.1

101

could be due to the presence of the carbonate and oxyhydroxide species. Another reason lies in the

difference in ionic radii and electronegativities of the metal ions found in the supports as

demonstrated in Table 4.5: [54]

Table 4.5: Comparison of ionic radii and element electronegativities

Ion Radius

(nm)

Electronegativity

of the element

Mg2+ 0.072 1.31

Al3+ 0.054 1.61

La3+ 0.116 1.10

Due to the relatively large ionic radius of La3+ and the low electronegativity of lanthanum, it is

difficult for it to be incorporated into the brucite-like layers of the hydrotalcite. Therefore, even if

it were to be incorporated, this would lead to a distortion of the layers as a result of their radii and

electronegativities. This also helps explain the poorer crystallization of the lanthanum containing

catalysts [54].

Lattice parameters and crystallite sizes obtained for the uncalcined supports are shown in Table

4.6. The formulas used to calculate these values are detailed in Appendix B.

Table 4.6: Lattice parameters and crystallite sizes of the uncalcined Mg6Al2, Mg6Al1.8La0.2, Mg6Al1.6La0.4 and

Mg6Al1.2La0.8 supports

Catalyst 2 (o) d003(Å) c (Å) 2(o) d110 (Å) a (Å) Crystallite

size (nm)

Mg6Al2HT 11.452 7.721 23.163 60.500 1.529 3.058 8.931

Mg6Al1.8La0.2HT 11.319 7.811 23.433 60.329 1.533 3.066 8.419

Mg6Al1.6La0.4HT 11.241 7.865 23.595 60.322 1.534 3.068 7.308

Mg6Al1.2La0.8HT 10.863 8.138 24.414 59.756 1.546 3.092 5.313

With the increase of lanthanum content, an increase in the parameters “a” and “c” is noted. La3+

ions have relatively larger ionic radii compared to Al3+ as shown in Table 4.5 previously. Thus,

their insertion into the host brucite-like layers increases the distance between metal ions, therefore

parameter “a” increases. La3+ ions also have a lower charge density compared to Al3+, thus their

insertion also leads to a decrease in the charge density of the host layers and this in turn leads to a

decrease in the electrostatic forces between the layers and interlayers thereby explaining the

102

increase in parameter “c” [54]. Additionally, as the lanthanum content increased crystallite sizes

became smaller [147].

Figure 4.11 shows the XRD diffractograms recorded for the calcined supports.

20 30 40 50 60 70 80

# ##

Inte

nsity (

a.u

.)

2(o)

Mg6Al

2

Mg6Al

1.8La

0.2

Mg6Al

1.6La

0.4

Mg6Al

1.2La

0.8

#

#

##

##

# #

# #

# #

# #

La2O

2CO

3# MgO

Figure 4.11: XRD diffractograms of the calcined Mg6Al2, Mg6Al1.8La0.2, Mg6Al1.6La0.4 and Mg6Al1.2La0.8 supports

The characteristic hydrotalcite diffraction peaks that were evident in the diffractograms of the

uncalcined supports disappeared after calcination. This is in agreement with the thermal analysis

(Figure 4.9) where total destruction of the hydrotalcite structure was observed at this temperature.

The new phases observed correspond to the mixture of oxides formed after calcination [12]. These

phases include the periclase structure MgO (JCPDS 43-1022) for all supports and La2O2CO3

(JCPDS 48-1113) for the lanthanum containing supports. The La2O2CO3 phase results from

atmospheric CO2 adsorption, after calcination, on the surface of basic La2O3. The formation of this

phase is expected given the strong basicity of La2O3 as well as the “unprotected” location of the

lanthanum species that are outside the layers of the hydrotalcite [148]. The absence of peaks

corresponding to AlOx phases for all catalysts indicates that Al3+ ions are well dispersed in the

MgO lattice [149].

103

4.2.2.3 Specific surface area analysis (BET) of uncalcined and calcined supports

Figure 4.12 shows the specific surface areas of all supports recorded before and after calcination.

The surface area of calcined La2O3 is also included.

90

105

56

21

3

135

179

136

105

0

50

100

150

200

La 2O 3

Surf

ace a

rea (

m2/g

)

Uncalcined

Calcined

Mg 6Al 2

Mg 6Al 1.8

La 0.2

Mg 6Al 1.6

La 0.4

Mg 6Al 1.2

La 0.8

Figure 4.12: Specific surface areas of the uncalcined and calcined Mg6Al2, Mg6Al1.8La0.2, Mg6Al1.6La0.4, Mg6Al1.2La0.8

supports and calcined La2O3

The specific surface areas of all supports increased after calcination at 700oC. This increase

indicates the transformation of the crystalline hydrotalcite structure into metal oxides. Crystalline

phases in general have lower surface areas since there are fewer defects in their structures, thus

when these crystalline phases are destroyed by thermal treatment to form mixed metal oxides, the

specific surface area of the solid increases. Furthermore, thermal decomposition leads to the escape

of CO2 from the interlayer and the removal of water molecules from the structure. This results in

the creation of more pores and craters on the surface of the material thereby explaining the higher

surface area obtained after calcination [150]. Before calcination, as lanthanum content increased

in the catalyst composition, the surface area became lower. This difference could be explained by

the blocking and clogging of pores due to the presence of the formed species discussed in the XRD

section. After calcination, surface areas decreased with increasing lanthanum content due to the

low surface area of La2O3 (3 m2/g) [151].

104

4.2.2.4 Fourier Transform Infrared Spectroscopy (FTIR) study of uncalcined and calcined

supports

Figure 4.13 shows the FTIR spectra for the uncalcined supports.

4000 3600 3200 2800 2400 2000 1600 1200 800 400

15001370 1100

1650

3000

Mg6Al

1.2La

0.8HT

Mg6Al

1.8La

0.2HT

Tra

nsm

itta

nce (

a.u

.)

Wavenumber (cm-1)

Mg6Al

2HT

Mg6Al

1.6La

0.4HT

3550

Figure 4.13: FTIR spectra of the uncalcined Mg6Al2, Mg6Al1.8La0.2, Mg6Al1.6La0.4 and Mg6Al1.2La0.8 supports

These spectra reveal several bands that are characteristic of hydrotalcites. The band observed at

3550 cm-1 is attributed to the stretching vibration of OH (from structural hydroxyl groups, water

molecules in the interlayer zone and physisorbed water) [73],[152]. A shoulder due to hydrogen

bonding between H2O and CO32- in the interlayer is observed at 3000 cm-1 [73],[72]. The band at

1650 cm-1 is due to the H2O bending vibration [73],[72],[152]. At 1500 cm-1 the band is attributed

to vibration due to the symmetry reduction of CO32- ions as a result of the interaction of metal

cations with one of their oxygen atoms [152]. Bands at 1370 cm-1 and 850 cm-1 are observed for

CO32- due to its asymmetric stretching v3 mode and out of plane bending v2 mode respectively

[73],[152]. OH deformations are indicated by the band at 1100 cm-1 [152]. It is important to note

that bands observed lower than 800 cm-1 are due to M-O interactions (M= Al, Mg, La). However

it is difficult to strictly give an assignment for each band since there is an overlapping of the bands

due to the presence of more than one metal [152],[153].

105

The FTIR spectra of the calcined supports are presented in Figure 4.14.

4000 3600 3200 2800 2400 2000 1600 1200 800 400

3610

Tra

nsm

itta

nce (

a.u

.)

Wavenumber (cm-1)

Mg6Al

1.2La

0.8

Mg6Al

1.8La

0.2

Mg6Al

2

Mg6Al

1.6La

0.4

3550

1650

1400

Figure 4.14: FTIR spectra of the calcined Mg6Al2, Mg6Al1.8La0.2, Mg6Al1.6La0.4 and Mg6Al1.2La0.8 supports

The destruction of the hydrotalcite structure (characterized by the removal of CO2 and H2O during

decomposition) is evident in Figure 4.14 compared to Figure 4.13 due to several factors:

Weaker IR bands observed for: OH’s in the 3550 cm-1 region, H2O at 1650 cm-1, CO32- at

1370 cm-1

Disappearance of several bands: Shoulder due to interaction between H2O-CO32- at 3000

cm-1 [72]

Clear differences in the 1000-400 cm-1 region due to the transformation of the layered

structure to mixed oxides [152]

Water and carbonate bands are still observed thereby indicating that there are still carbonates and

water molecules present in the structure. It is possible that gaseous CO2 from the ambient air

became readsorbed onto the basic sites of the metal oxides [15]. Water molecules could have also

physically adsorbed onto the metal oxides explaining the presence of water bands [146]. The

intense sharp band at 3605 cm-1 for the lanthanum containing supports corresponds to the

stretching and bending vibrations of the O-H in La(OH)3 [154]. More intense bands at 1370 cm-1

for the lanthanum containing supports are an indication that the addition of this promoter led to

the formation of more basic sites which led to more CO2 adsorption.

106

4.2.2.5 Discussion

An Mg-Al and La modified Mg-Al supports were synthesized by co-precipitation of the metals at

a constant pH. Physico-chemical properties of the supports alone were studied both before and

after calcination to verify the presence of the hydrotalcite structure and ensure its destruction.

Thermal analysis demonstrated that a calcination temperature of 700oC was sufficient for the

destruction of the hydrotalcite phase. X-ray diffractograms verified the presence of hydrotalcite

peaks before calcination for all catalysts and verified the destruction of the hydrotalcite phase and

appearance of new phases after calcination at 700oC. FTIR spectra also confirmed the presence of

the hydrotalcite structure and its destruction following thermal treatment.

The influence of the promoter quantity (La) on support characteristics was also studied. Prior to

calcination, the addition of lanthanum resulted in a poorer crystallinity of the hydrotalcite phase

as lanthanum carbonate and oxyhydroxide species were formed. After calcination, the La2O2CO3

phase was observed on all La containing supports resulting from the adsorption of atmospheric

CO2 to basic La2O3 phase. With the increase in lanthanum content, specific surface areas of the

supports decreased.

4.2.3 Synthesis of Ni-based hydrotalcite derived mixed oxides

Following the preparation of the supports, the Ni-based hydrotalcite derived mixed oxides were

synthesized by wet-impregnation of the nickel precursor on the calcined supports.

An adequate amount of nickel nitrate Ni(NO3)2.6H2O (Sigma-Aldrich, 98% purity) was dissolved

in 50 mL of deionized water in order to obtain 5 wt% of Ni in the final solid. This solution is then

added to 1 g of the support and mixed during two hours. The excess of water is slowly eliminated

in a rotary evaporator (60°C, under vacuum). The obtained residue is kept in an oven at 60°C

during 24 h.

The dried solids are then thermally stabilized by calcination at 600°C (1°C/min) under an air flow

for four hours. Just like in Chapter 3, the calcination temperature of 600oC was chosen based on

literature study. The obtained calcined catalysts were named Ni/Mg6Al2, Ni/Mg6Al1.8La0.2,

Ni/Mg6Al1.6La0.4 and Ni/Mg6Al1.2La0.8.

107

4.2.4 Characterization of Ni-impregnated supports

4.2.4.1 X-ray diffractograms (XRD) of calcined Ni-impregnated supports

Figure 4.15 shows the X-ray diffractograms of the calcined Ni/Mg6Al2, Ni/Mg6Al1.8La0.2,

Ni/Mg6Al1.6La0.4 and Ni/Mg6Al1.2La0.8 solids.

20 30 40 50 60 70 80

*Ni/Mg

6Al

1.2La

0.8

Ni/Mg6Al

1.8La

0.2

Ni/Mg6Al

1.6La

0.4

Ni/Mg6Al

2

Inte

nsity (

a.u

.)

2(o)

#

##

#

#

*

# ##

##

*

#

#

#

#

##

La2O

2CO

3NiO# MgO* Hydrotalcite phase

Figure 4.15: XRD diffractograms of the calcined Ni/Mg6Al2, Ni/Mg6Al1.8La0.2, Ni/Mg6Al1.6La0.4 and Ni/Mg6Al1.2La0.8 solids

The NiO (JCPDS 44-1159) peaks for all catalysts are superimposed over the MgO peaks and thus

cannot be clearly differentiated [155]. This makes it difficult to study NiO sizes. La2O2CO3 still

appears for all the promoted catalysts. As the lanthanum content increases in the catalyst

composition, peaks corresponding to the La2O2CO3 phase become more intense indicating an

increased crystallinity of this phase. On the other hand, with increasing lanthanum content the

periclase peaks (# and x) have lower peak intensities. This reveals that the crystalline structures of

the NiO and MgO periclases are damaged to a certain degree by the incorporation of lanthanum

even after calcination. The MgO phase in particular is known to be a defective structure as it has

cationic vacancies generated by the introduction of M3+ in the octahedral sites [148].

It is important to note that for the Ni/Mg6Al1.8La0.2, Ni/Mg6Al1.6La0.4 and Ni/Mg6Al1.2La0.8

catalysts the hydrotalcite peak at 22o was still observed. This indicates that even after calcination

at 600oC, some of the hydrotalcite structure was still retained as a result of the reconstruction

during the impregnation step. Studies have shown that lanthanum plays a significant role in the

adsorption mechanism involved during reconstruction [156].

108

4.2.4.2 Specific surface area analysis (BET) of Ni-impregnated supports

Figure 4.16 displays the surface areas of the calcined supports alone compared to the surface areas

of the Ni-impregnated solids after calcination.

135

64

179

54

136

25

105

13

0

50

100

150

200 Support alone

Surf

ace a

rea (

m2/g

)

With Ni

Mg 6Al 2

Mg 6Al 1.8

La 0.2

Mg 6Al 1.6

La 0.4

Mg 6Al 1.2

La 0.8

Figure 4.16: Comparison of specific surface areas of calcined supports (Mg6Al2, Mg6Al1.6La0.4 and Mg6Al1.6Ce0.4) and Ni-

impregnated solids after calcination (Ni/Mg6Al2, Ni/Mg6Al1.6La0.4 and Ni/Mg6Al1.6Ce0.4)

Impregnation of the supports with Ni clearly led to a decrease in the specific surface areas for all

catalysts. Agglomerates of the nickel oxides formed probably blocked the pores on the support

surface thereby decreasing the specific surface area [157].

109

4.2.4.3 H2-Temperature Programmed Reduction (H2-TPR) of Ni-impregnated supports

H2-TPR profiles of the catalysts up to 700oC are shown in Figure 4.17.

100 200 300 400 500 600 700

570oC

575oC

440oC

290oC

340oC

310oC

300oC

520oC

450oC

Ni/Mg6Al

2

Inte

nsity (

a.u

.)

Temperature (oC)

Ni/Mg6Al

1.8La

0.2

Ni/Mg6Al

1.6La

0.4

Ni/Mg6Al

1.2La

0.8

470oC

Figure 4.17: H2-TPR profiles of calcined Ni/Mg6Al2, Ni/Mg6Al1.8La0.2, Ni/Mg6Al1.6La0.4 and Ni/Mg6Al1.2La0.8 solids

Since oxides of Mg, Al and La cannot be reduced below a temperature of 900oC, the peaks

observed are mainly attributed to the reduction of Ni species [158],[159]. The Ni/Mg6Al2 profile

exhibits 2 main reduction peaks. The low temperature peak at 290oC corresponds to reduction of

free NiO. The peak at 440oC is attributed to the reduction of NiO weakly interacting with the

support. A higher temperature reduction peak (not shown as it is beyond our reduction temperature

range) corresponds to the reduction of Ni2+ in thermally stable phases (the Mg(NiAl)O periclases).

These phases have high reduction temperatures due to the electron transfer between NiO and MgO

which results in strong interactions and therefore a decreased reducibility of NiO [160].

The presence of La results in a clear shift in the temperature of the reduction peaks indicating an

influence on the interaction between Ni and the support. For Ni/Mg6Al1.8La0.2, the presence of a

small quantity of La led to an increase in the reduction temperatures indicating a stronger

interaction between Ni and the support. Looking closely at the profile of Ni/Mg6Al1.8La0.2, it is

evident that there is a high temperature reduction peak outside the studied temperature range

corresponding to the reduction of NiO interacting with the support. This could also be an indication

of increased Ni dispersion [54]. As the quantity of lanthanum increases in the catalyst composition,

the temperature of this peak in Ni/Mg6Al1.6La0.4 and Ni/Mg6Al1.2La0.8 becomes lower. This is an

indication that larger quantities of La3+ can weaken interactions between NiO and the support

110

[159]. This could correspond to a decreased dispersion of the nickel species with excess lanthanum

content as speculated by some studies [54].

Experimental hydrogen consumptions for all the catalysts are shown in Table 4.7.

Table 4.7: Hydrogen consumption values of calcined Ni/Mg6Al2, Ni/Mg6Al1.8La0.2, Ni/Mg6Al1.6La0.4 and Ni/Mg6Al1.2La0.8

solids

All experimental hydrogen consumption values were lower than the theoretical hydrogen

consumption (896 µmol/g) indicating that not all Ni oxide species were reduced in the considered

temperature range. These values also verify that with the addition of a small amount of La in the

support, a higher metal-support reaction resulted in less reducible Ni oxide species in the studied

temperature range. However, larger quantities of La in the support weakened metal-support

interactions, resulting in more reducible Ni oxide species in the studied temperature range as

evidenced by higher hydrogen consumption values.

Therefore, the quantity of lanthanum influences the reducibility of Ni species as well as its

dispersion.



Ni/Mg6Al1.2La0.8 526


Ni/Mg6Al1.8La0.2 70

Ni/Mg6Al2 300

111

4.2.4.4 CO2-Temperature Programmed Desorption (CO2-TPD) of Ni-impregnated supports

Basic properties of the catalysts were studied using CO2 as an acid probe molecule, and the

obtained CO2-TPD profiles are shown in Figure 4.18.

100 200 300 400 500 600 700

Inte

nsity (

a.u

.)

Temperature (oC)

Ni/Mg6Al

1.2La

0.8

Ni/Mg6Al

1.6La

0.4

Ni/Mg6Al

1.8La

0.2

Ni/Mg6Al

2I II

III

Figure 4.18: CO2-TPD profiles of calcined Ni/Mg6Al2, Ni/Mg6Al1.8La0.2, Ni/Mg6Al1.6La0.4 and Ni/Mg6Al1.2La0.8 solids

Three desorption peaks were observed for all catalysts: below 200oC (weak basic sites), between

200-400oC (medium strength basic sites) and above 400oC (strong basic sites). Weak sites are

associated with CO2 interacting with surface hydroxyl groups (OH), medium strength sites with

Lewis acid-base pairings namely Mg2+- O2- and Al3+- O2- and strong sites with isolated/surface O2-

anions [49].

The obtained profiles revealed that as lanthanum content increased, the proportion of the weak

basic sites (I) decreased while the proportions of the medium strength (II) and strong basic sites

(III) increased. This shows that the addition of lanthanum increases Lewis acid-base pairings and

surface O2- anions resulting in the higher contribution of medium strength and strong basic sites

[161]. The proportion of strong basic sites become notably larger with increasing La content due

to cationic vacancies induced by La3+ incorporation in the mixed oxide lattice as discussed in the

XRD section.

The basicity values shown in Table 4.8, are those corresponding only to the contribution of weak

(I) and medium strength (II) basic sites, as the strong sites (III) are partially outside our studied

temperature range.

112

Table 4.8: Basicity values of calcined Ni/Mg6Al2, Ni/Mg6Al1.8La0.2, Ni/Mg6Al1.6La0.4 and Ni/Mg6Al1.2La0.8 solids

Ni/Mg6Al1.8La0.2 showed a larger basicity value compared to Ni/Mg6Al2. As lanthanum content

increased, this value became lower for Ni/Mg6Al1.6La0.4 and Ni/Mg6Al1.2La0.8 as the contribution

of weak and medium sites decreased whereas as that of strong sites (not quantified) increased.

These results verify that lanthanum enhances the basicity of the catalyst. They can also be

correlated to the FTIR spectra of the supports alone (Figure 4.14) where more intense bands at

1370 cm-1 were observed for the lanthanum containing supports due to an improved basicity.

Catalyst Basicity (μmol CO2/ gcatalyst)




Ni/Mg6Al2 358

113

4.2.5 Steam reforming of glycerol over Ni/Mg6Al2, Ni/Mg6Al1.8La0.2, Ni/Mg6Al1.6La0.4

and Ni/Mg6Al1.2La0.8


selectivities of the Ni/Mg6Al2, Ni/Mg6Al1.8La0.2, Ni/Mg6Al1.6La0.4 and Ni/Mg6Al1.2La0.8 catalysts.

400 450 500 550 600 650 7000.0

0.2

0.4

0.6

0.8

1.0

1.2

1.4

1.6

1.8

2.0

2.2 Ni/Mg

6Al

2

Ni/Mg6Al

1.8La

0.2

Ni/Mg6Al

1.6La

0.4

Ni/Mg6Al

1.2La

0.8

Hydro

gen y

ield

(m

ol H

2/m

ol gly

cero

l)

Temperature (oC)

(a)

400 450 500 550 600 650 7000

10

20

30

40

50 Ni/Mg

6Al

2

Ni/Mg6Al

1.8La

0.2

Ni/Mg6Al

1.6La

0.4

Ni/Mg6Al

1.2La

0.8

Gly

cero

l C

onvers

ion G

aseous (

%)

Temperature (oC)

(b)

400 450 500 550 600 650 7000

10

20

30

40

50

60

70

80

90

100

CO

2 S

ele

ctivity (

%)

Temperature (oC)

Ni/Mg6Al

2

Ni/Mg6Al

1.8La

0.2

Ni/Mg6Al

1.6La

0.4

Ni/Mg6Al

1.2La

0.8

(c)

400 450 500 550 600 650 7000

10

20

30

40

50

60

70

80

90

100(d)

Ni/Mg6Al

2

Ni/Mg6Al

1.8La

0.2

Ni/Mg6Al

1.6La

0.4

Ni/Mg6Al

1.2La

0.8

CO

Sele

ctivity (

%)

Temperature (oC)

400 450 500 550 600 650 7000

10

20(e) Ni/Mg

6Al

2

Ni/Mg6Al

1.8La

0.2

Ni/Mg6Al

1.6La

0.4

Ni/Mg6Al

1.2La

0.8

CH

4 S

ele

ctivity (

%)

Temperature (oC)


reaction over Ni/Mg6Al2, Ni/Mg6Al1.8La0.2, Ni/Mg6Al1.6La0.4 and Ni/Mg6Al1.2La0.8 (flow rate: 0.025 mL/min,


114

In terms of hydrogen yield and glycerol conversion to gaseous products, a clear trend is observed:

Ni/Mg6Al1.6La0.4 > Ni/Mg6Al1.2La0.8 > Ni/Mg6Al2 > Ni/Mg6Al1.8La0.2. The increase in glycerol

conversion to gaseous products can be directly correlated to the increase in hydrogen yield. It is

important to note that for the lanthanum containing catalysts, although the La2O2CO3 phase was

observed in the XRD diffractograms and not the La2O3 phase, upon reduction prior to the catalytic

test (detailed in Appendix B) the following reaction restored the La2O3 phase [159]:

La2O2CO3 + H2 La2O3 + H2O + CO

Or the La2O2CO3 decomposed to La2O3 due to the high temperature involved during the reduction

step:

La2O2CO3 La2O3 + CO2

La2O3 is a hydrophilic oxide and therefore improves the activation of adsorbed H2O [162]. This

helps favor the water-gas shift reaction particularly at lower temperatures:


This explains why at lower temperatures, all the lanthanum containing catalysts exhibited higher

CO2 and lower CO selectivities than the Ni/Mg6Al2. Beyond 550oC, favoring of the endothermic

reverse water gas shift results in a decrease in CO2 selectivities and increase in CO selectivities for

all catalysts. The basic sites observed for all catalysts in CO2-TPD profiles (Figure 4.18) further

favor CO2 adsorption:


Regarding CH4 selectivities, at lower temperatures methanation reactions appear to be more

favored over Ni/Mg6Al1.6La0.4 and Ni/Mg6Al1.2La0.8. However, beyond 550oC, Ni/Mg6Al1.6La0.4

was the only catalyst that demonstrated the capacity to reform this produced CH4 into H2 and CO

via the methane steam reforming reaction:

CH4 + H2O CO + 3H2 ΔH298K = 206 kJ/mol

The Ni/Mg6Al1.8La0.2 was the least active catalyst and displayed an activity even lower than

Ni/Mg6Al2. This can be explained by H2-TPR analysis which showed that even though the

Ni/Mg6Al1.8La0.2 had an increased metal-support interaction and higher dispersion compared to all

115

the other catalysts, the amount of reducible NiO in the studied temperature range was very low

(hydrogen consumption value of 70), even lower than the Ni/Mg6Al2 (hydrogen consumption value

of 300). Thus, there was significantly less reduced NiO available on this catalyst to cleave the C-

C bonds in the glycerol molecule explaining the lower activity. Ni/Mg6Al1.2La0.8 was less active

than Ni/Mg6Al1.6La0.4 despite having a higher hydrogen consumption value. This could be due to

the lower metal-support interaction and decreased dispersion of Ni resulting from excess La3+ in

the catalyst composition. Therefore, it appears from these results that the Ni/Mg6Al1.6La0.4

contained the optimum lanthanum content providing both a good metal-support interaction as well

as a good Ni dispersion.


The total glycerol conversions and the concentrations of some obtained liquid by-products are

shown in Appendix B. All four catalysts showed increasing values in total glycerol conversion

until 550oC. Beyond 550oC, the Ni/Mg6Al1.6La0.4 exhibited a superior performance whereby its

total glycerol conversion kept increasing with temperature whereas that of the others either

remained stable or decreased. For all catalysts, larger quantities of acetaldehyde and

hydroxyacetone (both dehydration products) were produced compared to acetone and ethylene

glycol at all temperatures. Although hydrotalcites are mostly known for their basic sites, it should

not be overlooked that they also have acidic sites mainly in the form of Lewis acid sites. These

acidic sites are known to favor dehydration reactions which result in the formation of by-products

such as acetaldehyde and hydroxyacetone. As temperature increased, liquid product concentrations

decreased over Ni/Mg6Al1.6La0.4. Whereas for all the other catalysts, liquid product concentrations

either remained stable or increased at higher temperatures. It therefore appears that the

Ni/Mg6Al1.6La0.4 favored the conversion of oxygenated hydrocarbons (OHCs) into CH4, CO, CO2

and H2. This validates the higher total glycerol conversions as well as glycerol conversion to

gaseous products.

116


Catalysts were also characterized after test to help study the coke deposition in order to help make

correlations with the catalytic activity. The TG and DSC curves recorded for the used catalysts are

presented in Figure 4.20 (a) and (b) respectively.

Figure 4.20: a) TG and b) DSC curves obtained for used Ni/Mg6Al2, Ni/Mg6Al1.8La0.2, Ni/Mg6Al1.6La0.4 and Ni/Mg6Al1.2La0.8

The TG curves of the spent catalysts indicate weight losses due to the removal of coke. All four

catalysts produced small quantities of coke as evidenced by the small weight losses. Mg-Al

catalysts in general are known to exhibit minimal coke deposition due to their basic properties

(shown in CO2-TPD, Figure 4.18). Their basic sites enhance the oxidation of CHx fragments

adsorbed on the active metal. The addition of La2O3 as a promoter further led to carbon gasification

[35]. The hydrophilic character of La2O3 facilitates the gasification of coke via the following

endothermic reactions:

C + H2O CO + H2 ΔH298K = 131 kJ/mol

C + 2H2O CO2 + 2H2 ΔH298K = 90 kJ/mol

Additionally, the basic sites of lanthanum containing catalysts enhanced the adsorption of CO2 (as

shown in CO2-TPD) and this also helps in the removal of coke via the reverse Boudouard reaction

[164]:

CO2 + C 2CO ΔH298K = 172 kJ/mol

Ni/Mg6Al2 and Ni/Mg6Al1.2La0.8 exhibited the lowest weight losses of 1.5% and 1.8% respectively.

The low catalytic activity Ni/Mg6Al2 explains the least quantity of coke produced. Whereas for

100 200 300 400 500 600 700 80094

96

98

100

102(a)

1.5%

2.5%

We

igh

t L

oss (

%)

Temperature (oC)

Ni/Mg6Al

2

Ni/Mg6Al

1.8La

0.2

Ni/Mg6Al

1.6La

0.4

Ni/Mg6Al

1.2La

0.8

1.8%

2.5%

100 200 300 400 500 600 700 800

(b)

379oC 442

oC

450oC

440oC

566oC

580oC

Ni/Mg6Al

2

Ni/Mg6Al

1.8La

0.2

Ni/Mg6Al

1.6La

0.4

Ni/Mg6Al

1.2La

0.8

DS

C s

ign

al (a

.u.)

Temperature (oC)

460oC

117

Ni/Mg6Al1.2La0.8, the low coke formation is attributed to carbon gasification resulting from the

large quantity of La.

Ni/Mg6Al1.8La0.2 and Ni/Mg6Al1.6La0.4 exhibited similar weight losses (2.5%) despite the

significant difference of activity observed between both. It appears that the higher concentration

of liquid products produced over Ni/Mg6Al1.8La0.2 led to the formation of undesired coke

precursors.

Looking at the DSC profiles in Figure 4.20(b), the exothermic peaks observed correspond to the

different types of coke formed. Ni/Mg6Al2 produced two types of coke. The first type is oxidized

at a higher temperature of 580oC and the second type at a lower temperature of 460oC. As

previously discussed, coke oxidized at a lower temperature (<500oC) is generally referred to as

encapsulating or non-filamentous coke whereas coke oxidized at a higher temperature (>500oC) is

whisker-like or filamentous coke. Therefore, the Ni/Mg6Al2 catalyst produced both filamentous

and non-filamentous coke. The Ni/Mg6Al1.8La0.2 also produced both filamentous and non-

filamentous coke but with lower oxidation temperatures. In fact, with increasing quantities of

lanthanum in the catalyst composition, the coke formed became more amorphous and easier to

oxidize as evidenced by lower temperatures.

As discussed previously, dehydration products in particular formed on acidic sites evolve into coke

precursors for encapsulating coke. This can be linked to the results as all catalysts produced

encapsulating coke resulting from the high concentrations of the dehydration products,

acetaldehyde and hydroxyacetone.

4.2.8 Discussion

Following impregnation of the prepared supports with 5 wt% Ni and calcination at 600oC, the

obtained solids were characterized. X-ray diffractograms showed that the periclases (MgO and

NiO) were damaged to a certain degree by lanthanum incorporation. H2-TPR analysis showed

varying metal-support interactions and dispersions depending on the lanthanum content. A small

quantity of La resulted in a stronger metal-support interaction whereas larger quantities weakened

this interaction. This was attributed to an influence on metal dispersion. CO2-TPD profiles showed

an enhanced basicity of all lanthanum containing catalysts as indicated by higher proportions of

medium strength and strong basic sites as lanthanum content increased. Thus, the addition of La

results in structural modifications that directly affect the acid-basic properties of the catalyst.

118

Following catalytic tests, the Ni/Mg6Al1.6La0.4 catalyst exhibited the best catalytic performance as

its lanthanum content was optimum giving the highest hydrogen yield and glycerol conversion to

gaseous products. Liquid product analysis showed that the Ni/Mg6Al1.6La0.4 also exhibited the

highest total glycerol conversions at higher temperatures and as temperature increased liquid

product concentrations decreased.

Thermal analysis of the used catalysts showed the different types of coke formed over each

catalyst. The presence of encapsulating coke over all catalysts is attributed to the dehydration

products formed from the Lewis acid sites of the supports.

119

4.3 Parameter Modification: Effect of a higher Ni loading (15%Ni)

and bimetallic effect (1%Ru-5%Ni)

As the Ni/Mg6Al1.6La0.4 catalyst exhibited the best catalytic performance, it is important to study

the effect of using a higher nickel loading (15%) as well as the bimetallic effect (1%Ru-5%Ni) on

the catalytic performance.

4.3.1 Synthesis

The same procedure of wet-impregnation described previously was used but in this case the

amount of nickel nitrate added was varied such as to obtain a 15 wt% loading of Ni in the final

solid. A bimetallic catalyst was prepared by simultaneous addition of an adequate quantity of: 1)

the ruthenium precursor (Ru(NO)(NO3)3 (1.5 wt% Ru)) to obtain a 1 wt% Ru loading and 2) the

nickel precursor to obtain a 5 wt% Ni loading. Both prepared catalysts were calcined at 600oC.

Catalysts were named: Ru-Ni/Mg6Al1.6La0.4 and 15Ni/Mg6Al1.6La0.4. For more clarity, the

Ni/Mg6Al1.6La0.4 was renamed to 5Ni/Mg6Al1.6La0.4 in this section.

4.3.2 Characterization

4.3.2.1 X-ray diffractograms (XRD)

X-ray diffractograms of the calcined catalysts are shown in Figure 4.21.

20 30 40 50 60 70 80

Ru-Ni/Mg6Al

1.6La

0.4

15Ni/Mg6Al

1.6La

0.4

5Ni/Mg6Al

1.6La

0.4

##

#

#

Inte

nsity (

a.u

.)

2(o)

#

#

#

#

#

#

* Hydrotalcite phase # MgO NiO La2O

2CO

3

Figure 4.21: XRD diffractograms of 5Ni/Mg6Al1.6La0.4, Ru-Ni/Mg6Al1.6La0.4 and 15Ni/Mg6Al1.6La0.4

NiO peaks became more intense for the 15Ni/Mg6Al1.6La0.4 due to the higher Ni loading [157].

This indicates a larger NiO crystallite size. Peaks corresponding to the La2O2CO3 phase are also

120

observed for all catalysts. It is noted that the bimetallic catalyst exhibits more intense peaks for all

phases (NiO, MgO, La2O2CO3) compared to 5Ni/Mg6Al1.6La0.4. This indicates a higher level of

crystallinity. Many authors have reported that for bimetallic catalysts, the disappearance of NiO

peaks or difficulty in observing them is an indication of an improved dispersion of the nickel

species [165]. However, this was not the case for the bimetallic catalyst in this study.

4.3.2.2 H2-Temperature Programmed Reduction (H2-TPR)

H2-TPR profiles of the catalysts are shown in Figure 4.22.

100 200 300 400 500 600 700

190oC

Ru-Ni/Mg6Al

1.6La

0.4

15Ni/Mg6Al

1.6La

0.4

Inte

nsity (

a.u

.)

Temperature (oC)

5Ni/Mg6Al

1.6La

0.4

220oC 360

oC

508oC

540oC

423oC

357oC

220oC

310oC 470

oC

575oC

Figure 4.22: H2-TPR profiles of 5Ni/Mg6Al1.6La0.4, Ru-Ni/Mg6Al1.6La0.4 and 15Ni/Mg6Al1.6La0.4

The profiles of all catalysts show three peaks in common: a low temperature peak corresponding

to the reduction of free NiO and higher temperature peaks attributed to the reduction of NiO weakly

interacting with the support. In addition, the Ru-Ni/Mg6Al1.6La0.4 has a very small peak at around

190oC corresponding to the reduction of RuO2. A shoulder peak is also additionally observed for

15Ni/Mg6Al1.6La0.4 at 423oC also attributed to reduction of NiO weakly interacting with the

support. Differences in the temperatures of the reductions peaks are an indication of differing

metal-support interactions in each catalyst. The lower reduction temperatures of peaks in both

15Ni/Mg6Al1.6La0.4 and Ru-Ni/Mg6Al1.6La0.4 compared to 5Ni/Mg6Al1.6La0.4 suggest weaker

metal-support interactions as well as a decreased Ni dispersion on these catalysts. The larger NiO

crystallites observed in XRD diffractograms (Figure 4.21) for both these catalysts could account

for the weaker metal-support interactions [166].

121

Hydrogen consumption values are compared in Table 4.9:

Table 4.9: Hydrogen consumption values of 5Ni/Mg6Al1.6La0.4, Ru-Ni/Mg6Al1.6La0.4 and 15Ni/Mg6Al1.6La0.4

With the additional reduction of RuO2, the Ru-Ni/Mg6Al1.6La0.4 had a higher hydrogen

consumption value compared to 5Ni/Mg6Al1.6La0.4. The higher Ni loading of 15Ni/Mg6Al1.6La0.4

also explains the higher hydrogen consumption value as more NiO was reduced in the studied

temperature range. All catalysts had lower experimental H2 consumption values than the

corresponding theoretical values indicating that not all oxides (NiO, RuO2) were reduced in the

studied temperature range.

4.3.2.3 CO2-Temperature Programmed Desorption (CO2-TPD)

CO2-TPD profiles are shown in Figure 4.23.

100 200 300 400 500 600 700

15Ni/Mg6Al

1.6La

0.4

Ru-Ni/Mg6Al

1.6La

0.4

5Ni/Mg6Al

1.6La

0.4

Inte

nsity (

a.u

.)

Temperature (oC)

I

IIIII

Figure 4.23: CO2-TPD profiles of 5Ni/Mg6Al1.6La0.4, Ru-Ni/Mg6Al1.6La0.4 and 15Ni/Mg6Al1.6La0.4

CO2-TPD profiles showed that the bimetallic catalyst Ru-Ni/Mg6Al1.6La0.4 had a lower basicity

than the 5Ni/Mg6Al1.6La0.4. This is further verified by the basicity values of the weak and medium


experimental


H2 consumption

theoretical


15Ni/Mg6Al1.6La0.4 795 3007

Ru-Ni/Mg6Al1.6La0.4 310 1096

5Ni/Mg6Al1.6La0.4 255 896

122

basic sites shown in Table 4.10. The 15Ni/Mg6Al1.6La0.4 on the other hand exhibited a higher

basicity than 5Ni/Mg6Al1.6La0.4.

Table 4.10: Basicity values of 5Ni/Mg6Al1.6La0.4, Ru-Ni/Mg6Al1.6La0.4 and 15Ni/Mg6Al1.6La0.4

Catalyst Basicity


15Ni/Mg6Al1.6La0.4 362

Ru-Ni/Mg6Al1.6La0.4 261

5Ni/Mg6Al1.6La0.4 304

123



selectivities of 5Ni/Mg6Al1.6La0.4, Ru-Ni/Mg6Al1.6La0.4 and 15Ni/Mg6Al1.6La0.4.

400 450 500 550 600 650 7000

10

20

30

40

50

60

70

80

90

100(c)

5Ni/Mg6Al

1.6La

0.4

Ru-Ni/Mg6Al

1.6La

0.4

15Ni/Mg6Al

1.6La

0.4CO

2 S

ele

ctivity (

%)

Temperature (oC)

400 450 500 550 600 650 7000

10

20

30

40

50

60

70

80

90

100(d) 5Ni/Mg

6Al

1.6La

0.4

Ru-Ni/Mg6Al

1.6La

0.4

15Ni/Mg6Al

1.6La

0.4

CO

Sele

ctivity (

%)

Temperature (oC)

400 450 500 550 600 650 7000

10

20(e) 5Ni/Mg

6Al

1.6La

0.4

Ru-Ni/Mg6Al

1.6La

0.4

15Ni/Mg6Al

1.6La

0.4

CH

4 S

ele

ctivity (

%)

Temperature (oC)

400 450 500 550 600 650 7000.0

0.2

0.4

0.6

0.8

1.0

1.2

1.4

1.6

1.8

2.0

2.2

2.4(a)

Hydro

gen y

ield

(m

ol H

2/m

ol gly

cero

l)

Temperature (oC)

5Ni/Mg6Al

1.6La

0.4

Ru-Ni/Mg6Al

1.6La

0.4

15Ni/Mg6Al

1.6La

0.4

400 450 500 550 600 650 7000

10

20

30

40

50(b) 5Ni/Mg

6Al

1.6La

0.4

Ru-Ni/Mg6Al

1.6La

0.4

15Ni/Mg6Al

1.6La

0.4

Gly

cero

l C

onvers

ion G

aseous (

%)

Temperature (oC)


reaction over 5Ni/Mg6Al1.6La0.4, Ru-Ni/Mg6Al1.6La0.4 and 15Ni/Mg6Al1.6La0.4 (flow rate: 0.025 mL/min,


124

Test results demonstrated that in terms of hydrogen yields, catalytic activities were somewhat

similar between all catalysts. More evident differences were observed in glycerol conversions to

gaseous products and gaseous product selectivities revealing different mechanisms over each

catalyst.

At first, the bimetallic catalyst exhibited a similar but slightly lower glycerol conversion to gaseous

products as 5Ni/Mg6Al1.6La0.4 until 550oC. Beyond 550oC, this conversion became even lower.

Looking at the selectivities, higher CO2 and lower CO selectivities are an indication of the favoring

of the water-gas shift reaction due to the presence of ruthenium [133]. The abrupt change in trends

observed beyond 550oC is also attributed to the ruthenium which clearly resulted in a change in

the reaction mechanism. Methanation reactions were more favored over the bimetallic catalyst at

higher temperatures.

For 15Ni/Mg6Al1.6La0.4, at lower temperatures (400-500oC), the higher nickel loading resulted in

a higher hydrogen yield and glycerol conversion to gaseous products. However, beyond 500oC the

higher nickel loading did not improve the catalytic performance. This could be attributed to the

agglomeration of nickel which inhibits the catalytic activity [167]. XRD showed larger NiO

crystallite sizes and H2-TPR analysis speculated a decreased Ni dispersion in the

15Ni/Mg6Al1.6La0.4 catalyst. Similar results were also observed with Zarei Senseni et al. [67].

CO2 selectivity increased with the nickel loading whereas CO and CH4 selectivities decreased. The

conversion of glycerol to the gaseous products H2, CO and CO2 depends on the preferential

cleaving of C-C bonds instead of C-O bonds. Since nickel is known for its strong affinity to

breaking C-C bonds, this is why a decrease in CO selectivity was observed for the catalyst with a

higher Ni loading. The lower CO and higher CO2 selectivities could also be an indication of a

favoring of the Boudouard reaction:


The decreasing methane selectivities for the 15Ni/Mg6Al1.6La0.4 could also be due to the methane

decomposition reaction:


125


The figures showing the results of the liquid product analysis are in Appendix B. Total glycerol

conversion values showed that the 15Ni/Mg6Al1.6La0.4 experienced the highest total glycerol

conversion over all temperatures. This is expected given nickel’s ability to cleave C-C bonds.

Larger quantities of acetaldehyde as well as ethylene glycol were produced at lower temperatures

by 15Ni/Mg6Al1.6La0.4 indicating that there is a higher formation of liquid by-products with

increased Ni loading. The bimetallic catalyst experienced the lowest total glycerol conversion

values. This can be correlated to the liquid concentration values where the Ru-Ni/Mg6Al1.6La0.4

also experienced a lower formation of by-products compared to 5Ni/Mg6Al1.6La0.4 and

15Ni/Mg6Al1.6La0.4.


The TG and DSC curves recorded for the used catalysts are presented in Figure 4.25(a) and (b)

respectively.

TG curves show that the bimetallic Ru-Ni/Mg6Al1.6La0.4 experienced a lower weight loss (1.4%)

compared to 5Ni/Mg6Al1.6La0.4 (2.5%). This could be due to the synergetic bimetallic effect which

results in a lower coke formation [48]. The 15Ni/Mg6Al1.6La0.4 exhibited a significantly larger

weight loss (10%) indicating the formation of much larger quantities of coke than the

Figure 4.25: a) TG and b) DSC curves obtained for used 5Ni/Mg6Al1.6La0.4, Ru-Ni/Mg6Al1.6La0.4 and 15Ni/Mg6Al1.6La0.4

100 200 300 400 500 600 700 80080828486889092949698

100102104

We

igh

t L

oss (

%)

Temperature (oC)

5Ni/Mg6Al

1.6La

0.4

Ru-Ni/Mg6Al

1.6La

0.4

15Ni/Mg6Al

1.6La

0.4

1.4%

2.5%

10%

100 200 300 400 500 600 700 800

5Ni/Mg6Al

1.6La

0.4

Ru-Ni/Mg6Al

1.6La

0.4

15Ni/Mg6Al

1.6La

0.4

425oC

440oC

DS

C s

ign

al (a

.u.)

Temperature (oC)

451oC

590oC

126

5Ni/Mg6Al1.6La0.4. This is expected given the higher quantity of nickel as nickel is known to cause

coke deposition.

DSC profiles showed that the bimetallic catalyst produced more amorphous coke as indicated by

the lower temperature peak. Interestingly, the 15Ni/Mg6Al1.6La0.4 produced filamentous coke in

addition to the non-filamentous coke. It appears that the lower CO and higher CO2 selectivities in

the test results were indeed as a result of the occurrence of the Boudouard reaction. Furthermore,

the methane decomposition reaction also led to lower CH4 selectivities. As mentioned previously,

CO from the Boudouard reaction and CH4 from the methane decomposition reaction are precursors

of filamentous coke and this was clearly observed in these results. This therefore indicates that the

increased nickel loading resulted in a change in the coke formation mechanism.

4.3.6 Discussion

This section studied the bimetallic effect and effect of a higher Ni loading to see if it would lead

to an improved activity compared to 5Ni/Mg6Al1.6La0.4. The 5Ni/Mg6Al1.6La0.4 remained the most

active in terms of glycerol conversions to gaseous products and hydrogen yields. The bimetallic

effect did not improve the catalytic activity despite the additional presence of a noble metal. The

higher Ni loading resulted in an agglomeration of Ni which later hindered the catalytic activity.

The higher Ni loading also resulted in a different coke formation mechanism as verified by thermal

analysis results after test.

127

4.4 Comparison to Ni-based pure metal oxide

After several catalytic tests, the most efficient catalyst amongst the hydrotalcite derived mixed

oxides was found to be Ni/Mg6Al1.6La0.4. In this section, its activity will be compared to the most

efficient catalyst amongst pure metal oxides from Chapter 3: Ni/ZrO2. Both catalysts have a similar

Ni loading of 5 wt%. Comparison of the hydrogen yields and glycerol conversions to gaseous

products are shown in Figure 4.26:

Hydrogen yields are somewhat similar between both catalysts despite the different supports used.

Nevertheless, a significant difference is noted when comparing glycerol conversion to gaseous

products. The Ni/ZrO2 is clearly more active in converting the glycerol molecule to gaseous

products. Total glycerol conversions (Appendix B) reveal that conversions for Ni/Mg6Al1.6La0.4

are slightly lower than Ni/ZrO2 yet still close. Liquid product concentrations (Appendix B)

obtained over Ni/Mg6Al1.6La0.4 are significantly higher than Ni/ZrO2. This explains the difference

observed in glycerol conversion to gaseous products. The hydrotalcite support clearly favors a

higher formation of liquid by-products whereas the pure metal oxide produces lower quantities.

Nevertheless, both catalysts are successfully able to convert these liquid products to gaseous

products at higher temperatures as demonstrated by decreasing trends of liquid product

concentrations.

Figure 4.26: (a) H2 yields (b) Glycerol conversion to gaseous products of GSR reaction over Ni/ZrO2 and Ni/Mg6Al1.6La0.4

(flow rate: 0.025ml/min, steam/glycerol molar ratio: 9)

400 450 500 550 600 650 7000

1

2

3(a)

Hydro

gen y

ield

(m

ol H

2/m

ol gly

cero

l)

Temperature (oC)

Ni/ZrO2

Ni/Mg6Al

1.6La

0.4

400 450 500 550 600 650 7000

10

20

30

40

50(b) Ni/ZrO

2

Ni/Mg6Al

1.6La

0.4

Gly

cero

l C

onvers

ion G

aseous (

%)

Temperature (oC)

128

Comparing thermal analysis results of the used catalysts (Figure 4.27), it is clear that a significantly

lower quantity of coke is produced over Ni/Mg6Al1.6La0.4 (2.5%) compared to Ni/ZrO2 (16%). This

is attributed to the higher basicity of the hydrotalcite support especially due to the addition of

lanthanum. Furthermore, DSC profiles show that coke produced over the hydrotalcite support is

non-filamentous whereas that of the pure metal oxide support is filamentous.

Figure 4.27: (a) TG and (b) DSC curves obtained for used Ni/ZrO2 and Ni/Mg6Al1.6La0.4

100 200 300 400 500 600 700 80084

86

88

90

92

94

96

98

100

102

104(a)

16%

We

igh

t L

oss (

%)

Temperature (oC)

Ni/ZrO2

Ni/Mg6Al

1.6La

0.4

2.5%

100 200 300 400 500 600 700 800

(b) Ni/ZrO2

Ni/Mg6Al

1.6La

0.4

DS

C s

ign

al (a

.u.)

Temperature (oC)

129

4.5 Stability Test

A stability test was conducted at 600oC to study the activity of the Ni/Mg6Al1.6La0.4 catalyst at a

fixed temperature over a longer period of time. Glycerol conversions to gaseous products,

hydrogen yields and CO2, CO, CH4 selectivities were monitored over 6 hours on stream at 600oC

and the results are presented in Figure 4.28 (a) and (b). Total glycerol conversions and liquid

product distributions taken at 2 hour intervals are presented in Figure 4.28 (c) and (d).


Liquid product concentrations (d) Total glycerol conversions of stability test over Ni/Mg6Al1.6La0.4 (flow rate: 0.025

mL/min, steam/glycerol molar ratio: 9)

0 60 120 180 240 300 3600

10

20

30

40

50

60

70

80

90

100(a) Ni/Mg

6Al

1.6La

0.4

Sele

ctivity to C

pro

du

cts

(%

)

Time (min)

CO2

CH4

CO

0 60 120 180 240 300 3600

10

20

30

40

50

60(b)

Time (min)Gly

cero

l con

vers

ion

to g

aseo

us (

%)

Ni/Mg6Al

1.6La

0.4

0

1

2

3

4

5

6

7 H2 y

ield

(mol H

2 /mol o

f gly

cero

l)

120 240 3600.00

0.05

0.10

0.15

0.20

0.25

0.30

0.35

0.40

0.45

0.50(c)

Co

ncentr

ation (

mol/L)

Time (min)

Acetaldehyde

Acetone

Hydroxyacetone

Ethylene Glycol

Ni/Mg6Al

1.6La

0.4

120 240 3600

10

20

30

40

50

60

70

80

90

100(d) Ni/Mg

6Al

1.6La

0.4

Tota

l G

lycero

l C

onvers

ion (

%)

Time (min)

130

The Ni/Mg6Al1.6La0.4 catalyst experienced a dramatic decrease in both glycerol conversion to

gaseous products and hydrogen yield. The drop became less expressed after the first three hours.

This decrease in activity could possibly be due to an accumulation of coke on the surface of the

catalyst.

Looking at selectivities, the higher CO2 than CO selectivity is an indication of the occurrence of

the water gas shift reaction. The production of small quantities of CH4 was observed over the entire

stability test.

Liquid product analysis verified the loss of activity with time on stream through decreasing total

glycerol conversion values. The concentrations of all the studied liquid products in general were

low at 600oC since liquid product formation is more favored at lower temperatures. As mentioned

previously, at higher temperatures, these products form intermediates that lead to the production

of carbon oxides and hydrogen [141]. Nevertheless, the variation of these liquid product

concentrations with time on stream verifies that coke deposition is associated with the dehydration,

dehydrogenation and hydrogenolysis reactions that take place on catalyst active sites in GSR [142].

Concentrations of hydroxyacetone in particular increased with time on stream.

TG analysis after stability test (Figure 4.29) showed coke formation on the catalyst corresponding

to a weight loss of 8%. The DSC profile revealed an exothermic peak whose maxima had a

temperature of 452oC. This can be attributed to non-filamentous or encapsulating coke. According

to literature, encapsulating coke can block active metal sites which leads to a decrease in

conversion [126]. The results of this stability test therefore verify that the formation of

encapsulating coke affected the catalytic activity via blockage of active metal sites. It also verified

that by-products are the precursors of encapsulating coke as indicated by the varying liquid product

concentrations with time on stream.

131

Figure 4.29: TG and DSC curves obtained for used Ni/Mg6Al1.6La0.4 after stability test (flow rate: 0.025 mL/min)

100 200 300 400 500 600 700 80086

88

90

92

94

96

98

100

102

TG

DSC

Temperature (oC)

Weig

ht Loss (

%)

Ni/Mg6Al

1.6La

0.4

DS

C s

ign

al (a

.u.)

452oC

8%

132

4.6 Optimization

In an attempt to optimize and improve the activity of the Ni/Mg6Al1.6La0.4 which quickly

deactivated in the previous stability test, the flow rate was reduced threefold to 0.008 mL/min to

study the activity with time on-stream. Once again, the stability test was conducted at 600oC and

results are presented in Figure 4.30.

Several differences were noted. Most importantly, the catalyst lasted for 24 hours. There was still

a decrease in the glycerol conversion and hydrogen yields indicating a deactivation but this time

it was more gradual. In addition, glycerol conversion values and hydrogen yields were significantly


Liquid product concentrations (d) Total glycerol conversions of stability test over Ni/Mg6Al1.6La0.4 (flow rate: 0.008


0 500 1000 15000

10

20

30

40

50

60

70

80

90

100(a) Ni/Mg

6Al

1.6La

0.4

Sele

ctivity to C

pro

du

cts

(%

)

Time (min)

CO2

CH4

CO

0 500 1000 15000

10

20

30

40

50

60

70

80

90

100Ni/Mg

6Al

1.6La

0.4

Time (min)Gly

cero

l con

vers

ion

to g

aseo

us (

%)

0

1

2

3

4

5

6

7(b)

H2 y

ield

(mol H

2 /mol o

f gly

cero

l)

1260 14400.00

0.05

0.10

0.15

0.20

0.25

0.30

0.35

0.40

0.45

0.50(c) Ni/Mg

6Al

1.6La

0.4

Co

ncentr

ation (

mol/L)

Time (min)

Acetaldehyde

Acetone

Hydroxyacetone

Ethylene Glycol

1260 14400

10

20

30

40

50

60

70

80

90

100(d) Ni/Mg

6Al

1.6La

0.4

Tota

l G

lycero

l C

onvers

ion (

%)

Time (min)

133

higher than in the previous stability test. This can be attributed to the lower flow rate. By lowering

the flow rate, the gas hourly space velocity was also reduced. Lower gas hourly space velocities

allow more contact time between the reactants and catalyst and therefore result in a higher glycerol

conversion value [140], [167]. In terms of selectivities, higher CO2 than CO selectivities still

verified the occurrence of the water-gas shift reaction. The fluctuations observed for glycerol

conversions to gaseous and hydrogen yields during the entire stability test can be correlated to the

fluctuations of CO2 and CO selectivities. They could be possibly due to the reverse Boudouard

reaction resulting from the presence of La2O3:

CO2 + C 2CO ΔH298K = 172 kJ/mol

It appears that in this stability test, the La2O3 was able to gasify the coke produced to a certain

extent. This could be a reason why the catalyst took a longer time to deactivate. Nevertheless, with

time on stream the catalyst’s ability to gasify the produced coke became lower as indicated by the

weaker fluctuations as well as decreasing glycerol conversions to gaseous and hydrogen yields. It

is also worth noting that CH4 selectivities decreased with time on stream indicating the possible

occurrence of the methane decomposition reaction:


The overall loss of activity is verified by the total glycerol conversions obtained from liquid

products collected during the last 6 hours of the stability test. Total glycerol conversions decreased

from 69% to 51% during the last 6 hours. Liquid product concentrations during these last 6 hours

revealed that hydroxyacetone concentrations increased with time on stream.

Thermal analysis after test (Figure 4.31) showed a small amount of coke formation corresponding

to a weight loss of 3%. Interestingly, the DSC profiles revealed two distinct exothermic peaks

corresponding to two types of coke (filamentous and non-filamentous). The maximas of each peak

were 456 and 580oC respectively. The presence of filamentous coke indicates a change in the coke

formation mechanism resulting from the reduced flow rate. This filamentous coke was formed due

to the occurrence of the Boudouard and methane decomposition reactions under the new flow

conditions. It appears the observed fluctuating the CO2 and CO selectivities were also due to the

Boudouard reaction. The decreasing CH4 selectivities were therefore definitely due to the methane

decomposition reaction. The nonfilamentous, encapsulating coke was formed due to liquid-

134

product formation especially hydroxyacetone and acetaldehyde. Since encapsulating coke is

known to decrease activity, the deactivation observed was ultimately due to the formation of non-

filamentous coke. Nevertheless, less encapsulating coke was produced given the additional

formation of filamentous coke thereby explaining the longer deactivation time. The formation of

filamentous coke ultimately does not affect the stability of the catalyst.

This section clearly demonstrated that the decrease in gas hourly space velocity resulted in a higher

hydrogen yield as well as a higher stability [168]. A lower coke quantity was also observed in

addition to a different coke mechanism which helped prolong the eventual deactivation.

4.7 Conclusion

In this chapter, the activity of Ru-based and Ni-based hydrotalcite derived mixed oxides was

studied in the glycerol steam reforming reaction. Several key effects were studied including: the

effect of preparation method (for Ru catalysts), the effect of promoter addition and promoter

quantity as well as the effect of using a higher nickel loading and a bimetallic composition (for Ni

catalysts). After all the catalyst screening steps, the Ni/Mg6Al1.6La0.4 showed the best activity

amongst the Ni catalysts due to its enhanced catalytic properties resulting from its composition.

These properties include a good metal-support interaction and enhanced basicity. Compared to the

activity of Ni/ZrO2 from Chapter 3, the Ni/Mg6Al1.6La0.4 produced similar hydrogen yields with

lower quantities of coke given its higher basicity. In addition, the hydrotalcite support is much

Figure 4.31: TG and DSC curves obtained for used Ni/Mg6Al1.6La0.4 after stability test (flow rate: 0.008 mL/min)

100 200 300 400 500 600 700 80096

98

100

102

TG

DSC

Temperature (oC)

Weig

ht Loss (

%)

580oC

Ni/Mg6Al

1.6La

0.4

DS

C s

ignal (a

.u.)

3%

456oC

135

cheaper than the pure metal oxide support. Nevertheless, during a stability test the Ni/Mg6Al1.6La0.4

catalyst deactivated within six hours due to coke deposition particularly encapsulating coke which

blocked active Ni sites. An optimization step with a reduced flow rate demonstrated that higher

glycerol conversions and hydrogen yields could be obtained over Ni/Mg6Al1.6La0.4. A higher

stability was also observed resulting from both a decrease in the quantity of coke formed as well

as a change in the coke formation mechanism. The findings of this chapter show that the

Ni/Mg6Al1.6La0.4 can be a promising candidate for the glycerol steam reforming reaction provided

test conditions are further optimized.

136

Conclusion and Perspectives

137

The main aim of this thesis was to evaluate the use of Mg-Al hydrotalcite derived mixed oxides as

catalytic support materials in the transesterification reaction for biodiesel production and the steam

reforming of glycerol reaction for hydrogen production. Different active phases were used for each

reaction: calcium for biodiesel production, ruthenium and nickel for glycerol steam reforming.

This shows the versatility of using Mg-Al hydrotalcites in the catalysis field.

For biodiesel production, the activity of Ca-based Mg-Al catalysts was tested in the

transesterification of sunflower oil. Prior to catalytic testing, sunflower oil was characterized to

ensure its suitability for the transesterification reaction. Following testing, it was found that the

thermal treatment and preparation method both played a role in the catalytic activity. Impregnation

of the Ca active phase on the uncalcined Mg4Al2 support followed by a thermal treatment (600oC)

was clearly the best preparation method as the highest yield was obtained over this catalyst

(Ca600/Mg4Al2HT). Characterization results showed an enhancement of basic properties on this

catalyst. In addition, the optimum reaction conditions for biodiesel synthesis over

Ca600/Mg4Al2HT were found to be a methanol to oil molar ratio of 15:1, a catalyst to oil ratio of

2.5 wt% and a reaction time of 6 hours which gave a final FAME yield of 95%. The properties of

the produced biodiesel were analyzed and satisfied the ASTM (American Society for Testing and

Materials) requirements.

The activity of different Ni-based pure metal oxides was first studied for catalytic glycerol steam

reforming since pure metal oxides are commonly used as support materials. XRD analysis showed

different NiO sizes depending on the support. CO2-TPD results showed varying basicities and H2-

TPR profiles revealed varying stabilities of the active phase depending on the nature of the support.

In the glycerol steam reforming reaction, it was found that the type of support had a significant

impact on the reaction pathway and distribution of obtained products. The most efficient catalyst

for hydrogen production was the monoclinic Ni/ZrO2 catalyst. A tetragonal Ni/ZrO2 was also

prepared for comparison. The zirconia crystalline phase clearly influenced the catalytic properties.

The tetragonal catalyst had smaller NiO crystallites and a larger specific surface area than the

monoclinic catalyst. H2-TPR profiles revealed an increased metal-support interaction for the

tetragonal catalyst due to an improved Ni dispersion. However, a lower hydrogen consumption

value indicated that less NiO was reduced in the studied temperature range. This was why the

tetragonal catalyst was less active than the monoclinic catalyst for glycerol conversion and

138

hydrogen production. The tetragonal catalyst was additionally less stable on stream at 600oC due

to the formation of encapsulating coke which blocked active metallic Ni sites.

For catalytic glycerol steam reforming using Mg-Al hydrotalcite derived mixed oxides, Ru-based

catalysts were first studied. Two Ru-Mg-Al catalysts were synthesized using two different

preparation methods: impregnation and grafting. The physico-chemical properties and catalytic

activities of both catalysts were compared. The grafted catalyst displayed a lower catalytic activity

above 600oC in the GSR reaction. It was therefore found that the impregnation technique was more

effective as it resulted in a higher accessibility of the metallic ruthenium active phase. For the

grafted Ru-Mg-Al catalyst, some of the active phase was “trapped” in the bulk of the material and

thus became less accessible. This was verified by H2-TPR analysis which revealed higher metal-

support interactions in the grafted catalyst evidenced by higher temperature reduction peaks but

overall lower hydrogen consumptions. The activity of the Mg-Al support alone was also studied

and its low activity demonstrated that an active phase is necessary in the GSR reaction.

The activity of Ni-based Mg-Al hydrotalcite derived mixed oxides was then tested. The Mg-Al

supports were modified with La. Characterization results of the supports alone and Ni-impregnated

supports revealed the effect of lanthanum incorporation. For the supports alone, prior to calcination

the addition of lanthanum resulted in a poorer crystallinity of the hydrotalcite phase as lanthanum

carbonate and oxyhydroxide species were formed. After calcination, the La2O2CO3 phase was

observed on all La containing supports resulting from the adsorption of atmospheric CO2 to basic

La2O3 phase. With the increase in lanthanum content, specific surface areas of the supports

decreased. For the Ni-impregnated supports, XRD showed that the periclases (MgO and NiO) were

damaged to a certain degree by lanthanum incorporation. H2-TPR analysis showed varying metal-

support interactions and dispersions depending on the lanthanum content. CO2-TPD profiles

showed an enhanced basicity of all lanthanum containing catalysts. The Ni/Mg6Al1.6La0.4 catalyst

exhibited the best catalytic performance as its lanthanum content was optimum giving the highest

hydrogen yield and glycerol conversion to gaseous products. The effects of a higher Ni loading

and bimetallic effect were then both studied. The bimetallic effect did not improve the catalytic

activity despite the additional presence of a noble metal. The higher Ni loading resulted in an

agglomeration of Ni which later hindered the catalytic activity. The Ni/Mg6Al1.6La0.4 therefore

remained the most active in terms of glycerol conversions to gaseous products and hydrogen

139

yields. The activity of Ni/Mg6Al1.6La0.4 was then compared to the monoclinic Ni/ZrO2. In terms

of hydrogen production, both catalysts exhibited similar activities. However, more liquid by-

products were produced over the hydrotalcite derived catalyst due to its Lewis acid sites.

Nevertheless, the higher basicity resulting from the support composition (Mg-Al-La) led to a much

lower formation of coke. The relatively low cost of Ni/Mg6Al1.6La0.4 also makes it a more

promising candidate for GSR. During a stability test at 600oC, the Ni/Mg6Al1.6La0.4 catalyst

deactivated within six hours due to coke deposition particularly encapsulating coke which blocked

active metallic Ni sites. An optimization step with a reduced flow rate demonstrated that higher

glycerol conversions and hydrogen yields could be obtained over Ni/Mg6Al1.6La0.4. Additionally,

the catalyst lasted for 24 hours under the optimized conditions.

Due to the flow rates utilized in this study, it was possible to conduct a study on coke deposition

as well as some liquid product formation. Although the GSR process is much more complicated,

the following schematic summarizes in a simplistic manner some of the findings and validations

of this work:

The glycerol molecule (C3H8O3) can undergo various pathways during the steam reforming

reaction. Physico-chemical properties of the catalyst utilized determine which pathways are more

favored. The main aim of the steam reforming reaction is the production of CO2 and H2. The

140

glycerol molecule is thermally unstable and at higher temperatures decomposes to give H2, CO

and CH4.

CO and CH4 are both precursors of filamentous coke via the Boudouard and methane

decomposition reactions respectively. CO in the presence of steam (H2O) can undergo the water

gas shift reaction to produce CO2 and H2. The glycerol molecule can also undergo dehydration,

dehydrogenation and polymerization reactions which produce precursors of encapsulating coke.

Such reactions are particularly favored on acidic sites. In this study, acetaldehyde (C2H4O) and

hydroxyacetone (C3H6O2) both stood out as important precursors of encapsulating coke. The

basicity or basic sites of the catalyst governs which secondary reactions are favored and the extent

of carbon gasification/production.

Ni particle size also determines which type of coke is formed. Larger Ni particle sizes favor the

formation of filamentous coke whereas smaller ones favor the formation of encapsulating coke.

Filamentous coke does not affect the stability of the catalyst as it does not block active metallic Ni

sites whereas encapsulating coke affects the stability leading to deactivation due to the

encapsulation of these active metallic Ni sites. These findings can contribute to the search for an

ideal catalyst for industrial application as they help gain more insight into the GSR reaction.

To follow up on this work, it is important to perform additional studies that can help in scaling up

the processes to pilot plant scale. This will make the usage of the catalysts more attractive for an

eventual industrial application. For example, for biodiesel production, it would be interesting to

further optimize the catalyst composition and/or preparation method to obtain higher FAME yields

in shorter time periods. In addition, a catalyst is not considered efficient unless it can be reused for

several cycles. Therefore, a reusability study is essential to guarantee a good activity over several

reaction cycles.

For the glycerol steam reforming reaction, it would also be interesting to vary the catalyst

composition and preparation methods. Further optimization of the test conditions could also be

conducted to increase hydrogen yields and glycerol conversions to gaseous products. For industrial

application, a catalyst must be stable for long periods of time. Therefore, a thorough deactivation

study should also be conducted. Mechanistic studies could also be done where reactive

intermediates are identified (in-situ FTIR) to understand the role of each catalytic component

141

(support/active phase) in the reaction. All catalytic tests in this study were conducted using pure

glycerol purchased from Sigma Aldrich. The glycerol obtained directly from biodiesel production

is actually considered “crude” as it contains several impurities including methanol and soap. Since

the purification of crude glycerol is quite expensive, it would also be interesting to conduct the

steam reforming reaction on this crude glycerol and study the influence of these impurities on the

activity of the catalysts.

142

Appendix A: Biodiesel

143

This appendix contains all the experimental details related to the biodiesel part. This includes the

characterization of the vegetable oil, catalysts characterization, homogeneous catalysis, biodiesel

test set-up and characterization of the produced biodiesel and conventional diesel.

A.1 Characterization of vegetable oil

Four important properties were determined:

The amount of free fatty acid (FFA) (wt. %) in the oil

The saponification value of the oil (in mg KOH/ g of oil)

The amount of moisture/water present in the oil (in wt. %)

The density and viscosity of the oil

Determination of the FFA content of oil

The objective of this test is to determine the amount of free fatty acids in the vegetable oil. As

previously discussed in Chapter 1, free fatty acids could undergo side reactions (mainly

saponification) and this could lower the obtained biodiesel yield. Saponification generally takes

place when the FFA content of the oil exceeds 0.5 wt. %. Oleic fatty acid was considered as the

main analyte since it is the most common fatty acid found in all types of oil. Therefore, for this

test, oleic fatty acid (Molecular Weight: 282 g/mol) was titrated with a KOH solution.

The test was conducted according to the AOCS Ca-5a-40 method. 10 g of refined vegetable oil

were first weighed in an Erlenmeyer flask and 50 mL of isopropyl alcohol and ethanol (V = 1:1)

solution were added to the oil. Phenolphthalein indicator (V = 0.5 mL) was then added and the

solution was mixed until the pink color disappeared. A 50 mL burette was filled with KOH solution

(C = 0.1 M) by dissolving 2.805 g of KOH in 50 mL of distilled water. Titration was stopped when

the pink color appeared and persisted. The titration was repeated three times and the average FFA

content was noted. The FFA (%) was calculated using the following equation:

𝐹𝐹𝐴 (%) =𝑉𝑠 × 𝐶 × 282

𝑚× 100 (A. 1)

Where:

Vs is the volume of titrant (in L)

C is the KOH titrant concentration (in mol/L)

144

m is the mass of the oil (in g)

282 is the molecular weight of the analyte oleic free fatty acid (in g/mol)

Determination of Saponification Value of Oil

This test was conducted according to the AOCS Cd-3-25 method. In this test, the oil was titrated

with a solution of HCl (0.5N). The saponification value is an indication of the amount of KOH (in

mg) required to saponify 1 g of fat.

2 g of the vegetable oil sample were weighed in an Erlenmeyer flask. 25 mL of 0.5 N KOH solution

(prepared by dissolving 2.805 g of KOH in 100 mL of ethanol) were added to the flask. The

resulting solution was boiled on a heater for 15 min and 1 mL of phenolphthalein indicator was

then added to the boiled solution. Titration was stopped when the pink color disappeared. The

volume of titrant required was Vs in mL. The same titration procedure was repeated but without

the oil. This sample was noted as the blank and the volume of titrant used was Vb in mL. The

saponification value was calculated using the following equation:

𝑆𝑉 (𝑚𝑔(𝐾𝑂𝐻)

𝑔) =

(𝑉𝑏 − 𝑉𝑠) × 𝐶 × 56.1

𝑚 (A. 2)

Where:

C is the concentration of hydrochloric acid HCl (0.5 N)

Vb is the volume of titrant required for the blank (in mL)

Vs is the volume of titrant required for the sample (in mL)

m is the mass of the sample (in g)

56.1 is the molecular weight of KOH (in g/mol)

Determination of Water Content in Oil

As previously stated in Chapter 1, the amount of water in the oil should not exceed 0.5 wt. % as it

enhances triglyceride hydrolysis and lowers the FAME yield.

This test was conducted according to the AOCS Ca-2c-25 test method. A well-dried and cooled

Petri dish was weighed and the mass was recorded as mtare. 5g of oil were then weighed in the dish

145

after it was tared. This mass of oil was noted as msample. The filled dish was then placed in the oven

at 105 °C for 4 hours to dry. After 4 hours, the dish was removed, placed in a desiccator for 7

minutes and then weighed once again (mcooled). The moisture or water content was determined

using the equation:

𝑀𝑜𝑖𝑠𝑡𝑢𝑟𝑒 (%) =(𝑚𝑡𝑎𝑟𝑒 + 𝑚𝑠𝑎𝑚𝑝𝑙𝑒) − 𝑚𝑐𝑜𝑜𝑙𝑒𝑑

𝑚𝑠𝑎𝑚𝑝𝑙𝑒× 100 (A. 3)

Where;

mtare is the mass of the empty Petri dish (in g)

msample is the mass of oil (in g)

mcooled is the mass of the filled dish after drying (in g)

Determination of Density and Viscosity of Oil

The density of the oil was measured at 25 °C using an automatic density meter according to the

ASTM D4052 test method. The viscosity of the oil was also determined using a Saybolt and

Redwood viscometer at 40 °C according to the ASTM D445 test method. All measurements were

repeated 3 times and the average value was calculated.

A.2 Catalyst characterization

When developing catalytic systems for any reaction, it is important to characterize and study the

physical and chemical properties of the material in order to gain a better understanding of its

catalytic performance. Diffraction, spectroscopy and adsorption/desorption methods are the usual

characterization techniques used for materials. They give information on the crystallinity,

crystallite size, nature of the active sites, basicity and other characteristics. One of the major goals

in catalysis research is to study how the activity and selectivity of a catalyst can change with its

properties as this helps establish a relationship between the physico-chemical properties and the

activity of a catalyst. This information can also be used to gain insight into the mechanism of the

reaction in question. This section provides a brief review of the characterization techniques used.

All catalysts used in biodiesel production were characterized using the following techniques: N2

adsorption-desorption, X-ray diffraction (XRD) and CO2-Temperature Programmed Desorption

146

(CO2-TPD). Gas chromatography was used to find the FAME yield (%) obtained over each

catalyst.

X-ray Diffraction (XRD)

The crystalline structures of the solid catalysts were analyzed using X-ray diffraction technique.

This technique involves sending X-rays of wavelength to a sample placed on a support. When

these X-rays reach the sample, they are diffracted at an angle by the sample’s adjascent planes

such that they satisfy Bragg’s law:

2dhklsin = n

Where:

dhkl= interplanar spacing

= angle of incidence

n= the order of the reflection (an integer)

= wavelength of incident X-rays

The XRD experiments were conducted on a BRUKER AXS D8 Advance X-ray diffractometer at

ambient temperature. This diffractometer is equipped with a copper anticathode that emits Kα

radiation (=1.5405Å), a LynxEye detector, a /2 goniometer and a rotating sample holder.

Before the actual analysis, the catalyst sample was crushed to a fine powder and then placed on a

glass holder and pressed using a glass slide. Scattering intensities were measured at an angular

range of 5° < 2 < 80° for uncalcined samples and 20° < 2 < 80° for calcined samples. All samples

were measured with a step-size of 2 = 0.02° and a step time of 2 seconds. Peak and phase

identification were conducted by comparing the obtained diffraction patterns with standard XRD

reference patterns from the JCPDS (Joint Committee on Powder Diffraction Standards) database

established by the ICDD (International Center for Diffraction Data). This comparison was made

using EVA software.

Nitrogen adsorption-desorption

The presence of pores at the surface of the catalyst increases the contact surface and the probability

of adsorption of a gaseous mixture on the catalyst. To study textural properties, surface areas and

pore size distributions, N2 adsorption-desorption analysis can be used. This method relies on the

physical adsorption of nitrogen gas at liquid nitrogen temperature (77K). The adsorbed amounts

Figure A.1: Illustration of Bragg’s Law

147

of gas are measured as a function of relative pressure giving an adsorption isotherm. The shape of

the isotherm indicates the nature of the catalyst surface as well as that of the pores.

An ASAP 2020 Micromeritics equipment was used to conduct the N2 adsorption-desorption

experiments. Samples were degassed at 250oC for 4 hours under vacuum to remove impurities

from the sites of adsorption. Textural properties, surface areas and pore diameters were obtained

through analysis of the obtained N2 sorption isotherms. The BET equation was used to calculate

the surface area at various relative pressures:

Where:

P is the partial pressure of Nitrogen

P0 is the saturation pressure at the experimental temperature

V is the volume of N2 adsorbed (cm3/g)

Vm is the volume of N2 adsorbed at monolayer coverage (cm3/g)

C is a constant

S (m2

g) =

𝛼 𝑣𝑚𝑁𝐴

𝑉 (A. 5)

Where:

NA = 6.023*1023 molecules /mol

Vm = 22400 cm3/mol

𝜶 = 16.2*10-20

The desorption curve of the N2 isotherm was used to determine the pore diameter by applying

Barret-Joyner-Halenda (BJH) equation.

𝑃

𝑃0

𝑉(1−𝑃

𝑃0)=

1

𝑉𝑚𝐶 +

(𝐶−1)𝑃

𝑉𝑚𝐶 𝑃0 (A.4)

148

CO2-Temperature Programmed Desorption (CO2-TPD)

Temperature-programmed desorption (TPD) is carried out to measure both qualitatively and

quantitatively the desorbed species from a catalyst under a certain heating rate. The catalyst is

placed in contact with an adsorbate for a certain period of time and this is followed by a temperature

ramp at a specific rate. As the temperature is increased, the chemical bond between the adsorbate

and adsorbent at some point will be overcome. This results in the desorption of the adsorbate. A

TCD detector monitors the changes in the concentration of the desorbed gas.

In this work, CO2-TPD experiments in particular were carried out to study the basicity of each

catalyst. CO2 is used as a probe gas as it has sufficient acidity to evaluate all basic sites. Based on

the amount of desorbed CO2 molecules and the temperature of desorption, the total basicity as well

as the strength of the basic sites can be determined respectively.

The CO2-TPD experiments were conducted on a Zeton Altamira AMI-200 apparatus. Samples (20

mg) were placed in a U-tube quartz reactor and pretreated under helium flow (30 mL/min) at

600 °C for 0.5 h. This is done to ensure the removal of any adsorbed species on the surface such

as water. The catalysts were then cooled to room temperature and exposed to a flow of CO2

(30 mL/min) in helium (4.5% CO2/He) for an hour for adsorption. The samples were then purged

with helium (30 mL/min) for 0.5 h and heated to 700 °C at a rate of 10 °C/min to desorb the CO2.

Gas Chromatography (GC)

The fatty acids methyl esters (FAME) were quantified using a gas chromatography (Agilent)

7890A series equipped with a flame ionization detector (FID) and an HP-5 (30m x 0.321mm)

column. The method utilized is similar to that of Saba et al. [21]. The column temperature was

first kept at 60°C for 2 minutes, raised to 200°C at a rate of 10°C/min, then raised to 240°C at a

rate of 5°C/min and kept at 240°C for 7 min. 50 mg of biodiesel were weighed and then dissolved

in 1 ml of internal standard (C17) solution (10 mg/ml of C17 in hexane). The resulting methyl ester

peaks were identified by comparing them to those of the standard FAME mix C4-C24. The

calculation was done using equation A.6.

𝐹𝐴𝑀𝐸 (%) =∑𝐴

𝐴𝐼𝑆 ×

𝐶𝐼𝑆 × 𝑉𝐼𝑆

𝑊× 100 (A. 6)

149

Where:

ΣA: Total peak area of the methyl esters

AIS: Area of the internal standard (methyl heptadecanoate, C17)

CIS: Concentration of the internal standard solution (10 mg/mL)

VIS: Volume of the internal standard solution (in mL)

W: Weight of the biodiesel sample (in mg)

A.3 Homogeneous Catalysis

KOH was used as a representative homogeneous catalyst since it is able to produce a very high

FAME yield. Therefore, the peaks obtained will be used to calculate the FAME content.

The reaction conditions adopted were the following:

Temperature of 60 °C

Pressure of 1 atm

Stirring rate of 400 rpm

Methanol to oil molar ratio (MOMR) of 6:1

Catalyst to oil ratio (CTOR) of 1 wt%

Reaction time of 1 hour

200 g of sunflower oil were preheated at 60 °C. 2 g of KOH pellets were mixed with 56 mL of

methanol. These quantities were calculated using the following equations:

𝑚𝑐𝑎𝑡𝑎𝑙𝑦𝑠𝑡 (𝑔) =𝐶𝑇𝑂𝑅 × 𝑚𝑜𝑖𝑙

100 (A. 7)

Where:

CTOR: Catalyst to oil mass ratio (in wt%)

moil: mass of oil (in g)

𝑉𝑚𝑒𝑡ℎ𝑎𝑛𝑜𝑙 (𝑚𝐿) =𝑀𝑂𝑀𝑅 × 𝑀𝑊𝑚𝑒𝑡ℎ𝑎𝑛𝑜𝑙 × 𝑚𝑜𝑖𝑙

𝜌𝑚𝑒𝑡ℎ𝑎𝑛𝑜𝑙 × 𝑀𝑊𝑜𝑖𝑙 (A. 8)

150

Where:

MOMR: Methanol to oil molar ratio

MWmethanol: Molecular weight of methanol; 32.04 g.mol-1

ρmethanol: Density of methanol; 0.791 g.cm-3

MWoil: Molecular weight of average sunflower oil; 883.27 g.mol-1

The preheated oil was then added to the mixture of KOH and methanol. The temperature was set

at 60 °C and the stirring rate at 400 rpm. Upon completion of the reaction, the reaction medium

was transferred to a separating decanter to allow complete separation of the obtained biodiesel and

glycerol layers under gravity. The biodiesel was then collected to determine its FAME content by

GC analysis using equation A.6.

151

Figure A.2: Fatty acid methyl esters chromatogram of standard FAME mix C8 – C24

As mentioned previously, in order to help identify the methyl ester peaks produced in each experiment, it is important to compare them

to the peaks obtained by a standard FAME mix C4-C24. The chromatogram obtained from the standard FAME mix C4-C24 is presented

in Figure A.2 and each methyl ester peak is labelled.

152

Figure A.3: Biodiesel chromatogram from homogeneous catalysis

After analyzing the biodiesel produced by homogeneous catalysis (using KOH), nine main fatty acid methyl esters were detected in the

obtained chromatogram. These are represented in Figure A.3 by the asterisks. The calculation of the FAME content was based on the

peak area of each methyl ester (Table A.1).

153

Table A.1: Retention time and peak areas of produced FAMEs from homogeneous catalysis

Retention Time Corresponding Methyl Ester Peak Area Relative Percentage (%)

14.547 Methyl Dodecanoate (C12:0) 4.03162 0.06

16.838 Methyl Myristate (C14:0) 6.457185 0.09

19.214 Methyl Palmitoleate (C16:0) 8.62133 0.12

19.502 Methyl Palmitate (C16:1) 440.4118 6.15

20.947 Methyl Heptadecanoate (IS) (C17:0) 1422.944 -

22.077 Methyl Linoleate (C18:2) 4207.279 58.70

22.160 Methyl Oleate (C18:1) 2185.211 30.50

22.424 Methyl Stearate (C18:0) 243.5072 3.40

25.412 Methyl Arachidate (C20:0) 17.89418 0.25

29.307 Methyl Lignocerate (C24:0) 51.47168 0.72

According to equation A.6:

𝐹𝐴𝑀𝐸 (%) =∑𝐴

𝐴𝐼𝑆 ×

𝐶𝐼𝑆 × 𝑉𝐼𝑆

𝑊× 100

The total FAME yield (%) obtained from the homogeneous catalyst KOH was:

𝐹𝐴𝑀𝐸 (%) =7164.885

1422.944 ×

10 × 1

51.2× 100 = 98.3%

The relative percentages of each fatty acid methyl ester, as shown in the table, were calculated

using the following equation:

𝐹𝐴𝑀𝐸 𝑅𝑒𝑙𝑎𝑡𝑖𝑣𝑒 (%) =𝑃𝑒𝑎𝑘 𝐴𝑟𝑒𝑎

𝑇𝑜𝑡𝑎𝑙 𝑃𝑒𝑎𝑘 𝐴𝑟𝑒𝑎× 100

For example, the relative FAME % of methyl linoleate was:

𝐹𝐴𝑀𝐸 𝑅𝑒𝑙𝑎𝑡𝑖𝑣𝑒 (%) =4207.279

7164.885× 100 = 58.70%

154

A.4 Heterogeneous Catalysis

Biodiesel Test Set-Up

All experiments were conducted in a 500 mL round bottomed flask placed on a heater-magnetic

stirrer and attached to a water cooled condenser (Figure A.4). Prior to the reaction, methanol with

a fixed methanol to oil molar ratio (MOMR) of 12:1 and the catalyst with a specified catalyst to

oil ratio (CTOR) were mixed in the flask at 60oC and a stirring rate of 300 rpm for 5 minutes under

reflux. 20 g of preheated oil at 60 ̊C were then added to the mixture and the stirring rate was

increased to 400 rpm. The reaction was stopped after 4 hours and the reaction medium was allowed

to cool down. Excess methanol was recovered/removed using a rotary evaporator at 60 ̊C for 20

min. The obtained residue was then centrifuged for 30 minutes at 2500 rpm to separate the

biodiesel from the glycerol. Then, GC was used to quantify the biodiesel samples according to

equation A.6. This protocol was carried out for all the prepared catalysts. Optimization of the

reaction conditions of the best performing catalyst was done by varying the catalyst to oil mass

ratio, methanol to oil molar ratio and reaction time.

Figure A.4: Biodiesel test set-up

Heater and

magnetic stirrer

500 mL flask

containing reaction

medium and

magnetic bar

Water in

Water out Reflux condenser

155

Biodiesel and conventional diesel characterization

The properties of the produced biodiesel (FAME) and conventional diesel (purchased from an IPT

gas station in Lebanon) were evaluated and compared using standard ASTM methods. The

equipment and test methods used to determine these fuel properties (kinematic viscosity, density,

flash point, acid number and copper strip corrosion) are highlighted in the table below:

Table A.2: Equipment and test methods used to determine fuel properties

Property Equipment used Test Method

Kinematic Viscosity Saybolt-Redwood

Viscometer (SETA)

ASTM D445

Density Automatic Density meter

(RUDOLPH Research

Analytical)

ASTM D4052

Flash Point SETA PENKSY Closed

Cup Tester

ASTM D93

Acid Number Manual Titration AOCS Ca-5a-40

Copper Strip Corrosion SETA Copper Corrosion ASTM D130

156

Appendix B: Glycerol Steam Reforming

157

This appendix contains all the experimental details related to the glycerol steam reforming

parts. This includes characterization techniques, catalytic test set-up, liquid product

analysis and reaction metrics. The supplementary results of liquid product analysis for

Chapters 3 and 4 are also included.

B.1 Characterization Techniques

All prepared catalysts for GSR were characterized using the following techniques: X-Ray

Diffraction (XRD), Specific surface area by BET method, Fourier Transform Infrared

Spectroscopy (FTIR), H2-Temperature Programmed Reduction (H2-TPR), CO2-Temperature

Programmed Desorption (CO2-TPD) and Simultaneous Thermogravimetric- Differential Scanning

Calorimetry (TG-DSC).

X-Ray Diffraction (XRD)

The same method description given in Appendix A was also utilized.

Lattice Parameters

Hydrotalcites crystallize in rhombohedral 3R symmetry with hexagonal unit cells. The lattice

parameters of interest are “a” which is a function of the average difference between the metal ions

in the host layers, and “c” which is equal to 3c’ where c’ is the distance from the edge of one layer

to edge of the next layer (Figure B.1).

C’

C’

C’

C

a

a a

Figure B.1: Schematic diagram of lattice paramters “a” and “c”

158

These parameters are calculated knowing the lattice planes (hkl) and the spacing between adjacent

(hkl) lattice planes; dhkl given by the formula:

dhkl = 1

√4(h2 + k2 + hk)|3a2 + l2|c2 (𝐵. 1)

Lattice parameter “a” is calculated from the (110) plane and a relation of a=2d110 is obtained.

Lattice parameter “c” is calculated from the (003) plane and a relation of c=3d003 is obtained.

Crystallite Sizes

To determine the size of the crystallites in the catalyst powders, the Scherrer equation is used:

𝜏 =𝜅𝜆

𝛽𝑐𝑜𝑠𝜃 (𝐵. 2)

where:

τ is the size of the crystallites

κ is a dimensionless shape factor, with a value close to unity. The shape factor has a typical

value of 0.9

λ is the X-ray wavelength (1.5406Å)

β is the line broadening at half the maximum intensity (FWHM) in radians

θ is the Bragg angle

Determination of specific surface area using BET method

The surface of solids contains defects and pores of variable sizes. The specific surface area is

directly related to the number of pores on the surface of the solid. Determination of this specific

surface area is essential because a higher surface area is often an indication of a more active

catalyst. A larger surface area provides more possible space for deposition of the active phase, thus

increasing possible adsorption sites for reactant molecules.

Specific surface areas of the catalysts were calculated by “Brunauer-Emmett-Teller” method using

a Thermo Electron Qsurf M1 apparatus where N2 was used as an adsorbate. A mass within the

range of 20 to 75 mg was introduced into a BET sample tube. A degas step was conducted before

the actual measurement of the surface area to clean the catalyst sample from impurities and any

organic or water vapors. This is done by heating under vacuum. Before measurement, the samples

http://en.wikipedia.org/wiki/Intensity_(physics)

http://en.wikipedia.org/wiki/Full_width_at_half_maximum

http://en.wikipedia.org/wiki/Radian

http://en.wikipedia.org/wiki/Bragg_diffraction

159

were degassed as follows: Uncalcined samples were heated for 45 minutes at 55oC. Calcined

samples were heated for 30 minutes at 130oC.

The sample is weighed again to determine the mass lost during the degassing step. The

measurement then takes place in two steps: adsorption and desorption, during which the change in

thermal conductivity of a gas mixture of 30% nitrogen (adsorbed gas) and 70% helium (carrier

gas) is measured. Adsorption takes place at the temperature of liquid nitrogen (-196 ° C). This

temperature allows the nitrogen molecules to be physisorbed. When this step is completed, the

sample is brought back to room temperature which then induces a desorption of the nitrogen

molecules. This desorption is then quantified using a thermal conductibility detector (TCD).

Fourier Transform Infrared Spectroscopy (FTIR)

Fourier Transform Infrared (FTIR) is a spectroscopic method that involves passing an IR radiation

through a sample. Some of this infrared radiation is absorbed by the sample while a fraction is

transmitted. The signal obtained from this process is measured and digitized, then sent to a

computer where the mathematical Fourier transformation takes place to convert the raw data into

an actual spectrum. This spectrum represents the sample’s energy absorption and transmission.

Each peak represents different molecular vibration modes which can be uniquely associated with

specific chemical bonds.

In this work, a Perkin Elmer System 2000 FTIR Spectrometer was used. Samples were prepared

using the KBr pellet method. Before measurement, the background spectrum of the KBr was

recorded. For sample analysis, 2mg of each catalyst were grinded in a mortar with 198mg of KBr

such that the quantity of each solid in KBr is 1%. The resulting powder was then pelletized under

a pressure of 10 tonnes during 5 minutes. IR spectrums were generated in the range of 400-4000

cm-1 with a total of 32 scans and a 2cm-1 resolution. Spectra processing was carried out using

OPUS spectroscopy software.

H2-Temperature Programmed Reduction (H2-TPR)

H2-Temperature-programmed reduction (H2-TPR) is a technique widely used for the

characterization of metal oxides. A calcined catalyst is subjected to a programmed temperature

rise while a reducing gas mixture is passed over it. This method helps study the reducibility of the

160

oxides in the catalyst in order to determine the most efficient reduction conditions required for the

studied reaction. The reduction process is best expressed by the following equation:

MxOy + yH2 → xM + yH2O (B.3)

where MxOy represents the metal oxide while M is the reduced metal

TPR experiments were conducted on a Zeton Altamira AMI-200 apparatus to determine the

reducible species and their corresponding reduction temperatures. The calcined catalyst powder

(50 mg) was placed in a U-tube quartz reactor and purged under an argon flow (30 mL/min) for 1

hour at 150oC to eliminate any physisorbed water. A 5 volume% of H2/Ar was then fed over the

sample under atmospheric pressure, at a continuous flow rate of 30 mL/min. The temperature was

increased from room temperature to 700°C at a rate of 5oC/min. During the analysis, water was

trapped in a bath of ice and salt. The variation of the amount of H2 consumed as a function of

temperature is recorded by means of the thermal conductivity detector (TCD).

CO2 – Temperature Programmed Desorption (CO2-TPD)

The same method description given in Appendix A was also utilized.

Thermal Analysis: Simultaneous Thermogravimetric- Differential Scanning

Calorimetry (TG-DSC)

Under the effect of temperature, a solid can undergo changes in its physical and chemical

properties. Such changes can be accompanied by a release (exothermic reaction) or absorption

(endothermic reaction) of heat and by a loss or gain of mass. Simultaneous TG/DSC combines two

thermal analysis techniques: The first being thermogravimetric analysis (TG) which is used to

determine the thermal stability of a solid by monitoring the weight changes that occur as this solid

is heated. The second is differential scanning calorimetry (DSC) which is used to measure the

endothermic and exothermic transitions of a solid as a function of temperature.

These two techniques were carried out simultaneously using a Setaram Labsys EVO apparatus.

Two alumina crucibles are symmetrically placed on a support inside a furnace. The first crucible

is the reference crucible and is empty. Solids are introduced into the second crucible and then

heated from room temperature to 900°C at a rate of 5°C/min and under an air flow equal to 100

ml/min. A thermocouple system controls and measures the temperature of the sample. The

161

measured difference between the sample and the reference allows gravimetric analysis (loss or

gain of mass of the sample) and thermal differential (temperature difference) analysis.

B.2 Catalytic Test Set-Up

The glycerol steam reforming (GSR) catalytic tests were conducted at atmospheric pressure at a

temperature range of 400 to 700oC. The experimental set-up is shown in Figure B.2:

Figure B.2: Catalytic test set-up for glycerol steam reforming

The experimental set-up consists of a fixed-bed quartz reactor placed in a temperature programmed

electrical furnace. The temperature of the catalyst bed is measured using a thermocouple

positioned at the level of the catalyst. The liquid glycerol-water feed mixture is injected by means

of a TELEDYNE ISCO 500D syringe pump and passes through a liquid evaporator where it is

heated to 250oC such as to evaporate the solution completely before the reaction. It is then mixed

with the carrier gas (Ar) before entering the reactor. The outlet of the reactor is connected to a

condenser immersed in a salty ice bath.

Glycerol/water

feed mixture

Liquid Evaporator

Thermocouple

Furnace

Reactor

Acquisition

Micro GC

Condenser

Condensate

Thermocouple

He

Ar

Mass flow

meter

162

Catalysts (0.1g; particle size: 0.355-0.85 mm) were placed in the reactor. To avoid hot spots and

preferential gas flows, catalysts were diluted with an equal amount of inert SiC (0.1g). Prior to

catalytic testing, all catalysts were reduced in-situ under a 50ml/min H2 flow (50 vol% in Ar) at

700oC for 1 hour at a heating rate of 10oC/min. The catalyst was then purged with Argon (60

mL/min) for 1 hour, the temperature was lowered to 400oC and the feed mixture was introduced

into the reactor. The reactant solution is a 36 wt% aqueous glycerol solution prepared by dissolving

glycerol (Sigma Aldrich 99.0%) in distilled water (9:1 steam/glycerol molar ratio). It is introduced

at a certain flow rate, vaporized at 250oC in the evaporator and further eluted by the carrier gas

Argon (60 mL/min). The catalyst was left for approximately 25 minutes at each temperature

interval step in order to ensure operation at steady state conditions. Condensable, liquid products

were obtained/recuperated for each interval at the end of this 25-minute period. At the end of each

test, all collected condensates were analyzed in a GC-FID (Agilent 7890A) apparatus in order to

identify the by-products formed during the reaction and the extent of total glycerol conversion.

Details of the liquid product analysis method are explained in the next section.

Non-condensable, gaseous products were also analyzed on-line for each temperature interval at

the end of the 25-minute period. This was done by a connected gas chromatograph (Agilent 3000A

MicroGC) with two parallel columns; a molecular sieve column that separates CH4, H2 and CO

using argon as a carrier gas and a Plot U column separating CO2 and with helium as a carrier gas.

The detection of these products was done using a thermal conductivity detector.

Liquid Product Analysis: Gas Chromatography

Liquid products were analyzed by gas chromatography using a 7890A GC equipped with a

split/splitless inlet (Agilent Technologies). The method utilized is similar to that of Papageridis et

al. [45]. Chromatographic separation was achieved using a 30 m x 250 μm HP-5MS (5% phenyl,

95% methylpolysiloxane) capillary column with film thickness of 0.25 μm. Helium 5.0 (99.999%)

was used as the carrier gas at 1 mL/min in a constant flow rate mode. The oven temperature

program was first at 35oC for 5 min, it was then increased at a rate of 10oC/min to 250oC and held

at 250oC for 10 min. The temperature of the split/splitless injector was 280oC and the volume of

the samples injected were 1 μL with a split ratio of 100:1 using ultra inert liner with glass wool

(Agilent Technologies). The reagents and standards used to prepare standard solutions in order to

create calibration curves for each compound were: glycerol, hydroxyacetone, acetone,

163

acetaldehyde and ethylene glycol. 2-butanol was the solvent used to prepare all standard solutions.

A minimum of 4 standard solutions were prepared for the calibration of each compound.

All compounds were purchased from Sigma-Aldrich.

It is important to note that results for all liquid product analysis began from 450oC and not 400oC.

This is because, as mentioned previously, the catalyst was left for 25 minutes at each temperature

interval step in order to ensure operation at steady state conditions. By the end of this 25-minute

interval during the 400oC temperature interval step, no liquid products were obtained. They were

only obtained starting from 450oC. Furthermore, to ensure a homogenous protocol for the

collection of all liquid products, no additional time was added to the 400oC interval to guarantee a

collection.

Reaction Metrics

Catalytic performance is reported in terms of H2 yield, CO, CO2, CH4 selectivity, glycerol

conversion to gaseous products and total glycerol conversion. These performance parameters were

calculated based on the following equations:

H2Yield = H2 moles produced

moles of glycerol in the feedstock (𝐵. 4)

% Selectivity of i = (C atoms in species i

C atoms produced in the gas phase) × 100 (𝐵. 5)

Where species i refers to CO, CO2 and CH4.

% Glycerol conversion (gaseous products) = C atoms in the gaseous products

Total C atoms in the feedstock × 100(𝐵. 6)

% Glycerol conversion (total) = Glycerolin − Glycerolout

Glycerolin × 100 (𝐵. 7)

164

Water-Glycerol System

In this work, a water-glycerol mixture with a specific ratio is prepared and placed in the syringe

pump. Table B.1 shows weight percentages of glycerol solutions and their corresponding molar

compositions in terms of water to glycerol feed ratios and steam to carbon ratios:

Table B.1: Weight percentages, water to glycerol feed ratios and steam to carbon ratios of glycerol solutions

C3H8O3 wt% H2O : C3H8O3 H2O:C

10 46.00 15.33

20 20.44 6.81

30 11.93 3.98

40 7.67 2.56

50 5.11 1.70

60 3.41 1.14

70 2.19 0.73

80 1.28 0.43

For example, a 10 wt% glycerol solution is composed of a 46:1 molar ratio of water and glycerol

respectively.

165

Calculations for Preparation of Water to Glycerol Feed Ratio (WGFR)

A sample calculation for the preparation of a 250 ml water/glycerol mixture is demonstrated in

Table B.2. For example, to prepare a water/glycerol mixture with a feed ratio of 9:1, the molecular

weight of water is multiplied by nine and that of glycerol is multiplied by one (molar fraction of

water and glycerol) to obtain the mass of each component. The volume is obtained by dividing the

mass of each liquid by its corresponding density. The volume percent of each component is

obtained and from this volume percent, the volumes of water and glycerol needed to prepare 250

ml are calculated. Liquids are then measured and mixed together. For a water/glycerol feed ratio

of 9, the steam-to-carbon ratio will be 9/3= 3.

Table B.2: Sample calculation for the preparation of a water/glycerol mixture = 9 (S/C=3) with a total volume of 250 mL

H2O C3H8O3

Molar ratio (H2O : C3H8O3) 9 1

Molar mass (g/mol) 18 92

Density (g/mL) 1 1.26

Mass (g) 162 92

Volume (mL) 162 73.02

Volume Percent (%) 68.93 31.07

Total volume to prepare (mL) 250

Volume to take (mL) 172.33 77.67

166

B.4 Liquid Product Analysis Results – Chapter 3

Figure B.3: Total glycerol conversions and liquid product concentrations of GSR reaction over empty reactor (flow rate: 0.025


450 500 550 600 650 7000.0

0.1

0.2

0.3

0.4

0.5

Concentr

ation (

mol/L)

Temperature (oC)

Acetaldehyde

Acetone

Hydroxyacetone

Ethylene Glycol

Empty Reactor

450 500 550 600 650 7000

10

20

30

40

50

60

70

80

90

100

Tota

l G

lycero

l C

onvers

ion (

%)

Temperature (oC)

Empty Reactor

167

Figure B.4: Total glycerol conversions and liquid product concentrations over Ni/CeO2, Ni/Y2O3 and Ni/ZrO2 (flow rate:

0.025 mL/min, steam/glycerol molar ratio: 9)

450 500 550 600 650 7000.0

0.1

0.2

0.3

0.4

0.5

Concentr

ation (

mol/L)

Temperature (oC)

Acetaldehyde

Acetone

Hydroxyacetone

Ethylene Glycol

Ni/ZrO2

450 500 550 600 650 7000.0

0.1

0.2

0.3

0.4

0.5

Concentr

ation (

mol/L)

Temperature (oC)

Acetaldehyde

Acetone

Hydroxyacetone

Ethylene Glycol

Ni/Y2O

3

450 500 550 600 650 7000.0

0.1

0.2

0.3

0.4

0.5Ni/CeO

2

Co

ncentr

ation (

mol/L)

Temperature (oC)

Acetaldehyde

Acetone

Hydroxyacetone

Ethylene Glycol

450 500 550 600 650 7000

10

20

30

40

50

60

70

80

90

100

Tota

l gly

cero

l convers

ion (

%)

Temperature (oC)

Ni/CeO2

Ni/Y2O

3

Ni/ZrO2

168

Figure B.5: Total glycerol conversions and liquid product concentrations over Ni/ZrO2 (M) and Ni/ZrO2 (T) (flow rate:


450 500 550 600 650 7000.0

0.1

0.2

0.3

0.4

0.5Ni/ZrO

2 (M)

Co

ncentr

ation (

mol/L)

Temperature (oC)

Acetaldehyde

Acetone

Hydroxyacetone

Ethylene Glycol

450 500 550 600 650 7000.0

0.1

0.2

0.3

0.4

0.5Ni/ZrO

2 (T)

Co

ncentr

ation (

mol/L)

Temperature (oC)

Acetaldehyde

Acetone

Hydroxyacetone

Ethylene Glycol

450 500 550 600 650 7000

10

20

30

40

50

60

70

80

90

100

Tota

l gly

cero

l convers

ion (

%)

Temperature (oC)

Ni/ZrO2 (M)

Ni/ZrO2 (T)

169

B.3 Liquid Product Analysis Results – Chapter 4

Figure B.6: Total glycerol conversions and liquid product concentrations over Ni/Mg6Al2, Ni/Mg6Al1.8La0.2, Ni/Mg6Al1.6La0.4 and

Ni/Mg6Al1.2La0.8 (flow rate: 0.025 mL/min, steam/glycerol molar ratio: 9)

450 500 550 600 650 7000.0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1.0Ni/Mg

6Al

2

Concentr

ation (

mol/L)

Temperature (oC)

Acetaldehyde

Acetone

Hydroxyacetone

Ethylene Glycol

450 500 550 600 650 7000.0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1.0Ni/Mg

6Al

1.6La

0.4

Concentr

ation (

mol/L)

Temperature (oC)

Acetaldehyde

Acetone

Hydroxyacetone

Ethylene Glycol

450 500 550 600 650 7000.0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1.0

1.1

1.2Ni/Mg

6Al

1.8La

0.2

Co

ncentr

ation (

mol/L)

Temperature (oC)

Acetaldehyde

Acetone

Hydroxyacetone

Ethylene Glycol

450 500 550 600 650 7000.0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1.0Ni/Mg

6Al

1.2La

0.8

Concentr

ation (

mol/L)

Temperature (oC)

Acetaldehyde

Acetone

Hydroxyacetone

Ethylene Glycol

450 500 550 600 650 7000

10

20

30

40

50

60

70

80

90

100

Tota

l gly

cero

l convers

ion (

%)

Temperature (oC)

Ni/Mg6Al

2

Ni/Mg6Al

1.8La

0.2

Ni/Mg6Al

1.6La

0.4

Ni/Mg6Al

1.2La

0.8

170

Figure B.7: Total glycerol conversions and liquid product concentrations over 5Ni/Mg6Al1.6La0.4, Ru-Ni/Mg6Al1.6La0.4 and

15Ni/Mg6Al1.6La0.4 (flow rate: 0.025 mL/min, steam/glycerol molar ratio: 9)

450 500 550 600 650 7000.0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1.0Ru-Ni/Mg

6Al

1.6La

0.4

Concentr

ation (

mol/L)

Temperature (oC)

Acetaldehyde

Acetone

Hydroxyacetone

Ethylene Glycol

450 500 550 600 650 7000.0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1.015Ni/Mg

6Al

1.6La

0.4C

oncentr

ation (

mol/L)

Temperature (oC)

Acetaldehyde

Acetone

Hydroxyacetone

Ethylene Glycol

450 500 550 600 650 7000

10

20

30

40

50

60

70

80

90

100

Tota

l gly

cero

l convers

ion (

%)

Temperature (oC)

5Ni/Mg6Al

1.6La

0.4

Ru-Ni/Mg6Al

1.6La

0.4

15Ni/Mg6Al

1.6La

0.4

450 500 550 600 650 7000.0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1.05Ni/Mg

6Al

1.6La

0.4

Co

ncentr

ation (

mol/L)

Temperature (oC)

Acetaldehyde

Acetone

Hydroxyacetone

Ethylene Glycol

171

Figure B.8: Total glycerol conversions and liquid product concentrations over Ni/ZrO2 and Ni/Mg6Al1.6La0.4 (flow rate:


450 500 550 600 650 7000

20

40

60

80

100T

ota

l gly

cero

l convers

ion (

%)

Temperature (oC)

Ni/ZrO2

Ni/Mg6Al

1.6La

0.4

450 500 550 600 650 7000.0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1.0

Co

ncentr

ation (

mol/L)

Temperature (oC)

Acetaldehyde

Acetone

Hydroxyacetone

Ethylene Glycol

Ni/ZrO2

450 500 550 600 650 7000.0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1.0Ni/Mg

6Al

1.6La

0.4

Concentr

ation (

mol/L)

Temperature (oC)

Acetaldehyde

Acetone

Hydroxyacetone

Ethylene Glycol

172

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Dissemination of Research Results

Some of the results of this thesis have already been published, and were also presented as oral or

poster communications. Additional papers are submitted or in preparation. Below is a list of the

scientific work during the past three years:

Publications

Glycerol steam reforming over Ru-Mg-Al hydrotalcite-derived mixed oxides: Role of

the preparation method in catalytic activity

Eliane Dahdah, Samer Aouad, Cédric Gennequin, Jane Estephane, Bilal Nsouli, Antoine

Aboukaïs, Edmond Abi-Aad

International Journal of Hydrogen Energy. 43 (2018) 19864-19872

https://doi.org/10.1016/j.ijhydene.2018.09.042

Biodiesel production from refined sunflower oil over Ca-Mg-Al catalysts: Effect of

the composition and the thermal treatment

Eliane Dahdah, Jane Estephane, Samer Aouad, Cédric Gennequin, Reem Haydar, Yara

Youssef, Bilal El Khoury; Antoine Aboukais, Edmond Abi Aad

Renewable Energy, Submitted for publication (2018)

Zirconia supported nickel catalysts for glycerol steam reforming: Effect of zirconia

structure

Eliane Dahdah, Samer Aouad, Cédric Gennequin, Jane Estephane, Bilal Nsouli, Antoine

Aboukaïs, Edmond Abi-Aad

In preparation (2018)

Others:

CO2 reforming of methane over NixMg6−xAl2 catalysts: Effect of lanthanum doping

on catalytic activity and stability

E. Dahdah, J. Abou Rached, S. Aouad, C. Gennequin, H.L. Tidahy, J. Estephane, A.

Aboukaïs, E. Abi Aad



Ni based catalysts promoted with cerium used in the steam reforming of toluene for

hydrogen production

Jihane Abou Rached, Christelle El Hayek, Eliane Dahdah, Cédric Gennequin, Samer

Aouad, Haingomalala Lucette Tidahy, Jane Estephane, Bilal Nsouli, Antoine Aboukaïs,

Edmond Abi-Aad

184



Steam reforming of ethanol for hydrogen production over Cu/Co-Mg-Al-based

catalysts prepared by hydrotalcite route

Doris Homsi, Jihane Abou Rached, Samer Aouad, Cédric Gennequin, Eliane Dahdah, Jane

Estephane, Haingomalala Lucette Tidahy, Antoine Aboukaïs, Edmond Abi-Aad

Environmental Science and Pollution Research. 24 (2017) 9907–9913

https://doi.org/10.1007/s11356-016-7480-9

Dry reforming of methane over NixMg6−xAl1.8La0.2 catalysts

Eliane Dahdah, Jihane Abou Rached, Jane Estephane, Samer Aouad, Cédric Gennequin,

Antoine Aboukaïs, Edmond Abi-Aad

2016 7th International Renewable Energy Congress (IREC)

https://doi.org/10.1109/IREC.2016.7478874

Steam reforming of toluene for hydrogen production over NiMgALCe catalysts

prepared via hydrotalcite route

Jihane Abou Rached, Eliane Dahdah, Cédric Gennequin, Haingomalala Lucette Tidahy,

Samer Aouad, Jane Estephane, Bilal Nsouli, Antoine Aboukaïs, Edmond Abi-Aad

2016 7th International Renewable Energy Congress (IREC)

https://doi.org/10.1109/IREC.2016.7507629

Oral Presentations:

Renewable Energy Sources - Research and Business (RESRB) 2017 Conference (19-

21 June 2017, Wrocław, Poland):

Glycerol steam reforming over Ru-Mg-Al hydrotalcites: role of the preparation method in

catalytic activity

Eliane Dahdah, Samer Aouad, Cédric Gennequin, Jane Estephane, Antoine Aboukaïs, Bilal

Nsouli, Edmond Abi-Aad

Les Journées de l'ARCUS (9-11 April 2018, Université Mohammed V, Rabat,

Morocco) :

Zirconia supported nickel catalysts for glycerol steam reforming: Effect of zirconia

structure

Eliane Dahdah, Samer Aouad, Jane Estephane, Cédric Gennequin, Antoine Aboukaïs, Bilal


https://www.researchgate.net/publication/303563505_Dry_reforming_of_methane_over_NixMg6-xAl18La02_catalysts?_sg=6WUCl1utv8x9JWIzf0dWSPrfi-AuNslihL-yC3T7zyr533mERz4bl7NsTp56VTQS0DgTPAjY6JmwKSbudBUelZSqElYrjpVnbUP_o1AJ.pv64wecT2U0iDdtAjGVvQ0yfF97SBHKcduXfjlDpw4ksgaC9xfDiZ9gzZ_X1jBf5Gxhj8KnfaZv6k-G_fYXweA

https://www.researchgate.net/publication/305332468_Steam_reforming_of_toluene_for_hydrogen_production_over_NiMgALCe_catalysts_prepared_via_hydrotalcite_route?_sg=6WUCl1utv8x9JWIzf0dWSPrfi-AuNslihL-yC3T7zyr533mERz4bl7NsTp56VTQS0DgTPAjY6JmwKSbudBUelZSqElYrjpVnbUP_o1AJ.pv64wecT2U0iDdtAjGVvQ0yfF97SBHKcduXfjlDpw4ksgaC9xfDiZ9gzZ_X1jBf5Gxhj8KnfaZv6k-G_fYXweA

https://www.researchgate.net/publication/305332468_Steam_reforming_of_toluene_for_hydrogen_production_over_NiMgALCe_catalysts_prepared_via_hydrotalcite_route?_sg=6WUCl1utv8x9JWIzf0dWSPrfi-AuNslihL-yC3T7zyr533mERz4bl7NsTp56VTQS0DgTPAjY6JmwKSbudBUelZSqElYrjpVnbUP_o1AJ.pv64wecT2U0iDdtAjGVvQ0yfF97SBHKcduXfjlDpw4ksgaC9xfDiZ9gzZ_X1jBf5Gxhj8KnfaZv6k-G_fYXweA

185

The 24th International Scientific Conference, LAAS24 (25-26 April 2018, University

of Balamand, Koura, Lebanon):

Effect of the preparation method on the catalytic activity of Ru-Mg-Al hydrotalcites for the

steam reforming of glycerol



Others:

The 7th International Renewable Energy Congress (IREC) (22-24 March 2016,

Hammamet, Tunisia)

Dry reforming of methane over NixMg6-xAl1.8La0.2 catalysts

Eliane Dahdah, Jihane Abou Rached, Jane Estephane, Samer Aouad, Cédric Gennequin,

Antoine Aboukaïs, Edmond Abi-Aad

The 7th International Renewable Energy Congress (IREC) (22-24 March 2016,

Hammamet, Tunisia)

Steam reforming of toluene for hydrogen production over NiMgAlCe catalysts prepared

via hydrotalcite route

Jihane Abou Rached, Eliane Dahdah, Cédric Gennequin, Haingomalala Lucette Tidahy,

Samer Aouad, Jane Estephane, Bilal Nsouli, Antoine Aboukaïs, Edmond Abi-Aad

Poster Presentations:

The 23rd International Scientific Conference, LAAS23 (6-7 April 2017, Hadath,

Lebanon):

Glycerol steam reforming over Ru-Mg-Al hydrotalcites: Role of the preparation method in

catalytic activity

Eliane Dahdah, Jane Estephane, Samer Aouad, Cédric Gennequin, Antoine Aboukaïs, Bilal


3ème Colloque International Francophone en Environnement et Santé, (23-24

October 2017, Dunkerque, France):

Vaporeformage du glycérol en présence d’oxydes mixtes Ru-Mg-Al : Effet de la méthode

de préparation sur l'activité catalytique

Eliane Dahdah, Cédric Gennequin, Samer Aouad, Jane Estephane, Antoine Aboukaïs, Bilal


186

Les Journées de l'ARCUS (9-11 April 2018, Université Mohammed V, Rabat,

Morocco):

Catalytic activity of hydrotalcite derived Ru-Mg-Al mixed oxides in the steam reforming

of glycerol



The 24th International Scientific Conference, LAAS24 (25-26 April 2018, University

of Balamand, Koura, Lebanon):

Structural modification of ZrO2 and its influence on the catalytic activity for glycerol steam

reforming



Others:

Journées Franco-libanaises 3 (29-30 October 2015, Université Libanaise, Hadath,

Lebanon) :

L’importance de l’ajout du lanthane aux catalyseurs du type NiMgAl utilisés pour le

reformage à sec du méthane

Eliane Dahdah, Samer Aouad, Cédric Gennequin, Jihane Abou Rached, Mira Nawfal, Jane

Estephane, John El Nakat, Antoine Aboukaïs, Edmond Abi Aad

Journées Franco-libanaises 3 (29-30 October 2015, Université Libanaise, Hadath,

Lebanon) :

Vaporeformage du toluène sur des catalyseurs à base d’hydrotalcite

Jihane Abou Rached, Eliane Dahdah, Antoine Aboukaïs, Cédric Gennequin, Samer Aouad,

Edmond Abi Aad

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