ABSTRACT FORD, JEFFREY PETER. Semi-Batch Catalytic ...

ABSTRACT FORD, JEFFREY PETER. Semi-Batch Catalytic Deoxygenation of Biomass-Derived Fatty Acids and Model Compounds. (Under the direction of Harold Henry Lamb).

A series of supported Pd catalysts was screened for liquid-phase deoxygenation of

stearic, lauric and capric (decanoic) acids under 5% H2 at 300 °C and 15 atm. On-line

quadrupole mass spectrometry (QMS) was used to measure conversion, CO2 selectivity, H2

consumption, and initial decarboxylation rate. Post-reaction analysis of liquid products by

gas chromatograph (GC) was used to determine n-alkane yields. Pd-on-carbon (Pd/C)

catalysts were found to be highly active and selective for stearic acid (SA) decarboxylation

under the test conditions. In contrast, SA deoxygenation over Pd/SiO2 catalysts occurred

primarily via decarbonylation and at a much inferior rate. Pd/Al2O3 exhibited high initial SA

decarboxylation activity but deactivated under the test conditions. Similar CO2 selectivity

patterns among the catalysts were observed for deoxygenation of lauric and capric acid;

however, the initial decarboxylation rates tended to be lower for Pd/C and Pd/Al2O3 with

these substrates. The most active Pd/C catalyst was used to investigate the influence of alkyl

chain length on deoxygenation kinetics for C8-C18 fatty acids (FAs). Generally, as FA

carbon number decreases, reaction time and H2 consumption increase, and CO2 selectivity

and initial decarboxylation rate decrease. The increase in initial decarboxylation rates for

longer chain FAs is attributed to their greater propensities to adsorb onto the activated carbon

support. Literature pertinent to the research in this thesis was also summarized in the

introductory chapter.

Canola and lard-derived FAs were deoxygenated at 300°C in the liquid phase using a

5 wt.% Pd/C catalyst. On-line quadrupole mass spectrometry was used to monitor the

effluent streams from the 50- and 600-ml stirred autoclave reactors. Stearic, oleic, and

palmitic acids were employed as model compounds. H2 consumption during oleic acid (OA)

deoxygenation at the 50-ml scale under 10% H2 occurred during heating to reaction

temperature consistent with double bond hydrogenation. The initial decarboxylation rate of

palmitic acid (PA) under 5% H2 decreased with increasing initial FA concentration in

dodecane; specific semi-batch deoxygenation productivity exhibited a maximum with PA

concentration. Canola-derived fatty acids (CDFA) and a canola FA surrogate mixture also

were deoxygenated at the 50-ml scale. Decarboxylation was inhibited under 10% H2, and

there were indications of catalyst deactivation with CDFA. Low CO2 selectivities and

specific productivities were observed for model compounds and biomass-derived fatty acids

at the 600-ml scale due to the high initial FA concentrations employed. Complete

deoxygenation required substantially longer for OA than stearic acid. When the on-line

QMS traces for deoxygenation of OA and canola-derived FAs were superimposed, there was

no indication of catalyst deactivation attributable to impurities in the latter. However, when

the CDFA deoxygenation product was used as solvent in a subsequent run, the

decarboxylation pathway was inhibited.

Semi-Batch Catalytic Deoxygenation of Biomass-Derived

Fatty Acids and Model Compounds

by Jeffrey Peter Ford

A thesis submitted to the Graduate Faculty of North Carolina State University

in partial fulfillment of the requirements for the degree of

Master of Science

Chemical Engineering

Raleigh, North Carolina

2011

APPROVED BY:

_______________________________ ______________________________ Steven W. Peretti David F. Ollis

________________________________ H. Henry Lamb

Chair of Advisory Committee

ii

BIOGRAPHY

Jeffrey P. Ford was born to Jim and Coleen Ford on March 23rd, 1986 in Granite Falls,

Minnesota. He graduated from Central College in Pella, Iowa in 2008 with a Bachelors of

Arts in Chemistry and Spanish.

iii

DEDICATION

To my Father, who has been with me every step along the way.

iv

TABLE OF CONTENTS

LIST OF TABLES ….………………………………………………………………………vi LIST OF FIGURES ………………………………………………………………………vii LITERATURE REVIEW OF CATALYTIC DEOXYGENATION OF FATTY ACIDS …….……………………………………………1 References………………………………………………………………………......11 LIQUID-PHASE DEOXYGENATION OF FATTY ACIDS: A COMPARISON OF SUPPORTED PD CATALYSTS …….…………………………12 Abstract …………………………………………………………………………….13 Introduction ………………………………………………………………………...14 Experimental ……………………………………………………………………….16 Materials……………………………………………………………………………16 FA Deoxygenation Experiments …………………………………………………...17 Analytical Methods ………………………………………………………………...18 Catalyst Characterization …………………………………………………………..19 Reactor Modeling – Initial Rates Calculation ……………………………………...20 Results and Discussion …………………………………………………………….21 Catalyst Characterization …………………………………………………………..21 Stearic Acid Deoxygenation Under 5 and 10% H2 ……….………………………..22 Lauric Acid Deoxygenation Using Pd/SiO2 (A), Pd/Al2O3 and Pd/C (A) …………26 Decanoic Acid Deoxygenation Using Pd/SiO2 (A), Pd/Al2O3 and Pd/C (A) ……...27 Dependence of Decarboxylation Rate on Fatty Acid Carbon Number …………….28 Deoxygenation of FA Mixtures ………………………………………………….31 Conclusions ………………………………………………………………………32 References ………………………………………………………………………….33 Tables ………………………………………………………………………………35 Figures……………………………………………………………………………...40 CATALYTIC DEOXYGENATION OF BIOMASS-DERIVED FATTY ACIDS OVER 5% PD/C…………………………………………………….……48 Abstract …………………………………………………………………………….49 Introduction ………………………………………………………………………...50 Experimental ……………………………………………………………………….51 Materials……………………………………………………………………………51 50-ml Autoclave Deoxygenation Experiments ……………………………………52 Scale-up Deoxygenation Experiments ……………………………………………..53 Analytical Methods ………………………………………………………………...54 Reactor Modeling – Initial Rates Calculation ……………………………………...55 Results and Discussion …………………………………………………………….56 Analysis of Lard and Canola-Derived Fatty Acids ………………………………...56 Model Compounds …………………………………………………………………56 Canola Surrogate Mixture and Canola-Derived FAs ………………………………59

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Semi-Batch Deoxygenation Scale-up to 600-ml………………...…………………62 Conclusion …………………………………………………………………………68 References ………………………………………………………………….………70 Tables ………………………………………………………………………………72 Figures……………………………………………………………………………...76

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LIST OF TABLES

LITERATURE REVIEW OF CATALYTIC DEOXYGENATION OF FATTY ACIDS

Table 1 Kinetic results from Simakova’s structure sensitivity study ………………...7

LIQUID-PHASE DEOXYGENATION OF FATTY ACIDS: A COMPARISON OF

SUPPORTED PD CATALYSTS

Table 1 Catalyst characterization results……………………………………………35 Table 2 Results for SA deoxygenation using 5 wt.% supported Pd catalysts ………36 Table 3 Results for LA and DA deoxygenation over Pd/SiO2 (A), Pd/Al2O3, and Pd/C (A)………………………………………37 Table 4 Results for deoxygenation of FAs with varying carbon numbers………….38 Table 5 Results for the deoxygenation of SA, DA, and a 50-50 molar mixture of DA and SA …………………………………39

CATALYTIC DEOXYGENATION OF BIOMASS-DERIVED FATTY ACIDS OVER

5% PD/C

Table 1 Composition (mole percent) of biologically derived FAs as determined by GC-FID analysis…………………………...72 Table 2 Results for PA deoxygenation over 5 wt.% Pd/C in a 50-ml stirred autoclave reactor ……………………………………….73 Table 3 Deoxygenation results for canola-derived FAs and a canola FA surrogate mixture in a 50-ml stirred autoclave reactor ……………………74 Table 4 Deoxygenation results for model compounds and biomass-derived FAs in a 600-ml stirred autoclave reactor ………………75

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LIST OF FIGURES

LIQUID-PHASE DEOXYGENATION OF FATTY ACIDS: A COMPARISON OF

SUPPORTED PD CATALYSTS

Figure 1. Correlation between H2 consumption and CO2 selectivity for SA and other FA deoxygenations over 5% Pd/C (A)…………………………………………………………..40 Figure 2. Molar production rate (mmol/min) of CO2 (a) and CO (b), and molar consumption rate of H2 (b), from SA deoxygenation over various 5% Pd catalysts in dodecane solvent at 300ºC for 4 h under 60 ml/min flowing 5% H2(He) at 15 atm………………………………………………41 Figure 3. Molar production rate (mmol/min) of CO2 (a) and

CO (b), and molar consumption rate of H2 (b), from SA deoxygenation over various 5% Pd/C catalysts in dodecane solvent at 300ºC for 4 h under 60 ml/min flowing 5% H2(He) at 15 atm. ……………………………………………42

Figure 4. Molar production rate (mmol/min) of CO2 (a) and CO (b), and molar consumption rate of H2 (b), from LA deoxygenation over Pd/SiO2 (A), Pd/Al2O3, and Pd/C (A) in dodecane solvent at 300ºC for 4 h under 60 ml/min flowing 5% H2(He) at 15 atm ………………………………….43 Figure 5. Molar production rate (mmol/min) of CO2 (a) and CO (b), and molar consumption rate of H2 (b), from DA deoxygenation over Pd/SiO2 (A), Pd/Al2O3, and Pd/C (A) in dodecane solvent at 300ºC for 4 h under 60 ml/min flowing 5% H2(He) at 15 atm ………………………………….44 Figure 6. Molar production rate (mmol/min) of CO2 (a) and CO (b), and molar consumption rate of H2 (b), from FA deoxygenation over Pd/C (A) in dodecane solvent at 300ºC for 4 h under 60 ml/min flowing 5% H2(He) at 15 atm. 0.0056 mol of FA were added for each run …………………….45 Figure 7. Correlation between relative initial decaroboxylation rate, CO2 selectivity and FA carbon number……………………………….46 Figure 8. Molar production rate (mmol/min) of CO2 (a) and CO (b), and molar consumption rate of H2 (b), from the deoxygenation of SA, DA and a 50-50 molar mixture of DA and SA over Pd/C (A) in dodecane solvent at 300ºC for 4 h under 60 ml/min flowing 5% H2(He) at 15 atm. 0.0056 mol of FA were added for each run. …………………………………………………...…….47

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CATALYTIC DEOXYGENATION OF BIOMASS-DERIVED FATTY ACIDS OVER

5% PD/C

Figure 1. GC-FID chromatograms of lard-derived fatty acids (a) and canola-derived fatty acids (b) ……………………………………...76 Figure 2. CO2 and CO molar flow rates and effluent and effluent mol% H2 for OA deoxygenation at 300 °C for 3 h under 10% H2(He) purge in dodecane……………………...77 Figure 3. CO2 (a) and CO (b) production rates in mmol/min and effluent mol% H2 (b) for PA deoxygenation at

300 °C for 5 h in dodecane under 5% H2(He) 60 ml/min at 15 atm in a 50-ml autoclave. The total mass of PA and solvent was held constant at 23.94 g…………………78

Figure 4. CO2 and CO production rates and effluent mol% H2 for 48 wt.% PA deoxygenation at 300 °C for 5 h in dodecane under 5% H2(He) 60 ml/min at 15 atm in a 50-ml autoclave. The initial mass of PA added was 11.52 g. The total mass of PA and solvent was 23.94 g……………………79 Figure 5. CO2 and CO molar production rates and effluent mol% H2 for (a) canola surrogate mixture (b) canola-derived FAs. Reaction conditions: 300 °C for 4 h in dodecane under 5% H2(He) 60 ml/min at 15 atm in a 50-ml stirred autoclave. 5.6 mmol of FA was added to the reactor.……………………………………………………80 Figure 6. CO2 and CO production effluent mol.% H2 for deoxygenation of (a) canola FA surrogate mixture(b) canola-derived FAs. Reaction conditions: 300 °C in dodecane under 10% H2(He) 60 ml/min at 15 atm in a 50-ml stirred autoclave. 5.6 mmol of FA was added to the reactor. ……………………………………………………………….81 Figure 7. CO2 and CO molar production rates and H2 conversion for SA deoxygenation at 300 °C in dodecane under 10% H2(He) 600 ml/min at 15 atm in a 600-ml stirred autoclave. 200 g SA, 100 g dodecane, and 8.375 g Pd/C were added to the reactor. The reactor reached operating temperature (300 °C) at t = 1 h……………………………………………..82 Figure 8. CO2 and CO molar production rates and H2 conversion for OA deoxygenation at 300 °C in dodecane under 10% H2(He) 600 ml/min at 15 atm in a 600-ml stirred autoclave. 200 g OA, 100 g dodecane, and 8.375 g Pd/C were added to the reactor. The reactor reached operating temperature (300 °C) at t = 1 h.…………………………………………….83

ix

Figure 9. CO2 and CO molar production rates and H2 conversion for lard-derived FA deoxygenation at 300 °C in dodecane under 12% H2(He) 600 ml/min at 15 atm in a 600-ml stirred autoclave. 200 g lard-derived FA, 100 g dodecane, and 8.375 g Pd/C were added to the reactor. The reactor reached operating temperature (300 °C) at t = 1 h. ………………………...84 Figure 10. CO2 and CO molar production rates and H2 conversion for lard-derived FA deoxygenation at 300 °C in dodecane under 6% H2(He) 600 ml/min at 15 atm in a 600-ml stirred autoclave. 200 g lard-derived FA, 100 g dodecane, and 8.375 g Pd/C were added to the reactor. The reactor reached operating temperature (300 °C) at t = 1 h. ………………………...85 Figure 11. CO2 and CO molar production rates and H2 conversion for canola-derived FA deoxygenation at 300 °C in dodecane under 10% H2(He) 600 ml/min at 15 atm in a 600-ml stirred autoclave. 198 g canola-derived FA, 100 g dodecane, and 8.375 g Pd/C were added to the reactor. The reactor reached operating temperature (300 °C) at t = 1 h. ………………………...86 Figure 12. CO2 and CO molar production rates and H2 conversion

for canola-derived FA deoxygenation at 300 °C in dodecane under 10% H2(He) 600 ml/min at 15 atm in a 600-ml stirred autoclave. 198 g canola-derived FA, 100 g canola-derived FA deoxygenation product, and 8.375 g Pd/C were added to the reactor. The reactor reached operating temperature (300 °C) at t = 1 h.………………………………………........87

Literature Review of Catalytic Deoxygenation of Fatty Acids

Jeffrey P. Ford

Department of Chemical and Biomolecular Engineering,

North Carolina State Universty, Raleigh, NC 27695-7905, USA

2

Over the last 30 years, the US has become increasingly dependent on foreign oil. The

US supplied 60% of the total petroleum consumed in 1980; in 2008 it supplied 34% of its

total petroleum consumption.1 This leaves the US economy increasing susceptible to price

fluctuations in the global market. The 500% in the price of a barrel of oil from 2002 to 2008

had a drastic impact on the US economy.2 In recent years, the US government has legislated

a number reforms in an effort to establish energy security. The Energy Policy Act of 2005

established the first mandate for biofuels production in US history.3 In 2007, these mandates

were elaborated upon in the Energy Independence and Security Act.4 By 2008, 9 billion

gallons of biofuels were to be blended into existing fuel,4 and 26 billion gallons by 2022

(EPA Renewable Fuel Standard). US biofuel production has been increasing to meet the

legislated demand; however, problems exist with the current commercially produced

biofuels.

Ethanol and biodiesel are the most common commercially produced biofuels. Ethanol

is produced largely from the fermentation of carbohydrates; biodiesel is produced by the

transesterification of triglycerides. Though ethanol and biodiesel have both been successfully

commercially implemented, they must be mixed with petroleum-based transportation fuels to

be used in conventional engines. Both ethanol and biodiesel contain oxygen. Biodiesel is a

fatty acid methyl ester (FAME), and ethanol is an alcohol. Since they are partially oxidized,

their energy density is less than that of conventional fuels. The oxygen also increases their

polarity, making both ethanol and biodiesel fairly hygroscopic. Proper storage techniques

must be implemented to prevent water condensation within the fuel. Moreover, ethanol is a

3

corrosive solvent, and can be destructive to metal and polymer engine components. Biodiesel

can be a food source for microbes which can in turn create a film which can clog engine

parts. A biocide must be used to prevent this adverse effect. Biodiesel also becomes viscous

at temperatures below 0 °C which limits its implementation in cold climates. Ultimately,

since ethanol and biodiesel are chemically dissimilar from traditional transportation fuels,

they have different properties and are unable to completely replace them.

With these limitations in mind, another process has been developed which produces

biofuels that are chemically identical to current transportation fuels. To produce second

generation biodiesel, fats and oils from biological sources are first hydrolyzed to liberate

fatty acids (FAs) from their glycerol backbone. FAs are then catalytically deoxygenated to

produce linear alkanes. These alkanes can then be isomerized and cracked to obtain the right

distribution of compounds required for gasoline, jetfuel, and diesel fuel applications. The

resultant biomass-derived transportation fuels are chemically identical to the fuels currently

in use. With advent of algal oils technology, high quantities of FAs can be produced using

minimal farmland to provide biologically derived transportation fuels for the future.5

FA deoxygenation occurs via two pathways on the catalyst surface which are

displayed below for stearic acid.6

361723517 HCnCOCOOHHC −+→ (1)

The products of decarboxylation are CO2 and a linear alkane which is one carbon less than

the reacted FA. In this case, n-heptadecane, a linear alkane is produced from SA. No net H2

4

is consumed during decarboxylation. The decarbonylation pathway is displayed below for

SA:

34173517 HCCOCOOHHC +→ (2)

Decarbonylation yields, CO, H2O, and an alkene which is one carbon less than the reacted

FA. In this case, 1-heptadecene is produced from SA. In order to produce a linear alkane, the

alkene must be hydrogenated, resulting in a mole of hydrogen consumed per mole linear

alkane produced. The decarboxylation pathway is the preferred pathway because it neither

consumes H2 nor produces CO, which acts as a catalyst poison and inhibits FA

deoxygenation.

This research is devoted to the development of the catalytic deoxygenation of FAs for

the production of second generation biofuels. Although a basic understanding of this reaction

has been obtained, further research is needed to take the current base of knowledge from

model compounds to being able to effectively utilize biologically derived FAs as feedstocks.

The current knowledge in literature regarding FA deoxygenation is summarized with an

emphasis on the action that must be taken to take this knowledge from practicing solely on

model compounds to directly deoxygenating biologically derived sources. This provides a

background for the need and application of the knowledge gained by the research conducted

for chapters 2 and 3 in this thesis.

Maier et al. studied the decarboxylation of carboxylic acids in the vapor phase over

Ni/SiO2 and Pd/SiO2 catalysts.7 The reaction conditions were 330 °C for Pd and

5

180 °C for Ni; both catalysts were operated under 25 ml/min H2 at atmospheric pressure.

Decarboxylation of heptanoic and octanoic acids over Pd/SiO2 resulted in high yields of

hexane and heptane (98% and 97% respectively), but much lower yields were obtained over

Ni (26% and 64% respectively). No catalytic activity was observed under Argon flow. Maier

et al. concluded that H2 was necessary for the decarboxylation reaction to proceed and that

dissociated H2 on the Pd surface must play an important role in the reaction mechanism.

Murzin et al. has published several papers on the production of second generation

fuel precursors by the catalytic deoxygenation of FAs. Snare, a member of Murzin’s group,

published a catalyst screening paper in 2006.8 Several transition metal catalysts were tested

for SA deoxygenation in a 300-ml stirred autoclave. The reaction conditions were 6 bar, 300

°C, 86 g dodecane, 4.5 g SA (0.154 mol/L), under 25 ml/min flowing He for 6 h. A variety of

transition metals (Ni, NiMo, Ru, Pd, PdPt, Pt, Ir, Os, and Rh) supported on various metal

oxides and carbon were screened for deoxygenation activity. The catalysts were reduced

under conditions appropriate for each metal. Additionally, Pd/C was tested at 1, 5, and 10

wt.% metal loadings. The conversion and selectivity was determined by GC-FID analysis of

the reactor contents after the reaction. Out of all the catalysts tested, 5 wt.% Pd/C was the

most successful catalyst displaying 100% conversion and 95% selectivity toward n-

heptadecane. The total selectivity toward C17 products was 99%. The 1 wt.% Pd/C and 10

wt.% Pd/C had 33.4% and 48.1% conversion and 52% and 60% selectivity to n-heptadecane

product respectively. It has later been discovered that the nature of the support, nanoparticle

6

distribution throughout the catalyst support, and dispersion have an effect on catalyst

activity.9

The deoxygenation of molecules originating from triglycerides, and moieties from

biologically derived fats and oils, have also been tested over Pd/C.10 Kubickova tested the

deoxygenation of SA, ethyl stearate, and tristearine over 5 wt.% Pd/C under He, 5% H2, and

100% H2 purge gases in a 300-ml stirred autoclave. Liquid phase sampling was used to gain

kinetic data for these compounds. Sampling was performed throughout the reaction. The 5%

H2 flow resulted in the highest turnover frequency (TOF) as well as the highest selectivity

toward n-heptadecane, 126 x 103 s-1 and 62% conversion after 360 min respectively.

However, the liquid-phase sampling time resolution was insufficient for accurately

determining kinetic parameters. The main product of all the reactions under 5% H2 was n-

heptadecane

Snare et al. conducted semi-batch and fed-batch work on oleic acid and linoleic

deoxygenation over 5 wt.% commercial Pd/C.11 It was found that hydrogenation of oleic acid

proceeds first and is then followed by FA deoxygenation. Under an inert environment, it was

found that the oleic acid deoxygenation activity was very low and polyunsaturated products

were formed by dehydrogenation. Moreover, H2 uptake was not monitored throughout the

course of these experiments. Maintaining a low H2 consumption per mole alkane formed is

essential for the industrial application of this process.

The structure sensitivity of the reaction was tested by Simakova by synthesizing four

1 wt.% Pd/C catalysts of various dispersions by precipitation deposition of Pd chloride

7

solutions over a pH range (8-10).12 The dispersion of the catalysts was determined by TEM

and CO chemisorption to be 18% 47%, 65%, and 72%. A mixture of palmitic and stearic acid

(60-40%) was deoxygenated under 25 ml/min 5% H2(He) at 17.5 bar. The decarboxylation

reaction was found to be structure sensitive. The catalysts with dispersions of 47% and 65%

were the most productive catalysts. The initial rate of deoxygenation was highest for 65%.

Also, 65% dispersion showed the highest turnover frequency. The kinetic results are

displayed below in Table 1. These values are based solely on liquid phase analysis of the

products. The sampling frequency was quite low. Online QMS analysis of the effluent stream

would provide a much more detailed analysis of the deoxygenation reaction kinetics by

increasing the sampling frequency from once an hour to multiple times per minute if

necessary.

Table 1: Kinetic results from Simakova’s structure sensitivity study.

Metal Dispersion Initial rate (mmol/(min·gcat))

TOF (s-1) Conversion after 300 min

18% 0.03 30 68

47% 0.2 76 100

65% 0.4 109 99

72% 0.05 12 96

A range of FAs were deoxygenated over 1 wt.% Pd/C (sibunit).13 The dispersion of

the catalyst was 38% as determined by TEM and pulse CO chemisorption. Heptadecanoic

acid (C17:0), stearic acid (C18:0), nonadecanoic acid (C19:0), arachidic acid (C20:0), and

behenic acid (C22:0) were tested under 17 bar argon; the gas flow rate was not specified. The

purities of heptadecanoic acid and behenic acid were 90% and 80% respectively, and high

8

levels of phosphorous (226 ppm and ~150 ppm respectively) were found in the compounds

by atomic absorption. From these experiments, Simakova inferred that there was no effect of

chain length on the deoxygenation rate. However, the time scale and frequency of liquid

sampling were not sufficient to accurately measure the turnover frequency or rate of alkane

production for these FAs. Moreover, the data shows arachidic acid reaching completion 250

min before stearic acid. If sampling were more frequent, more accurate kinetic data could be

obtained allowing the turnover frequency and rate of alkane production to be more accurately

determined. Ultimately, a FA screening of high purity FAs under 5% H2 complete with

online QMS analysis of the reactor effluent stream would provide the ability to accurately

determine the effect of chain length on reaction kinetics.

Outside the work of Murzin’s group, Immer and Lamb also performed research on the

catalytic deoxygenation of FAs. Their work relied heavily upon analyzing the reactor effluent

stream by online QMS rather than just relying upon liquid phase sampling. The result is

much more frequent data and more accurate determination of kinetic parameters.

A catalyst screening similar to Murzin’s work was conducted for semi-batch stearic

acid deoxygenation over four distinct 5 wt.% Pd/C catalysts.9 The reaction took place in a

50-ml stirred autoclave reactor under He purge gas (60 ml/min) in dodecane. A uniformly

impregnated catalyst (E117), gave 100% conversion in 1 h with 99% selectivity toward

decarboxylation products. The same catalysts were tested under 60 ml/min 5% H2(He). The

uniformly impregnated catalyst still had the best performance under 5% H2 with the highest

9

CO2 selectivity, highest decarboxylation rate, and lowest H2 consumption. All of the catalysts

tested showed greater stability under 5% H2 than He.

E117 was used for the majority of Immer’s work. The work was performed by

analyzing the reactor effluent stream by QMS. The reactor contents were also analyzed by

GC-FID post reaction. Over E117, stearic acid deoxygenation under He occurred very

quickly with high CO2 selectivity. However, when the catalyst was reused, an order of

magnitude of activity was lost for the decarboxylation pathway and the CO selectivity greatly

increased. Fresh catalyst was limited to ~220 turnovers under He, but under 5% H2 over

2200 turnovers are possible.

Parametric studies were conducted for the semi-batch deoxygenation of SA. The

following parametric effects were observed. As SA concentration increases, selectivity

toward the decarboxylation pathway and the overall reaction rate decrease. As H2 and partial

pressure increases, the selectivity toward decarboxylation decreases. The presence of CO

from the decarbonylation pathway may inhibit the decarboxylation pathway while having no

effect on the decarbonylation pathway. CO and H2 showed additive inhibition. The

decarbonylation reaction is not affected by the parameters that inhibit the decarboxylation

reaction. Immer infers that decarboxylation and decarbonylation occur on distinct sites.

The fed-batch deoxygenation of stearic acid was studied in a 50-ml stirred autoclave

reactor with continuous feeding throughout the reaction for up to 24 h.14 Higher H2 partial

pressures resulted in dramatic switchover from high CO2 selectivities to high CO

selectivities. H2 consumption increased drastically with increasing CO selectivity. The time

10

of switchover increased as the hydrogen partial pressure decreased. Immer proposed that the

decarboxylation pathway was inhibited by the high H2 partial pressures and led to a build up

of stearic acid in the reactor. The higher FA to catalyst ratio resulted in increased

decarbonylation activity which ultimately led to CO poisoning of the catalyst. Reaction

inhibition due to high CO and H2 partial pressures as well as high FA to catalyst ratios is

reversible with lowering CO or H2 partial pressure or stopping stearic acid injection.

However, after 10 h of reaction under conditions favoring decarbonylation, the reaction

inhibition was irreversible.

Though much has been accomplished in understanding the catalytic deoxygenation of

FAs, there is still a strong need for continued research in the area. Neither Murzin nor Immer

have directly studied biologically derived FAs in their research. Catalyst deactivation due to

sulfur and phosphorous in biologically derived FAs is a major concern since biological

sources often contain phospholipids and sulfolipids. This issue must be addressed before

industrial production of second generation biodiesel can be implemented. Further work is

needed to adapt the catalytic FA deoxygenation to cyanobacteria-derived FAs. Cyanobacteria

provide the largest yield per acre of any triglyceride source.5 Cyanobacteria produce FAs

with a range of carbon number from C12 to C20 with varying unsaturations.15 Simakova’s

work was insufficient to determine the effect of FA carbon number on decarboxylation rate

over this range. Given that physical parameters vary considerably with carbon number (i.e.

volatility and heat of adsorption) there is a need to conduct an FA screening study to

determine the effect of FA carbon number on decarboxylation rate.

11

References

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2. World Crude Oil Prices. Administration, US Energy Information Administration

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134-139. 10. Kubickova, I.; Snare, M.; Eranen, K.; Maki-Arvela, P.; Murzin, D. Y., Catalysis

Today 2005, 106 (1-4), 197-200. 11. Snare, M.; Kubickova, I.; Maki-Arvela, P.; Chichova, D.; Eranen, K.; Murzin, D. Y.,

Fuel 2008, 87 (6), 933-945. 12. Simakova, I.; Simakova, O.; Maki-Arvela, P.; Simakov, A.; Estrada, M.; Murzin, D.

Y., Applied Catalysis a-General 2009, 355 (1-2), 100-108. 13. Simakova, I.; Simakova, O.; Maki-Arvela, P.; Murzin, D. Y., Catalysis Today 2010,

150 (1-2), 28-31. 14. Immer, J. G.; Lamb, H. H., Energy & Fuels 2010, 24, 5291-5299. 15. Thomas WH, T. T., Weissman J. Screening for Lipid Yielding Microalgae: Activities

for 1983; 1984

12

A Comparison of Supported Pd Catalysts for Liquid-Phase

Deoxygenation of Fatty Acids

Jeffrey P. Ford, Jeremy G. Immer, and H. Henry Lamb*

Department of Chemical and Biomolecular Engineering, North Carolina State

University, Raleigh, NC 27605-7905, USA

*Corresponding author. Email: [email protected]

13

Abstract

A series of supported Pd catalysts was screened for liquid-phase deoxygenation of stearic,

lauric and capric (decanoic) acids under 5% H2 at 300 °C and 15 atm. On-line quadrupole

mass spectrometry (QMS) was used to measure conversion, CO2 selectivity, H2 consumption,

and initial decarboxylation rate. Post-reaction analysis of liquid products by gas

chromatograph (GC) was used to determine n-alkane yields. Pd-on-carbon (Pd/C) catalysts

were found to be highly active and selective for stearic acid (SA) decarboxylation under the

test conditions. In contrast, SA deoxygenation over Pd/SiO2 catalysts occurred primarily via

decarbonylation and at a much inferior rate. Pd/Al2O3 exhibited high initial SA

decarboxylation activity but deactivated under the test conditions. Similar CO2 selectivity

patterns among the catalysts were observed for deoxygenation of lauric and capric acid;

however, the initial decarboxylation rates tended to be lower for Pd/C and Pd/Al2O3 with

these substrates. The most active Pd/C catalyst was used to investigate the influence of alkyl

chain length on deoxygenation kinetics for C8-C18 fatty acids (FAs). Generally, as FA

carbon number decreases, reaction time and H2 consumption increase, and CO2 selectivity

and initial decarboxylation rate decrease. The increase in initial decarboxylation rates for

longer chain FAs is attributed to their greater propensities to adsorb onto the activated carbon

support.

14

1. Introduction

US petroleum production peaked in the early 1970’s. Afterward, the US economy has

been increasingly dependent on foreign oil to maintain economic growth. The oil embargo of

1972, the Iranian revolution of 1978, and the Iraqi invasion of Kuwait in 1990 are examples

of the effect that the foreign petroleum market has on the US economy. Although predictions

for world peak oil production vary dramatically, this much is certain: petroleum is a finite

resource, and world demand for petroleum is increasing. In 2002, the average price for a

barrel of oil in the US was $22.86.1 The summer of 2008, the average price per barrel of oil

in the US increased to $121.69.1 From 2002 to 2008, world petroleum consumption increased

by 9.5%.2 In China alone, petroleum consumption increased by more than 50%.2 As the

developing world continues to become more industrialized, the demand for petroleum will

only increase. Outside of environmental concerns and climate change, securing alternative

source of transportation fuel is necessary to protect the US economy from price volatility in

the world petroleum market.

As far as transportation fuel is concerned, biofuels provide a practical method to

reduce dependence on foreign oil. The most promising of all biofuels are those produced by

cyanobacteria, or microalgae. The accolades of microalgae are widespread. First,

cyanobacteria use CO2 as their carbon source to build all of their biomass, including

triglycerides. They have higher photosynthetic efficiency than most plants.3 They are able to

produce up to 80% dry weight basis triglyceride content,4 and values of 20-50% are easily

attainable.5 The highest accolade is that when considering 30% triglyceride content on a dry

15

weight basis, only 2.5% of the United States’ current agricultural land would be able to

supply 50% of the current transportation fuel need.5 Most notably, this avoids the problems

with first generation biofuels which require a large percentage of land use to. For example,

soybeans, the main source for biodiesel fatty acids, would require 326% of the currently

available US agricultural land when microalgae would be able to do so with 1.1-2.5% of the

currently available land (30 to 70% dry weight biomass).5 Moreover, they can be grown in

sea water which is non-potable.6 Domestic wastewater has been hypothesized as a source for

algae. There are thousands of species of microalgae, and therefore finding a species that

would be suited for a specific environment is very viable.7 Oil palm can only grow in very

specific environments which are not available to the whole country. Also, microalgae can be

grown in tanks on nonarable land, which further makes their application more viable.8

Depending on the species, microalgae synthesize a range of fatty acids. C12:0, C14:0,

C:16:0 and C18:0, C:20 with variations in the unsaturation of the compounds have all been

synthesized.9 Cyanobacteria can be genetically modified to selectively produce unsaturated

compounds. Moreover, Vermaas has developed a cyanobacteria which produces lauric acid

without attaching it to a glycerol backbone.10 This pathway has been promoted so that the

fatty acid is secreted into solution. This simplifies the separation process which has currently

been a hurdle in the synthesis of microalgal oils.

Up until this point, the majority of deoxygenation of fatty acids for the production of

second generation biofuels has focused on stearic acid (C18:0). C18 fatty acids are the major

constituent of soybean oil which is currently the major source of biodiesel fatty acids.

16

Although the alkyl chain does not play a direct role in the catalytic reaction, several physical

properties (i.e. volatility and heat of adsorption) vary with alkyl chain length. However,

lower chain length fatty acids have not yet been tested. This work seeks to develop an

understanding of the reaction kinetics of fatty acids of varying chain lengths over 5% Pd/C,

5% Pd/Al2O3, and 5% Pd/SiO2. Snare et al. screened transition metal catalysts for stearic

acid (SA) deoxygenation under an inert purge and found that 5 wt.% Pd/C gave the highest

conversion and selectivity to n-heptadecane.11 Unfortunately, rapid catalyst deactivation is

observed under inert atmosphere. Simakova et al. reported that the catalytic deoxygenation of

fatty acids using supported Pd catalysts is structure sensitive (i.e., metal dispersion affects

specific catalytic activity).14 Here, supported Pd catalysts were screened for SA

deoxygenation under 5% H2.

2. Experimental Methods

2.1. Materials

Dodecane (99+%) and dodecanoic acid (99%) were purchased from Sigma-Aldrich

and used as received. Stearic acid (97%), palmitic acid (98%), myristic acid (99%), and

decanoic acid (99%) were purchased from Acros Organics and used as received. Certified

5% CO (balance He), 20% CO2 (balance He), ultra-high-purity He and H2, and zero grade air

were obtained from National Welders.

Commercial catalysts were obtained from a number of sources. 5 wt% Pd/C catalysts

were purchased from Alfa-Aesar (AA38300) and provided by Evonik-Degussa (E117PB/W

E101NN/W, and E199NN/W). The Evonik-Degussa catalysts were powders with the

17

following median particle sizes: 35 µm for E117, 23 µm for E101, and 17 µm for E199. The

E199 and E101 catalysts had an eggshell Pd distribution where the Pd particles are primarily

on or near the surface of the activated carbon support. The Pd in E117 has a uniform

distribution where the Pd particles are distributed throughout the C support. A 5% Pd/SiO2

was purchased from BASF Strem Chemicals (BASF Escat 1351), and 5% PdAl2O3 was

purchased from Johnson Matthey (A302013-5). Carbon-supported catalysts were dried at

40°C overnight before use. In this work, the following nomenclature will be used for the

Pd/C catalysts: E117 – Pd/C (A), AA38300 – Pd/C (B), E101 – Pd/C (C), E199 – Pd/C (D).

A 5 wt.% Pd/SiO2 was synthesized by incipient wetness impregnation of Aerosil 300

SiO2 (Degussa, wetted, dried, and crushed with mortar and pestle to increase the bulk

density) with Pd(NO3)2·XH2O dissolved in deionized water. The impregnated SiO2 was

calcined at 400 °C under flowing zero grade air for 2 h. In this work, the resulting catalyst

will be referred to as 5% Pd/SiO2 (B), and the commercial 5% Pd/SiO2 will be referred to as

5% Pd/SiO2 (A).

2.2. FA deoxygenation experiments

Semi-batch FA deoxygenation experiments were conducted in a 50-ml stirred

autoclave (Autoclave Engineers). Gas flow rate and purge gas composition were set by mass

flow controllers (Brooks 5850E series). A 20-ml condenser was used to collect condensable

vapors from the reactor effluent. Downstream from the condenser, the reactor pressure was

controlled by a manual back pressure regulator (Tescom).

18

In a typical experiment, 22.5 g dodecane and 336 mg of catalyst were added to the

reactor. The catalyst was suspended in solvent by stirring at 240 rpm. The reactor system was

flushed with He for 5 min to remove air. The reactor system was then flushed with 30 ml/min

H2 for 5 min. The pressure was then increased to 2 atm. The catalyst was reduced in situ at

200 °C with a 5 °C/min ramp and a 1-h soak. The reactor was cooled to 30 °C before the

purge gas was switched back to He. The reactor was removed under He flow and

approximately 5.6 mmol FA were added to the reactor system manually. The reactor was

then sealed and purged under He flow for 5 min while agitating at 240 rpm. Afterward, the

reactor purge gas was switched to 60 ml/min 5% (H2/He) for 5 min, and then the pressure

was raised to 15 atm and the agitation rate was increased to 1000 rpm. The reactor was

heated to 300 °C at 5 °C/min and held at 300 °C for the reaction time (typically 4 h). The

reactor was cooled to 30 °C before liquid samples were collected. The condensate was

sampled after reaction for analysis.

2.3. Analytical methods

The reactor effluent was analyzed online using a QMS (Pfeiffer Prismaplus) with a

heated capillary inlet and Quadstar 32-bit software. The H2 (2 m/z), He (4 m/z), CO (28 m/z),

and CO2 (44 m/z) signals were measured routinely. The CO and CO2 signals were calibrated

using 20% CO2 in He and 5% CO in He. The CO signals have been corrected for CO2

ionization to CO+ (28 m/z) by subtracting 10% of the CO2+ (44 m/z) intensity.

Reactor and condensate samples were analyzed using an HP5890 gas chromatograph

(GC) equipped with a flame ionization detector (FID) and an Econocap EC-5 30 m x 0.32

19

mm x 1.0 µm capillary column. Chromatograms were collected using an SRI Model 333

Peak Simple Chromatography Data System. Peak integrations were performed using

PeakSimple Software. The following GC oven temperature program was used: 5 °C/min

ramp from 80 to 300 °C, 1 min soak at 300 °C. Samples (0.05 µL) were injected onto the

column inlet (300 °C, 10 psig head pressure) with a 50:1 split ratio. The n-alkane and FA

FID responses were calibrated using an n-decane internal standard. Response factors of n-

alkanes and FAs were determined relative to n-decane, but equivalent response factors were

assumed for the alkenes and alkanes of the same carbon number.

2.4. Catalyst characterization

The catalysts were characterized by pulsed CO chemisorption, N2 physisorption, and

H2 chemisorption. Pulsed CO chemisorption measurements were made on a custom

apparatus.12 Catalyst samples were pretreated under flowing 5% H2 for 1 h at 250 °C

followed by a 30 min He purge at 250 °C. The sample was then cooled to room temperature.

The uptake of CO pulses (500 µL sample loop, 5% CO) by the catalyst was monitored by on-

line QMS.

BET surface area, total pore volume, and micropore volume were also determined for

all of the catalysts by N2 porosimetry using a Micromeritics ASAP 2020c instrument.

Carbon-supported catalysts were degassed for 8 h at 300°C while non-carbon-supported

catalysts were degassed for at least 4 h. 40-point adsorption-desorption isotherms were

measured, and the data was used to determine the BET surface area, total pore volume at

97.7% saturation, and the t-Plot micropore volume.

20

2.5 Reactor Modeling – Initial Rates Calculation

In order to quantitatively compare the online QMS data, the initial decarboxylation

rate is calculated via a material balance. The CO2 mole balance below assumes that the

stirred autoclave contents (liquid and gas phases) are well-mixed; therefore, the effluent

composition reflects the instantaneous gas-phase composition in the reactor.

{ }

=+−

dt

dP

RT

VWrPP

RT

Q COg

COCOCO outin

2

2,2,2 (3)

Where rCO2 is the rate of CO2 generation, Q = purge rate, W = catalyst weight, Vg = volume

of reactor head space, R is the gas constant, and T is the absolute temperature. Since the CO2

partial pressure in the feed is negligible (PCO2,in ~ 0), Eq. 3 simplifies to

+=

dt

dP

Q

VP

WRT

Qr

COg

COCO2

22. (4)

The CO2 partial pressure is proportional to the He-normalized 44 m/z signal. Before the

reaction commences PCO2, out is approximately zero, and the initial rate is proportional to

dt

dPCO2 . The material balance simplifies to the following equation for initial rates:

=

dt

dP

WRT

Vr

COg

CO2

2 (5)

dt

dPCO2 is calculated by fitting a linear region of the PCO2 vs. t plot with a linear regression.

The linear fit extrapolates to the point where it crosses the t-axis, determining the initial

decarboxylation rate within the reactor.

21

3. Results and Discussion

3.1 Catalyst Characterization

The supported Pd catalysts were characterized by N2 porosimetry. The specific

surface area, total pore volume, and micropore volume of each catalyst are given in Table 1.

The Pd/Al2O3 and Pd/SiO2 catalysts have specific surface areas between 200-300 m2/g. The

specific surface areas of the Pd/C catalysts are much higher (780-850 m2/g) and typical of

activated carbon supported noble metal catalysts. The total pore volumes and surface areas of

the Pd/SiO2 catalysts are indicative of mesoporosity ( average cylindrical pore diameter of ~8

nm). The Pd/Al2O3 catalyst has a lower pore volume, and an average cylindrical pore

diameter (~2 nm) at the lower end of the mesoporous range (2-50 nm). In contrast, Pd/C

catalysts are microporous with 28-34% of their pore volume in micropores. The micropore

volumes of the Pd/C catalysts are equivalent; however, the total pore volume of Pd/C (A) is

~15% less than the others. Consistent with their total pore volumes and specific surface

areas, the Pd/SiO2 and Pd/Al2O3 catalysts have negligible micropore volumes.

Pd dispersions (% metal exposed) were determined by pulse CO chemisorption; the

reported values (Table 1) assume a 2 to 1 Pd surface atom to adsorbed CO stoichiometry

(Pds:CO 2:1). The commercial Pd/SiO2 (A) catalyst has a very low dispersion consistent with

an average Pd particle size >10 nm. Pd/SiO2 (B), which was prepared in-house by incipient

wetness impregnation, exhibits a higher dispersion consistent with an average Pd particle size

of ~5 nm. Significant for comparing their catalytic performance, Pd/SiO2 (B) and Pd/C (A)

have closely similar dispersions. The Pd/Al2O3 catalyst has a dispersion of ~50% consistent

22

with an average Pd particle size of ~2 nm. The dispersions of Pd/C (B) and Pd/C (D) are

similar to that of the Pd/Al2O3 catalyst; however, the latter was prepared with an “eggshell”

Pd distribution. The dispersion of Pd/C (C) (also prepared with an eggshell Pd distribution)

is intermediate between those of Pd/C (A) and Pd/C (D).

3.2 Stearic Acid Deoxygenation under 5 and 10% H2

Catalyst performance data for semi-batch SA deoxygenation at 300 °C under a 5% H2

purge are summarized in Table 2. SA conversions approaching 100% after 4 h (as quantified

by on-line QMS) are seen for all the supported Pd catalysts evaluated. The n-heptadecane

yields from GC-FID are essentially 100% under these conditions. Since unsaturated products

were not detected by post-reaction GC-FID analysis, we infer that any alkenes formed during

the reaction were hydrogenated to n-heptadecane. CO2 selectivity and H2 consumption

exhibit a marked dependence on the catalyst support. Pd/C catalysts have CO2 selectivites of

90-95% and correspondingly low values of H2 consumption per mole of SA converted,

consistent with SA decarboxylation: 236173517 COHCCOOHHC +→− . In contrast, the

Pd/SiO2 catalysts exhibit much lower CO2 selectivites and greater H2 consumption consistent

with decarbonylation followed by heptadecene hydrogenation:

OH COCOOHHC 234173517 ++→− HC . The Pd/SiO2 catalyst with the higher Pd

dispersion shows greater CO2 selectivity. Although Pd/C (A) and Pd/SiO2 (B) have similar

Pd dispersions, they exhibit very different catalytic performance. Deoxygenation proceeds

mainly via decarboxylation over Pd/C (A) and mainly via decarbonylation over Pd/SiO2 (B).

23

The CO2 selectivity of the Pd/Al2O3 catalyst is intermediate between those of the Pd/C and

Pd/SiO2 catalysts.

CO2 selectivity decreases and H2 consumption increases markedly when the H2

percentage in the reactor purge stream is increased from 5 to 10% consistent with H2

inhibition of the decarboxylation pathway.13 This effect is particularly strong for the Pd/SiO2

catalysts (Table 2). Under a 10% H2 purge, Pd/SiO2 (A) is greater than 97% selective for

decarbonylation. CO2 selectivity and H2 consumption per mole of SA are strongly correlated

for all the catalysts tested for SA deoxygenation under 5% and 10% H2, as evidenced in

Figure 1. The negative slope of the regression line corresponds to 1.06 moles of H2

consumed per mole of CO produced. There are two potential interpretations of this

stoichiometry: conversion of CO2 to CO via the reverse water-gas shift reaction and

hydrogenation of the heptadecenes produced by SA decarbonylation. Evidence supporting

the latter interpretation is presented below.

On-line QMS data (expressed as molar flow rates) collected during liquid-phase SA

deoxygenation using the Pd/SiO2 (A), Pd/SiO2 (B), Pd/Al2O3 and Pd/C (A) catalysts are

compared in Figure 2. Note that the ordinate range in Figure 2a is an order of magnitude

larger than in Figure 2b. In addition to its very high CO2 selectivity, Pd/C (A) is much more

active for SA deoxygenation than the Pd/SiO2 catalysts. The reaction goes to completion in

~30 min with the Pd/C (A) catalyst; whereas, completion requires ~80 min with the Pd/SiO2

catalysts, and there is substantially greater decarbonylation activity. The Pd/Al2O3 catalyst

exhibits high initial decarboxylation activity with CO2 evolution beginning before the reactor

24

has reached operating temperature; however, CO2 production peaks early, and CO and CO2

evolution continue until the reactor is cooled. We infer that it is because of this deactivation

that the Pd/Al2O3 catalyst fails to achieve complete SA conversion in 4 h. The measured H2

consumption rate (defined as the difference between the inlet and outlet H2 flow rates) is

plotted on the secondary ordinate axis in Figure 2b. H2 evolution is observed coincident with

the initial decarboxylation activity of Pd/C (A) and Pd/Al2O3. H2 consumption occurs

subsequently on the trailing edge of the CO2 evolution peak. These temporal H2 features are

prominent for SA deoxygenation over Pd/C (A); however, net H2 consumption is small

(Table 2). When Pd/SiO2 catalysts are employed for SA deoxygenation, the initial H2

evolution feature is negligible. H2 consumption begins after CO production peaks, and H2

consumption reaches its maximum value after CO and CO2 production are nearing

completion. We infer that the observed H2 consumption arises from secondary hydrogenation

of heptadecenes to n-heptadecane.

Since all the catalysts achieved high SA conversions after 4 h under reaction

conditions and some (e.g., Pd/Al2O3) showed signs of deactivation, the on-line QMS data

were used to estimate initial decarboxylation rates. The relative initial decarboxylation rates

for each catalyst under 5% H2 are given in Table 2; the rates are normalized to Pd/C (A)

which exhibited the highest absolute rate at 300 °C. The initial decarboxylation rate of

Pd/Al2O3 at 300 °C is probably underestimated since the catalyst becomes active at a lower

temperature and there is evidence of catalyst deactivation. Not surprisingly, catalysts with

higher initial decarboxylation rates tend to have higher overall CO2 selectivities.

25

Because Pd/C (A) was the most active SA dexoygenation catalyst under 5% H2,

several other 5 wt.% Pd/C catalysts were investigated, including some with higher Pd

dispersions and some with “eggshell” Pd distributions. The CO2 and CO production rates,

and H2 consumption rates measured by on-line QMS during SA dexoygenation using the

Pd/C catalysts are compared in Figure 3. Pd/C (A) produces a higher CO2 maximum flow

rate and reaches completion ~0.5 h earlier than the other Pd/C catalysts. Decarbonylation is

suppressed over Pd/C (A). Moreover, Pd/C (A) is the only carbon-supported catalyst to

evolve a significant amount of H2 initially; it is also the only Pd/C catalyst to exhibit a

narrow H2 consumption peak after SA decarboxylation nears completion. Pd/C (B), Pd/C

(C), and Pd/C (D) display similar decarboxylation activities; however, Pd/C (B) exhibits

higher decarbonylation activity. Although the CO2 selectivities, n-heptadecane yields, and

SA conversions are comparable for the Pd/C catalysts, the relative initial decarboxylation

rate is significantly higher for Pd/C (A) than the other Pd/C catalysts (Table 2). Since Pd/C

(A) has less than half the available active sites when compared to Pd/C (B) and Pd/C (C), its

advantage in catalytic activity per active site (turnover frequency) is even greater than

indicated by the relative rates. Simakova et al. suggested the ideal dispersion to be ~50%,

which is much closer to that of Pd/C (B) and Pd/C (D) than Pd/C (A). Adsorption of fatty

acids on activated carbons also is known to be affected by the specific surface functional

groups on the carbon,15 and the nature and densities of surface groups are determined by the

carbon source and activation procedure. Further research is needed to determine how Pd

26

dispersion, carbon support characteristics, and other factors contribute to the superior

performance of Pd/C (A) in SA deoxygenation catalysis.

3.3 Lauric acid deoxygenation using Pd/SiO2 (A), Pd/Al2O3 and Pd/C (A)

Semi-batch lauric acid (LA) deoxygenation at 300 °C under a 5% H2 purge was

investigated using Pd/SiO2 (A), Pd/Al2O3, and Pd/C (A). As shown in Table 3, these

catalysts gave LA conversions and n-undecane yields above 90%; however, the CO2

selectivities are very different. LA deoxygenation over Pd/SiO2 (A) proceeds selectively via

the decarbonylation pathway; whereas, the CO2 producing decarboxylation pathway is

dominant for LA deoxygenation when using Pd/Al2O3 and Pd/C (A).

As evidenced in Table 2 and Figure 4, LA deoxygenation over these catalysts is

significantly slower than SA deoxygenation over the same catalysts under equivalent

conditions (Figure 2); however, the trends in catalytic performance are similar. Complete

LA conversion is achieved with Pd/C (A) after ~1 h under reaction conditions. In contrast,

LA deoxygenation requires nearly 3 h for completion over Pd/SiO2 (A). The initial

decarboxylation rate is higher for LA over Pd/Al2O3 than over Pd/C (A) (Table 3); however,

CO2 production with Pd/Al2O3 tails off as the reaction nears completion, suggesting catalyst

deactivation. The initial LA decarboxylation rate over Pd/C (A) is nearly an order of

magnitude slower than SA decarboxylation over the same catalyst. Overall, CO2 selectivity

is higher for Pd/Al2O3 than Pd/C (A), because the decarbonylation pathway is more active

over Pd/C (A). Small H2 evolution peaks are observed at the beginning of the LA

deoxygenation reaction over Pd/Al2O3 and Pd/C (A). These peaks are followed by H2

27

consumption troughs on the trailing edges of the CO2 evolution peaks, as also seen in Figure

2. The H2 consumption trough is much broader for Pd/Al2O3 extending virtually to the end

of the reaction time. Initial H2 evolution is minimal over Pd/SiO2 (A), and a broad H2

consumption feature is observed concurrent with CO production.

3.4 Decanoic acid deoxygenation over Pd/SiO2 (A), Pd/Al2O3 and Pd/C (A)

Decanoic acid (DA) deoxygenation also was investigated over Pd/SiO2 (A), Pd/Al2O3

and Pd/C (A), and catalyst performance trends similar to those described previously for SA

and LA deoxygenation were observed. Pd/C (A) and Pd/Al2O3 are much more selective for

DA decarboxylation than Pd/SiO2, and consequently, H2 consumption per mole of DA

converted is much higher over Pd/SiO2.

The initial rates of decarboxylation are lower and the completion times are longer for

DA deoxygenation than for LA, reinforcing the trend of declining reactivity with decreasing

FA carbon number. The initial DA decarboxylation rate for Pd/Al2O3 is higher than for Pd/C

(A); however, for DA deoxygenation (in contrast to SA and LA) significant tailing is

observed for CO2 production over Pd/C (A)—extending the completion time to almost 3 h

under reaction conditions. H2 consumption values over Pd/C (A) and Pd/Al2O3 was similar.

Both catalysts exhibit initial H2 evolution peaks and subsequent H2 consumption troughs.

These features are less prominent than those observed during SA and LA deoxygenation over

Pd/C (A) and Pd/Al2O3. Pd/SiO2 (A) exhibits negligible decarboxylation activity throughout

the reaction. The CO production pathway over Pd/SiO2 (A) increases sharply as the reactor

reaches temperature and falls as the DA concentration decreases within the reactor. DA

28

deoxygenation over Pd/SiO2 (A) also nears completion after ~3 h at reaction conditions. H2

consumption occurs concurrent with CO production throughout the reaction; however, peak

H2 consumption occurs approximately 1 h after peak CO production.

As shown in Table 3, the n-nonane yield measured by GC is lower than DA

conversion in each of the experimental runs. This deficit in the material balance is attributed

primarily to n-nonane loss through the condenser. Because of the higher vapor pressure of n-

nonane (than the other hydrocarbon products), it is more difficult to condense leading to

evaporative losses. DA is also more volatile than LA and SA and could be transported from

the reactor with the purge stream resulting in lower than expected DA conversions (based on

CO and CO2 evolution).

3.5 Dependence of decarboxylation rate on fatty acid carbon number

From the SA, LA, and DA deoxygenation experiments over Pd/C (A) and Pd/Al2O3,

trends of lower initial rates, longer completion times, and lower CO2 selectivities with

decreasing FA carbon number were observed. To gain further insight into these trends, semi-

batch deoxygenation experiments using the Pd/C (A) catalyst were performed under

equivalent conditions for all naturally occurring FAs with carbon numbers from 8 to 18. The

rates of CO2 and CO production and H2 consumption for these experiments are compared in

Figure 6. As the FA carbon number decreases, the maximum CO2 production rate decreases,

and the peak shifts toward longer reaction times. Consequently, the CO2 peak broadens, and

the reaction takes longer to reach completion. Moreover, as the FA carbon number decreases,

so does its initial decarboxylation rate. The trend of initial decarboxylation rate with FA

29

carbon number is shown in Figure 7. SA also exhibits the largest and sharpest initial H2

evolution feature. As FA carbon number decreases, this initial H2 evolution feature and

subsequent H2 consumption trough become broader and more diffuse. The initial CO

production (decarbonylation) rates are similar for all FAs studied; however, the

decarbonylation pathway remains active longer for lower molecular weight FAs. This effect

is due to the increased reaction time required for FAs with shorter alkyl chains.

The overall CO2 selectivities (Table 4) reflect the effect of the lower initial

decarboxylation rates and extended reaction times for lower molecular weight FAs. As FA

carbon number decreases, CO2 selectivity decreases and H2 consumption increases. The

linear correlation between H2 consumption per mole of FA converted and CO2 selectivity

(Figure 1) appears to be universal for FA deoxygenation over supported Pd catalysts. For

each FA, the only hydrocarbon product detected by GC-FID analysis was the n-alkane

corresponding to loss of the carboxyl group as CO2. Yields of the n-alkanes are greater than

95% for C14-C18 FAs. Alkane yields are lower for the lighter members of the series due to

the increasing volatilities of the FAs and their deoxygenation products. Octanoic acid, the

lightest member of the series, is a liquid at room temperature and has the highest volatility.

The low conversion and n-heptane yield for octanoic acid are attributed to evaporative losses

of the reactant (from the reactor) and n-heptane from the condenser.

Previous workers have concluded that there is not a significant effect of fatty acid

alkyl chain length on deoxygenation kinetics over Pd/C catalysts. Lestari, et al. investigated

SA and PA deoxygenation over a 4% Pd/C in the liquid phase and concluded that chain

30

length had no impact on deoxygenation kinetics.16 Simakova et al. determined the kinetic

parameters of fatty acids in the range between C17:0 to C22:0 by liquid phase sampling.17

Slight differences in deoxygenation activity were attributed to catalyst deactivation from

feedstock impurities. These kinetic studies obtained data by liquid sampling of the reactor.

Because long-chain FAs react very rapidly over Pd/C, only a few data points were collected

at low to moderate conversion; whereas, we have the advantage of obtaining essentially

continuous data via on-line monitoring of the reactor effluent. One potential explanation for

our contrasting conclusions is the range of FAs studied. It is feasible that alkyl chain length

effects are less prominent for FAs with carbon numbers of 16 and greater—as suggested by

the leveling off in the initial decarboxylation rate between SA and PA in Figure 7.

We attribute the increasing initial decarboxylation rate with fatty acid carbon number

to a greater propensity for adsorption on the activated carbon. Kipling and Wright studied the

adsorption of OA, LA, PA, and SA onto carbon blacks from cyclohexane.15 The higher the

carbon number of the acid, the more strongly it adsorbed at low concentrations. The

adsorption of FAs on carbon supports depends on the functional groups found on the

activated carbon surface.15

The differences in FA decarboxylation activity among the Pd/C, Pd/Al2O3 and

Pd/SiO2 catalysts also can be explained by a greater tendency of the FA to adsorb onto the

activated carbon and alumina supports. According to Traube’s rule, adsorption of long-chain

fatty acids on activated carbons is more favorable than on silica.18 SiO2 is ionic/covalent in

nature and has a slightly acidic isolectric point; whereas, activated carbons are comprised of

31

folded and disordered graphite sheets with slit-like pores. Activated carbon surfaces may

expose acidic and basic surface groups depending on the carbon source and activation

procedures.19 FAs typically adsorb on activated carbons with the alkyl chain parallel to the

surface plane.20 Al2O3 (depending on the specific phase and impurities) may have a basic

isoelectric point, which contributes to a coulombic interaction between the FA and the Al2O3

surface.21 Instead of adsorbing parallel to the Al2O3 surface, FAs adsorb perpendicular to the

Al2O3 surface.22 Therefore, since the coulombic interaction between Al2O3 and FAs is not

affected by chain length, it may explain why Pd/Al2O3 catalyst is more active initially than

the Pd/SiO2 and Pd/C catalysts for LA and DA deoxygenation. FAs adsorb parallel to the

SiO2 surface.22 However, due the acidic nature of the SiO2 support, coulombic interaction

with FAs is not favorable. This may explain why the decarboxylation activity over Pd/SiO2 is

considerably less than the decarboxylation activity over Pd/C and Pd/Al2O3.

3.4 Deoxygenation of FA mixtures

Semi-batch deoxygenation of a 50-50 DA-SA molar mixture was investigated at

300 °C under a 5% H2 purge using Pd/C (A). DA-SA deoxygenation exhibits CO2 selectivity,

H2 consumption, conversion and n-alkane yield intermediate to SA and DA deoxygenation in

Table 5. CO2 production for the DA-SA mixture is significantly more active than DA

deoxygenation initially (Figure 8). This is consistent with the explanation that initial CO2

production is attributed to the increased propensity for SA to adsorb to the carbon support.

The DA-SA mixture reaches maximum production at the same time that the SA approaches

completion. After CO2 production reaches a maximum, the SA completes reaction after

32

30 min at reaction conditions; whereas the DA-SA mixture requires ~2 h. The initial H2

evolution peak and subsequent consumption trough observed for the DA-SA mixture are

intermediate between that of DA and SA deoxygenation. The subsequent H2 consumption

trough is more pronounced than that of DA deoxygenation as CO production decreases zero

at an earlier batch time. CO production for DA, DA-SA and SA reach the same value

initially. The DA-SA CO trace decreases more quickly than that of DA deoxygenation,

indicating that the concentration of FA in the reactor decreases more quickly for DA-SA

deoxygenation.

4. Conclusions

Pd/C catalysts were more active for SA deoxygenation than the Pd/Al2O3 and Pd/SiO2

catalysts. Pd/C (A) was the most active Pd/C catalyst for SA deoxygenation. FA carbon

number had a pronounced effect on decarboxylation activity over Pd/C (A). The initial

decarboxylation rate and CO2 selectivity decreased while H2 consumption and reaction time

increased with increasing FA carbon number. We attribute the increase in initial

decarboxylation rate increases with FA carbon number due to an increased propensity for the

FA to adsorb on the carbon support. Similar trends were also observed for LA and DA

deoxygenation over Pd/Al2O3 and Pd/SiO2 (A). Decarboxylation trends over Pd/Al2O3 and

Pd/SiO2 were also attributed to FA-support interaction. Further research is needed to

determine how Pd dispersion and support characteristics promote decarboxylation activity

over Pd/C (A).

33

References

1. World Crude Oil Prices. US Energy Information Administration, 2011.

2. World Petroleum Consumption, Annual Estimates, 1980-2008. US Energy Information Administration, 2009.

3. Minowa, T.; Yokoyama, S.; Kishimoto, M.; Okakura, T., Fuel 1995, 74 (12), 1735-1738.

4. Patil, V.; Tran, K. Q.; Giselrod, H. R., Int. J. Mol. Sci. 2008, 9 (7), 1188-1195.

5. Chisti, Y., Biotechnology Advances 2007, 25 (3), 294-306.

6. Grima, E. M.; Belarbi, E. H.; Fernandez, F. G. A.; Medina, A. R.; Chisti, Y.,

Biotechnology Advances 2003, 20 (7-8), 491-515.

7. Mata, T. M.; Martins, A. A.; Caetano, N. S., Renewable & Sustainable Energy Reviews

2010, 14 (1), 217-232.

8. Amin, S., Energy Conversion and Management 2009, 50 (7), 1834-1840.

9. Thomas WH, T. T., Weissman J. Screening for Lipid Yielding Microalgae: Activities for

1983; 1984. 10. ARPA-E Cyanobacteria Designed for Solar-Powered Highly Efficient Production of

Biofuels. http://arpae.energy.gov/LinkClick.aspx?fileticket=NER0Iui8UXc%3D&tabid=212 (accessed 3/25/2011).

11. Snare, M.; Kubickova, I.; Maki-Arvela, P.; Eranen, K.; Murzin, D. Y., Industrial &

Engineering Chemistry Research 2006, 45 (16), 5708-5715. 12. Kelly, M. J.; Kim, J.; Roberts, G. W.; Lamb, H. H., Topics in Catalysis 2008, 49 (3-4),

178-186.

13. Immer, J. G.; Lamb, H. H., Energy & Fuels 2010, 24, 5291-5299. 14. Simakova, I.; Simakova, O.; Maki-Arvela, P.; Simakov, A.; Estrada, M.; Murzin, D. Y.,

Applied Catalysis a-General 2009, 355 (1-2), 100-108.

15. Kipling J., Wright E., Journal of the Chemical Society 1963, 3382-3389.

34

16. Lestari, S.; Maki-Arvela, P.; Simakova, I.; Beltramini, J.; Lu, G. Q. M.; Murzin, D. Y., Catalysis Letters 2009, 130 (1-2), 48-51.

17. Simakova, I.; Simakova, O.; Maki-Arvela, P.; Murzin, D. Y., Catalysis Today 2010, 150

(1-2), 28-31.

18. Adamson, A. W., Physical Chemistry of Surfaces. 5 ed.; Wiley: New York, 1990.

19. Bansal R.C., G. M., Activated Carbon Adsorption. CRC Press: 2005.


21. Marmier, N., Metal Ion Adsorption on Silica, Alumina, and Related Surfaces. In Encyclopedia of Surface and Colloid Science: Inv-Pol, Hubbard, A. T., Ed. CRC Press: 2002.


35

Table 1: Catalyst characterization results.

Pd Dispersion

Surface Area (m2/g)

Micropore Volume (cm3/g)

Total Pore Volume (cm3/g)

Pd/SiO2 (A) 8.4% 273 0.015 1.05

Pd/SiO2 (B) 23.6% 218 0.002 1.19

Pd/Al2O3 47.7% 254 0.001 0.264

Pd/C (A) 19.5% 797 0.21 0.617

Pd/C (B) 50.5% 817 0.20 0.710

Pd/C (C) 38.4% 842 0.20 0.721

Pd/C (D) 47.6% 784 0.21 0.706

36

Table 2: Results for SA deoxygenation using 5 wt.% supported Pd catalysts.a

Catalyst % H2 Relative Initial

Decarboxylation Rateb X SCO2

H2 consumptionc

n-Heptadecane Yield

Pd/SiO2 (A) 5 0.09 1.020 0.493 0.540 0.99

Pd/SiO2 (B) 5 0.34 1.107 0.661 0.303 1.06

Pd/Al2O3 5 0.53 0.931 0.842 0.192 ----

Pd/C (A) 5 1.00 0.960 0.952 0.110 1.00

Pd/C (B) 5 0.62 1.011 0.908 0.125 1.11

Pd/C (C) 5 0.55 1.001 0.935 0.097 1.05

Pd/C (D) 5 0.71 0.992 0.936 0.098 1.03

Pd/SiO2 (A) 10 ----- 1.050 0.028 1.12 0.96

Pd/SiO2 (B) 10 ----- 1.117 0.173 0.836 1.03

Pd/C (A) 10 ----- 0.986 0.772 0.370 1.00

a) Reaction conditions: 0.0056 mol SA, 336 mg catalyst, 22.5 g dodecane solvent,

5% H2(He), 60 ml/min purge gas flow rate, 15 bar, 300 °C, and 4 h reaction time. b) Relative to SA deoxygenation over Pd/C (A) under 5% H2. c) Moles H2 consumed per mol SA converted.

37

Table 3: Results for LA and DA deoxygenation over Pd/SiO2 (A), Pd/Al2O3, and Pd/C (A).a

Catalyst FA Relative Initial

Decarboxylation Rateb X SCO2

H2 Consumptionc

n-Alkane Yield

Pd/SiO2 (A) LA 0.004 0.941 0.127 0.925 0.96

Pd/Al2O3 LA 0.23 0.918 0.929 0.158 0.94

Pd/C (A) LA 0.12 0.971 0.869 0.152 0.90

Pd/SiO2 (A) DA 0.001 0.918 0.040 0.840 0.87

Pd/Al2O3 DA 0.10 0.912 0.870 0.102 0.82

Pd/C (A) DA 0.05 0.948 0.842 0.182 0.88

a) Reaction conditions: 0.0056 mol FA, 336 mg catalyst, 22.5 g dodecane solvent,

5% H2(He), 60 ml/min purge flow rate, 15 bar, 300 °C, and 4 h reaction time. b) Normalized to SA deoxygenation over Pd/C (A) under 5% H2. c) Moles H2 consumed per mol FA converted.

38

Table 4: Results for deoxygenation of FAs with varying carbon numbers.a

FA Relative Initial

Decarboxylation Rateb X SCO2 H2 Consumptionc n-Alkane Yield

Stearic (C18:0) 1.00 1.01 0.946 0.029 0.96

Palmitic (C16:0) 0.73 0.98 0.935 0.113 1.00

Myristic (C14:0) 0.24 1.01 0.916 0.104 0.96

Lauric (C12:0) 0.12 0.97 0.869 0.141 0.90

Decanoic (C10:0) 0.047 0.95 0.842 0.182 0.88

Octanoic (C8:0) 0.029 0.76 0.719 0.208 0.47


5% H2(He), 60 ml/min purge flow rate, 15 bar, 300 °C, and 4 h reaction time. b) Normalized to SA deoxygenation over Pd/C (A) under 5% H2. c) Moles H2 consumed per moles FA converted.

39

Table 5: Results for the deoxygenation of SA, DA, and a 50-50 molar mixture of DA and SA.a

FA Relative Initial

Decarboxylation Rateb X SCO2 H2 Consumptionc n-AlkaneYield

SA 1.00 1.01 0.946 0.029 0.960

50-50 DA-SA 0.27 0.98 0.895 0.121 0.897

DA 0.047 0.95 0.842 0.182 0.884


5% H2(He), 60 ml/min purge flow rate, 15 bar, 300 °C, and 4 h reaction time. b) Normalized to SA deoxygenation over Pd/C (A) under 5% H2. c) Moles H2 consumed per moles FA converted.

40

0

0.2

0.4

0.6

0.8

1

1.2

0 20 40 60 80 100

SAOther FAs

y = 1.081 - 0.010569x R= 0.98936

Mol H2 Consumed per Mol FFA Converted

CO2 Selectivity

Figure 1: Correlation between H2 consumption and CO2 selectivity for SA and other FA deoxygenations over 5% Pd/C (A).

41

0

0.05

0.1

0.15

0.2

0.25

0.3

50

100

150

200

250

300

0 1 2 3 4 5

Pd/SiO2 (A)

Pd/SiO2 (B)

Pd/Al2O

3

Pd/C (A)

T

CO

2 Flow Rate (mmol/min)

T (

oC)

Time (h)

a

0

0.01

0.02

0.03

0.04

0.05

0.06 -0.05

0

0.05

0.1

0.15

0.2

0.250 1 2 3 4 5

Pd/SiO2 (A)

Pd/SiO2 (B)

Pd/Al2O

3

Pd/C (A)

CO Flow Rate (mmol/min)

H2 C

onsumption Rate (m

mol/m

in)

Time (h)

b

Figure 2: Molar production rate (mmol/min) of CO2 (a) and CO (b), and molar consumption rate of H2 (b), from SA deoxygenation over various 5% Pd catalysts in dodecane solvent at 300ºC for 4 h under 60 ml/min flowing 5% H2(He) at 15 atm.

42

0

0.05

0.1

0.15

0.2

0.25

0.3

50

100

150

200

250

300

0 1 2 3 4 5

Pd/C (A)

Pd/C (B)

Pd/C (C)

Pd/C (D)

T

CO


T (

oC)

Time (h)

a

0

0.005

0.01

0.015

0.02

0.025

0.03 -0.05

0

0.05

0.1

0.15

0.20 1 2 3 4 5

Pd/C (A)

Pd/C (B)

Pd/C (C)

Pd/C (D)


H2 C

onsumption Rate (m

mol/m

in)

Time (h)

b

Figure 3: Molar production rate (mmol/min) of CO2 (a) and CO (b), and molar consumption rate of H2 (b), from SA deoxygenation over various 5% Pd/C catalysts in dodecane solvent at 300ºC for 4 h under 60 ml/min flowing 5% H2(He) at 15 atm.

43

0

0.05

0.1

0.15

50

100

150

200

250

300

0 1 2 3 4 5

Pd/SiO2 (A)

Pd/Al2O

3

Pd/C (A)

T

CO


T (

oC)

Time (h)

a

0

0.02

0.04

0.06

0.08

0.1 -0.05

0

0.05

0.1

0.15

0.20 1 2 3 4 5

Pd/SiO2 (A)

Pd/Al2O

3

Pd/C (A)


H2 C

onsumption Rate (m

mol/m

in)

Time (h)

b

Figure 4: Molar production rate (mmol/min) of CO2 (a) and CO (b), and molar consumption rate of H2 (b), from LA deoxygenation over Pd/SiO2 (A), Pd/Al2O3, and Pd/C (A) in dodecane solvent at 300ºC for 4 h under 60 ml/min flowing 5% H2(He) at 15 atm.

44

0

0.02

0.04

0.06

0.08

0.1

50

100

150

200

250

300

0 1 2 3 4 5

Pd/SiO2 (A)

Pd/Al2O

3

Pd/C (A)

T

CO


T (

oC)

Time (h)

a

0

0.01

0.02

0.03

0.04

0.05

0.06 -0.04

0

0.04

0.08

0.12

0.16

0.20 1 2 3 4 5

Pd/SiO2 A

Pd/Al2O

3

Pd/C A


H2 C

onsumption Rate (m

mol/m

in)

Time (h)

b

Figure 5: Molar production rate (mmol/min) of CO2 (a) and CO (b), and molar consumption rate of H2 (b), from DA deoxygenation over Pd/SiO2 (A), Pd/Al2O3, and Pd/C (A) in dodecane solvent at 300ºC for 4 h under 60 ml/min flowing 5% H2(He) at 15 atm.

45

0

0.05

0.1

0.15

0.2

0.25

0.3

0.35

0.4

50

100

150

200

250

300

0 1 2 3 4 5

C18:0

C16:0

C14:0

C12:0

C10:0

C8:0

T

CO


T (

oC)

Time (h)

a

0

0.005

0.01

0.015

0.02

0.025 -0.05

0

0.05

0.1

0.15

0.20 1 2 3 4 5

C18:0

C16:0

C14:0

C12:0

C10:0

C8:0


H2 C

onsumption Rate (m

mol/m

in)

Time (h)

b

Figure 6: Molar production rate (mmol/min) of CO2 (a) and CO (b), and molar consumption rate of H2 (b), from FA deoxygenation over Pd/C (A) in dodecane solvent at 300ºC for 4 h under 60 ml/min flowing 5% H2(He) at 15 atm. 0.0056 mol of FA were added for each run.

46

0

0.2

0.4

0.6

0.8

1

0.7

0.75

0.8

0.85

0.9

0.95

6 8 10 12 14 16 18 20

Initial Rate

CO2

Relative Initial Decarboxylation Rate

CO

2 Selectivity

Chain Length

Figure 7: Correlation between relative initial decaroboxylation rate, CO2 selectivity and FA carbon number.

47

0

0.05

0.1

0.15

0.2

0.25

0.3

0.35

0.4

50

100

150

200

250

300

0 1 2 3 4 5

SA

50-50 DA-SA

DA

T

CO


T (

oC)

Time (h)

a

0

0.005

0.01

0.015

0.02

0.025 -0.05

0

0.05

0.1

0.15

0.20 1 2 3 4 5

SA

50-50 DA-SA

DA


H2 C

onsumption Rate (m

mol/m

in)

Time (h)

b

Figure 8: Molar production rate (mmol/min) of CO2 (a) and CO (b), and molar consumption rate of H2 (b), from the deoxygenation of SA, DA and a 50-50 molar mixture of DA and SA over Pd/C (A) in dodecane solvent at 300ºC for 4 h under 60 ml/min flowing 5% H2(He) at 15 atm. 0.0056 mol of FA were added for each run.

48

Semi-Batch Deoxygenation of Canola and Lard-Derived Fatty Acids

Jeffrey P. Ford and H. Henry Lamb*

Department of Chemical and Biomolecular Engineering, North Carolina State

University, Raleigh, NC 27605-7905, USA

*Corresponding author. Email: [email protected]

49

Abstract

Canola and lard-derived fatty acids (FAs) were deoxygenated at 300°C in the liquid phase

using a 5 wt.% Pd/C catalyst. On-line quadrupole mass spectrometry (QMS) was used to

monitor the effluent streams from the 50- and 600-ml stirred autoclave reactors. Stearic,

oleic, and palmitic acids were employed as model compounds. H2 consumption during oleic

acid (OA) deoxygenation at the 50-ml scale under 10% H2 occurred during heating to

reaction temperature consistent with double bond hydrogenation. The initial decarboxylation

rate of palmitic acid (PA) under 5% H2 decreased with increasing initial FA concentration in

dodecane; specific semi-batch deoxygenation productivity exhibited a maximum with PA

concentration. Canola-derived fatty acids (CDFAs) and a canola FA surrogate mixture also

were deoxygenated at the 50-ml scale. Decarboxylation was inhibited under 10% H2, and

there were indications of catalyst deactivation with the canola-derived fatty acids. Low CO2

selectivities and specific productivities were observed for model compounds and biomass-

derived fatty acids at the 600-ml scale due to the high initial FA concentrations employed.

Complete deoxygenation required substantially longer for OA than stearic acid. When the

on-line QMS traces for deoxygenation of OA and canola-derived FAs were superimposed,

there was no indication of catalyst deactivation attributable to impurities in the latter.

However, when the CDFA deoxygenation product was used as solvent in a subsequent run,

the decarboxylation pathway was inhibited.

50

1. Introduction

Concerns about global warming, rising petroleum prices, and energy security have

increased the interest in renewable energy over the last decade. Biodiesel and ethanol are the

main renewable transportation fuels produced by the United States. The production of both

has increased over the last decade due an increased interest and higher oil prices. Biodiesel

production has increased from 25 million gallons in 2004 to 678 million gallons in 2008.1

Though biodiesel has been commercially implemented, there are a number of problems

preventing its use as a drop-in replacement for diesel fuel. Biodiesel is partially oxidized and

has lower energy density than diesel fuel.2 Moreover, biodiesel is less stable and requires

special storage and transportation methods.2 Microorganisms can consume biodiesel and

antimicrobials must be used if it will be stored for large amounts of time.2 Biodiesel must be

mixed with petroleum fuels to avoid cold flow issues in colder climates.3 The majority of

these issues arise from biodiesel being chemically distinct from traditional transportation

fuels. Biodiesel is a fatty acid methyl ester (FAME) whereas gasoline and diesel are

composed of hydrocarbons. In order for a biofuel to be used as a drop-in replacement for

gasoline, it must also be composed of hydrocarbons. Heterogeneous catalysis provides an

economically feasible method to remove the oxygen atoms from fatty acids (FAs) to produce

biologically derived fuels which are chemically identical to current transportation fuels.

A significant amount of research has been done on the catalytic deoxygenation of

fatty acids.4-14 Fatty acids are deoxygenated via two separate pathways, decarboxylation and

decarbonylation, which are displayed on the next page.

51

Decarboxylation:

2CORCOOHR +→− (1)

Decarbonylation:

OHCORCOOHR 2++→− = (2)

Decarboxylation is the preferred pathway because it does not consume H2. Decarbonylation

requires H2 to hydrogenate the alkene product, R=, and produces CO, a well-known catalyst

inhibitor. Immer et al. determined that high decarboxylation selectivity can be achieved by

maintaining low H2 and CO partial pressures within the reactor.13

The majority of fatty acid deoxygenation research has been conducted using model

compounds; little has been done with biologically derived fatty acids. Currently, there has

been limited research regarding the deoxygenation of fatty acids derived from biological

sources. The research in this chapter intends to fill in this gap of knowledge by studying the

deoxygenation of model compounds and biomass-derived fatty in a 50-ml autoclave. Model

and biomass-derived fatty acids will also be deoxygenated in a 600-ml autoclave to simulate

industrial production of fuel precursor hydrocarbons.

2. Experimental Methods

2.1. Materials

The following reagent-grade chemicals were purchased and used as received: 90%

oleic acid (Sigma-Aldrich), 97% stearic acid (Acros), 98% palmitic acid (Acros), and 99+%

dodecane (Alfa-Aesar). 5% Pd/C (E117PB/W) was obtained from Evonik-Degussa. The

median particle size of E117 is 35 µm; the E117 surface area was determined to be 797 m2/g.

52

The E117 Pd particles have an average dispersion of 19.5% and are uniformly distributed

throughout the support. Certified balance H2/He mix, and H2 and He ultra-high-purity gases

(99.9999%) were obtained from National Welders. Biologically derived fatty acids were

derived by thermal hydrolysis of food grade lard and canola oil in a 5-l batch reactor.15

2.2 50-ml Autoclave Deoxygenation Experiments

Semi-batch FA deoxygenation experiments were conducted in a 50-ml stirred

autoclave reactor (Autoclave Engineers). Gas flow rate and purge gas composition were set

by mass flow controllers (Brooks 5850E series). A 20-ml condenser was used to collect

condensable vapors from the reactor effluent. The reactor pressure was controlled by a

manual back pressure regulator (Tescom) downstream from the condenser.

In a typical experiment, 22.5 g dodecane and 336 mg of 5% Pd/C catalyst (E117) were added

to the reactor. The catalyst was suspended in solvent by stirring at 240 rpm. The reactor

system was flushed with He for 5 min to remove air. The reactor system was then flushed

with 30 ml/min H2 for 5 min. The pressure was then increased to 2 atm. The catalyst was

reduced in situ at 200 °C with a 5 °C/min ramp and a 1-h soak. The reactor was cooled to 30

°C before the purge gas was switched back to He. The reactor was removed under He flow

and approximately 5.6 mmol FA were added to the reactor system manually. The reactor

was then sealed and purged under He flow for 5 min while agitating at 240 rpm. Afterward,

the reactor purge gas was switched to 60 ml/min 5% (H2/He) for 5 min, and the pressure was

subsequently raised to 15 atm and the agitation rate was increased to 1000 rpm. The reactor

was heated to 300 °C at 5 °C/min and held at 300 °C for the reaction time (typically 4 h). The

53

reactor was cooled to 30 °C before liquid samples were collected. The condensate was

sampled after reaction for analysis.

2.3 Scale-up Deoxygenation Experiments

Semi-batch experiments were conducted in a 600-ml stirred autoclave reactor

(Autoclave Engineers) using 200 g FA, 100 g solvent, and 8.375 g of suspended 5% Pd/C

catalyst. Gas flow rates to the autoclave were controlled by mass flow controllers (Brooks

5850E Series). The effluent gases passed through a 300-ml stainless steel condenser. The

reactor pressure was controlled downstream from the condenser by a back pressure regulator

(Tescom) and monitored by a strain gauge transducer (Omega).

In a typical experiment, dodecane and Pd/C (dried over night at 40 °C) were added to

the reactor body. The catalyst was suspended in the solvent by stirring at 1,000 rpm. The

reactor was purged with 1 slm of He for at least 15 min to remove air. The purge gas was

switched to 200 sccm of 10% H2/He, and the reactor pressure was increased to 2 atm. The

reactor was heated to 200 °C at a rate of 5 °C/min. The reactor was held at 200 °C for 1 h.

The reactor was cooled overnight under 50 sccm He. The next day, the reactor was opened at

25 °C under He flow, and 200 g FA were added. The reactor was purged with 500 sccm He

for at least 15 min to purge out any residual air in the reactor. The flow was then switched to

600 sccm of H2/He flow (6, 10 or 12 % H2/He). The stirring was turned on to 1,000 rpm and

the pressure was increased to 15 atm. The reactor was heated at 5 °C/min until 300 °C. The

reactor was kept at 300 °C until the reaction was completed as evidenced by on-line QMS.

The reactor was allowed to cool overnight. The reactor was heated to 100 °C before its

54

contents were drained and sampled. The condenser was heated before its contents were

collected and sampled.

2.4 Analytical Methods

The effluent stream was analyzed on-line using a QMS (Pfeiffer Prismaplus with

Quadstar 32-bit software) and capillary inlet system. The H2 (2 m/z), He (4 m/z), CO (28

m/z) and CO2 (44 m/z) signals were measured every 1 min throughout the course of the

reaction. The CO signal was corrected for CO2 electron impact fragmentation to CO+ by

subtracting 10% of the CO2+ (44 m/z) intensity.

The 50-ml autoclave reactor and condensate samples were analyzed using an HP5890

gas chromatograph (GC) equipped with a flame ionization detector (FID) and an Econocap

EC-5 30 m x 0.32 mm x 1.0 µm capillary column. Chromatograms were collected using an

SRI Model 333 Peak Simple Chromatography Data System. Peak integrations were

performed using PeakSimple Software. The following GC oven temperature program was

used: 5 °C/min ramp from 80 to 300 °C, 1 min soak at 300 °C. Samples (0.05 µL) were

injected onto the column inlet (300 °C, 10 psig head pressure) with a 50:1 split ratio. The n-

alkane and FA flame ionization detector (FID) responses were calibrated using an n-decane

internal standard. Response factors of n-alkanes and FAs were determined relative to n-

decane; equivalent response factors were assumed for the alkenes and alkanes of the same

carbon number.

Scale-up liquid phase products and biomass-derived FA reactants were analyzed on a

Perkin Elmer AutoSystem GC equipped with an FID and an EC-1 capillary column (30 m x

55

0.32 mm x 1.0 µm). Peak integrations were performed using Chromulan software. The GC

program ramped from 80 to 300 °C at 5 °C/min followed by a 5 min soak at 300 °C. The

detector was held at 325 °C. Samples (1 µL) were injected onto the column inlet (300 °C, 10

psi head pressure) with a 50:1 split ratio. The concentrations were determined relative to n-

decane as an internal standard. FID response factors were determined for stearic and

palmitic acid, dodecane, and heptadecane. The same FID response factors were used for

stearic acid, oleic acid, and linoleic acid, as well as n-heptadecane and heptadecenes.

2.5 Reactor Modeling – Initial Rates Calculation

In order to quantitatively compare the online QMS data, the initial decarboxylation

rate is calculated via a material balance. The CO2 mole balance below assumes that the

stirred autoclave contents (liquid and gas phases) are well-mixed; therefore, the effluent

composition reflects the instantaneous gas-phase composition in the reactor.

{ }

=+−

dt

dP

RT

VWrPP

RT

Q COg

COCOCO outin

2

2,2,2 (3)

Where CO2 r is the rate of CO2 generation, Q = purge rate, W = catalyst weight, Vg = volume

of reactor head space, R is the gas constant, and T is the absolute temperature. Since the CO2

partial pressure in the feed is negligible (PCO2,in ~ 0), Eq. 3 simplifies to

+=

dt

dP

Q

VP

WRT

Qr

COg

COCO2

22. (4)

56

The CO2 partial pressure is proportional to the He-normalized 44 m/z signal. Before the

reaction commences PCO2, out is approximately zero, and the initial rate is proportional to

dt

dPCO2 . The material balance simplifies to the following equation for initial rates:

=

dt

dP

WRT

Vr

COg

CO2

2 (5)

dt

dPCO2 is calculated by fitting initial linear region of the PCO2 vs. t plot. The linear fit

extrapolates to the point where it crosses the t-axis, determining the initial decarboxylation

rate within the reactor.

3. Results and Discussion

3.1 Analysis of Lard and Canola-Derived Fatty Acids

The compositions of lard and canola-derived FAs as determined GC-FID analysis are

displayed in Table 1. The primary constituent of lard and canola-derived FAs is oleic acid

(OA, C18:1). Essentially, half of the lard-derived FAs (LDFA) is comprised of OA. LDFA

also contains a significant amount of saturated FAs, palmitic and stearic acid. Linoleic acid is

also found in LDFA, but only in smaller quantities. Canola-derived FAs (CDFA) consist

primarily of unsaturated FAs. Nearly 75% of CDFA is OA. CDFA also contains a significant

portion of linoleic acid (C18:2). Only small amounts of saturated FAs were found in CDFA.

3.2 Model Compounds

Deoxygenation of OA, a major constituent of canola oil and lard-derived FAs, was

studied as a model of biologically derived fatty acids in a 50-ml stirred autoclave under 10%

57

flowing H2. The CO2 and CO production rates and H2 conversion are shown in Figure 2. A

prominent feature is the initial H2 uptake while the reactor is increasing in temperature. This

uptake peak is absent for the saturated fatty acids such as stearic or palmitic acid (PA). H2

consumption decreases before CO2 production begins. Near peak CO2 production, a small

amount of H2 is produced. However, after CO2 production decreases, H2 is again consumed.

Overall, 0.98 mol H2 is consumed per mole OA reacted consistent with the hydrogenation of

the OA double bond. The CO2 production increases quickly. The deoxygenation reaction was

93% selective toward CO2 and yielded 99.9% n-heptadecane product.

A PA concentration series was deoxygenated in dodecane; the total mass of PA and

dodecane was held constant. PA is the major saturated constituent of LDFA. As the semi-

batch reaction is scaled-up to increase productivity, it is important to know how

deoxygenation kinetics are affected by FA concentration. Studying PA deoxygenation over a

range of concentrations allows the effect of FA concentration on deoxygenation kinetics to

be determined.

The CO2 and CO production rates and effluent H2 percentage for semi-batch PA

deoxygenation experiments are displayed in Figure 3. Deoxygenation with an initial PA

concentration of 6% was highly selective toward CO2 production. CO2 production begins to

increase as the reactor reaches 300 °C, and H2 is evolved concurrently. The initial H2

evolution peak is followed by an H2 consumption trough. CO production is minimal, and the

reaction is complete in ~1 h. The 12 wt.% PA deoxygenation is also highly selective toward

CO2; however, the initial rate of CO2 production (decarboxylation) is lower. CO2 production

58

increases production at a slower rate. The relative initial decarboxylation rate (Table 2)

decreases significantly as the PA concentration increases. As the reactor reaches 300 °C, an

initial H2 evolution peak occurs which is followed by a subsequent H2 evolution peak as CO2

production reaches its maximum. Afterward, an extended H2 consumption trough occurs as

well. CO production is increased but does not cause a significant change in overall CO2

selectivity (Table 2). When the initial PA concentration is increased to 24 wt.%, CO

production is much higher initially. It is apparent that the initial decarbonylation rate

increases with PA concentration in Figure 3b. The initial H2 evolution peak is absent for 24

wt.% PA deoxygenation, and the CO2 maximum production occurs later. H2 consumed per

mole of PA converted is significantly higher for the 24 wt.% PA run than the lower

concentration runs (Table 2). The H2 consumption scales directly with CO2 selectivity as H2

is required to hydrogenate the olefin product of the decarbonylation pathway.16 The

conversions of all the runs in Figure 3 are above 90% with near 100% n-pentadecane yields

(Table 2).

The CO2 and CO flow rates and H2 consumption for 48 wt.% initial PA

deoxygenation are plotted in Figure 4. For this reaction, the CO flow rate rises before the

CO2 flow rate. The CO2 flow rate quickly overtakes the CO flow rate; however

decarboxylation remains the preferred pathway for less than an hour. Afterward, the

decarbonylation pathway becomes the most active. CO2 production reaches a local minimum

after 3 h. As CO production decreases, CO2 production begins to increase once again. After 4

h into the temperature program, decarboxylation becomes the preferred pathway and

59

continues to increase until the end of the temperature program. The reaction does not reach

completion (Table 3). H2 evolution is absent for this run; H2 consumption begins as the

decarbonylation and decarboxylation pathways initially increase in activity.

Reversible CO inhibition of the decarboxylation pathway is observed as it exhibits the

highest CO production levels for 48 wt.% PA. The maximum CO2 production level was

considerably less for 48 wt.% PA than the lower concentration runs. Consequently, the CO2

selectivity for 48 wt.% PA was significantly less than that of the lower concentration PA runs

(Table 2). As the CO production decreased, the decarboxylation pathway increased in

activity. Specific productivity per gram of catalyst was calculated and is displayed in Table

2; maximum productivity was observed for 24 wt.% initial PA concentration. As a result of

CO inhibition, catalyst productivity and relative initial decarboxylation rate decrease

drastically when increasing from 24 to 48 wt.% PA (Table 2). Moreover, H2 consumption

increases by nearly 50% by increasing from 6 to 24 wt.% PA. Because H2 is useful as a fuel,

productivity must be weighed with other parameters.

3.3 Canola Surrogate Mixture and Canola-Derived FAs

A canola FA surrogate mixture was made consisting of 5.2% palmitic, 2.3% stearic,

69.5% oleic, and 23.0% linoleic acids on a molar basis. The make-up of the canola-derived

fatty acids (CDFA) can be found in Table 1. The CO2 and CO flow rates and H2 molar flow

composition for canola surrogate mixture deoxygenation under 5% H2 are plotted in Figure 5.

During the temperature ramp, H2 consumption gives evidence of the hydrogenation of

unsaturated fatty acids. CO2 production begins to increase as the reactor reaches 300 °C. H2

60

is evolved concomitantly with CO2. After H2 evolution, an immediate H2 consumption

occurs. CO production is minimal throughout the course of the reaction. The reaction is

highly selective toward CO2 production (Table 3). H2 consumption was much lower than

expected based on canola FA surrogate mixture composition.

CDFA deoxygenation was also studied at the 50-ml scale. The CO2 and CO flow

rates and H2 molar flow composition of the reactor effluent stream are also plotted for CDFA

deoxygenation under 5% H2 in Figure 5. The behavior of CDFAs was very similar to that of

the surrogate mixture. An initial H2 uptake occurred while the reactor was increasing in

temperature. H2 was evolved as CO2 production began to increase. Subsequent H2

consumption was also observed for CDFA deoxygenation. CO2 production increases less

quickly and reaches a maximum at a lower value than observed for the canola surrogate

mixture. CO production was also minimal for CDFA deoxygenation. CO2 selectivity and H2

consumption were similar for the both the canola surrogate and CDFA experiments (Table

3). Both of the reactions reached completion within 1 h at reaction conditions.

Deoxygenation of the canola surrogate mixture and CDFA were examined under 10%

H2. The CO2 and CO production rate and H2 molar flow composition of the reactor effluent

stream for both reactions are plotted in Figure 6. For the canola surrogate mixture, there is an

initial H2 uptake as the reactor temperature increases. The catalyst evolves a slight amount of

H2 as the reactor reaches 300 °C which is immediately followed by a broad H2 consumption

trough. The CO2 production increases more slowly under 10% H2 than 5% H2. This effect is

shown quantitatively by the relative initial decarboxylation rates in Table 5. The CO2

61

production peak is also much broader; consequently, the reaction takes much longer to reach

completion under 10% H2. CO production increases before CO2 production. Overall, the

decarbonylation pathway is more active under 10% H2 consistent with H2 inhibition of the

decarboxylation pathway;13 the CO2 selectivity given in Table 3. Significantly more H2 is

consumed under H2 as the olefin decarbonylation product must be hydrogenated to produce

n-alkane.

Several aspects of CDFA deoxygenation were similar to surrogate mixture

deoxygenation under 10% H2. First, the H2 traces of both runs were similar. There is an

initial uptake while reactor temperature increases. A slight evolution of H2 also occurs as the

reactor reaches 300 °C; a H2 consumption trough immediately follows. The CO production

pathway also increases before CO2 production. However, the CO2 production is significantly

lower than displayed by the surrogate mixture. CO2 production increases at a lower rate, does

not reach the same value, and does not reach completion before the end of the temperature

program. Ultimately, CDFA shows signs of catalyst deactivation. The conversion is slightly

lower, and the relative initial decarboxylation rates are significantly lower for CDFAs (Table

3). CDFA deoxygenation is less selective toward CO2 and consumes more H2 than surrogate

mixture deoxygenation. Overall, slower reaction kinetics are observed for CDFA

deoxygenation; however, the catalyst remains moderately active even though the CDFAs

were not purified from any potential catalyst poisons.

62

3.4 Semi-Batch Deoxygenation Scale-up to 600-ml

The temporal CO2 and CO production rates and effluent H2 percentage from a stearic

acid (SA) deoxygenation in a 600-ml stirred autoclave are displayed in Figure 7. The initial

SA concentration was 67 wt.% in this experiment. CO flow rate increases before the CO2

flow rate. This behavior is similar to what was observed during 48 wt.% PA deoxygenation

in the 50-ml autoclave (Figure 4). This indicates that the decarbonylation pathway is

preferred initially for high concentrations of saturated FAs. The deoxygenation behavior of

SA may be interpreted in light of H2 and CO inhibition as described by Immer et al.13 As the

H2 partial pressure falls due to hydrogenation of heptadecenes, CO2 production begins to

increase. CO2 production exhibits its maximum rate when the H2 partial pressure reaches a

minimum in the reactor. As the H2 partial pressure in the reactor increases, the

decarboxylation reaction is inhibited, and CO2 production quickly decreases. Near the end of

the batch time, the CO partial pressure is low enough that the decarboxylation pathway

becomes less inhibited, and CO2 production begins to increase. After 6 h batch time,

deoxygenation switches over to favoring decarboxylation. Once the decarboxylation pathway

becomes dominant, CO2 production quickly increases, and the reaction approaches

completion. This data illustrates a dynamic relationship between the decarboxylation and

decarbonylation pathways. The CO product of the decarbonylation pathway reversibly

inhibits the decarboxylation pathway.

This reaction reached 100% conversion with 100% n-heptadecane yield (Table 4).

The CO2 selectivity and H2 consumption in Table 5 show a direct relationship between CO

63

production and H2 consumption, as there are no unsaturated FAs to hydrogenate prior to

reaction. The specific productivity for SA deoxygenation is the highest of all the FAs tested

at the 600-ml scale; however, productivity is lower than the values observed in the 50-ml

batch reactor (Table 2).

The temporal CO2 and CO production rate and effluent H2 percentage for OA

deoxygenation in the 600-ml stirred autoclave are displayed in Figure 8. A major distinction

between OA and SA deoxygenation is that OA deoxygenation requires H2 to hydrogenate the

OA double bond. H2 consumption begins as the reactor is heated, and essentially all of the H2

fed to the reactor is consumed at the H2 minimum percent composition. Consequently, as the

reactor reaches operating temperature, a significantly different reaction environment exists

for OA deoxygenation than SA deoxygenation. Because the partial pressure of H2 in the

reactor is much lower for OA deoxygenation, CO2 production increases sharply before the

decarbonylation pathway becomes active. The CO2 and CO flow rates show local maxima at

approximately 2 h batch time. As H2 partial pressure in the reactor begins to increase, both

CO2 and CO production decrease. CO2 production decreases much more rapidly than CO

production, and decarbonylation becomes the dominant deoxygenation pathway. After OA

has been mostly hydrogenated to SA, there is a significant change in the slope of the H2

consumption. This change in slope occurs while CO production reaches its maximum value

suggesting that the SA concentration reaches a maximum value at this time. Afterward, there

is a steady decline in both CO production and H2 consumption. Similar to SA deoxygenation,

the decarboxylation pathway regains activity as the CO partial pressure in the reactor

64

decreases. However, the CO partial pressure decreases far below the level at which the SA

decarboxylation pathway regained activity during SA deoxygenation with only marginal

increase in CO2 production. This suggests that the catalyst has undergone a significant

amount of deactivation during the reaction. Unsaturated feeds (e.g. OA) were found to

increase deactivation rate for FA deoxygenation over Pd/C.5

The CO2 selectivity for OA deoxygenation was less than SA deoxygenation (Table

4). The specific productivity for OA deoxygenation was significantly lower than SA

deoxygenation (Table 4). Under the same conditions, OA deoxygenation 17 h for completion

compared to less than 7 h for SA. Overall, a much higher amount of H2 was required

including contributions from OA and heptadecene hydrogenation (Table 4). We infer that the

decrease specific productivity for OA deoxygenation due to deactivation most likely caused

by the very low initial H2 partial pressures in the presence of OA.

The temporal CO2 and CO production rates and effluent H2 percentage for lard-

derived fatty acid (LDFA) deoxygenation under 12% H2 are displayed in Figure 9. As the

reactor temperature increases to 300 °C, nearly 100% of the H2 fed to the reactor is

consumed to hydrogenate unsaturated FAs. The initial H2 uptake period is considerably

shorter than for OA deoxygenation, which we attribute to the lower unsaturated fatty acid

content in the feed. Similar to OA deoxygenation, CO2 production increases before CO

production. The CO2 production increases more quickly and obtains a higher reaction value

than OA. This is attributed to there being a higher concentration of saturated fatty acids

initially available for reaction. CO production which begins to increase as the H2 partial

65

pressure begins to increase within the reactor. The maximum CO2 production rate occurs as

the CO production rate changes slope dramatically. CO2 production decreases much more

quickly than during OA deoxgenation as the H2 partial pressure increases more quickly for

LDFA under 12% H2. Again, the abrupt change in the derivative of effluent H2 percentage

corresponds to the maximum CO production rate. We infer that the majority of unsaturated

FAs have been hydrogenated by this point. Thereafter, H2 consumption is related mainly to

olefin hydrogenation. The decarbonylation pathway remains dominant until the end of the

reaction when there is a second spike in CO2 production (associated with reactor cooling).

The decarboxylation pathway regains activity during higher levels of CO production than OA

deoxygenation; however, CO2 production does not regain the same level of activity as

displayed when SA deoxygenation approaches completion. This effect could be attributed to

the higher H2 concentration instead of actual catalyst deactivation. Overall, LDFA

deoxygenation under 12% H2 displayed higher specific productivity than OA deoxygenation

under 10% H2 but not as high as SA under 10% H2 (Table 4). LDFA deoxygenation under

12% H2 also had a higher CO2 selectvity and a lower H2 consumption when compared to OA

deoxygenation (Table 4).

The temporal CO2 and CO production rates and effluent H2 percentage for LDFA

deoxygenation under 6% H2 are displayed in Figure 10. Initially, nearly all the H2 fed to the

reactor is consumed to hydrogenate unsaturated FAs. The FA hydrogenation stage takes

much longer to reach completion under 6% H2 than under 12% H2; (~ 9 h under reaction

conditions versus ~ 2 h). Although present, the change in the H2 consumption slope is less

66

pronounced for this run. The initial CO2 production is significantly less than that of the 12%

H2 LDFA run. CO partial pressure rose quickly within the reactor inhibiting the

decarboxylation pathway. Decarbonylation remained the dominant reaction pathway until

the reactor was cooled after 25 h. Consequently, the CO2 selectivity was much lower under

6% H2 than under 12% H2, and the overall reaction was much slower. Even under conditions

favorable to decarboxylation activity toward the end of reaction (i.e. low CO partial pressure,

low H2 partial pressure) the decarboxylation pathway did not regain activity. Also, the

reaction did not reach completion giving further evidence of catalyst deactivation (Table 4).

Given the high concentration of unsaturated FAs and the abundance of unsaturated product

via the decarbonylation pathway, the H2 concentration most likely was not sufficient to

prevent coke build-up on the catalyst surface. This explains why this run had the lowest

conversion and n-alkane yield (Table 4). As the reaction did not reach completion after 25 h

at reaction conditions, the catalyst productivity was significantly lower than the other

reactions (Table 4).

The temporal CO2 and CO production rates and effluent H2 percentage for CDFA

deoxygenation in dodecane under 10% H2 are displayed in Figure 11. The CDFA and OA

runs under 10% H2 are highly comparable. The CDFA and OA CO2 production curves are

closely similar, and the CO production curves also follow closely. However, the temperature

program was ended after 15 h batch time whereas the OA run ended after 18 h. The

conversion (Table 4) indicates that CDFA deoxygenation was incomplete after 15 h. The

CDFA had a higher selectivity toward CO2 product but still consumes more H2 per mol FA

67

converted when compared to OA deoxygenation (Table 5). This difference in H2

consumption is due to the lower CDFA conversion; less FA was deoxygenated to offset the

H2 consumed during unsaturated FA hydrogenation. Otherwise, CDFA and OA

deoxygenation consume nearly equivalent amounts of H2 throughout the reaction. Catalyst

productivity for CDFA and OA deoxygenation under 10% H2 are comparable and show little

evidence of the untimely temperature program termination.

CDFAs were also deoxygenated using CDFA deoxygenation product

(n-heptadecane) as solvent under 10% H2. The temporal CO2 and CO production rates and

effluent H2 percentage for this reaction using CDFA deoxygenation product as a solvent are

displayed in Figure 12. The CDFA deoxygenation product was used as solvent in this

reaction to determine if deoxygenation product was a viable solvent for commercial use. The

H2 consumption trace is very similar to the CDFA run with dodecane solvent. However, the

initial CO2 production is much lower than that of CDFA in dodecane. Decarbonylation

quickly becomes the dominant reaction pathway. The decarboxylation pathway remained

inactive until the very end of the reaction. This run displayed the lowest CO2 selectivity of all

the deoxygenations in the 600-ml autoclave (Table 5). Consequently, CDFA deoxygenation

in CDFA deoxygenation product consumed the highest amount of H2 in the table as well. The

conversion was higher than the CDFA run in dodecane as the temperature program went until

the catalyst end of the reaction (Table 5).

It has been shown that an effective increase in H2 partial pressure occurs when

solvents with lower vapor pressures are used.14 The CDFA product is predominantly n-

68

heptadecane which has a lower vapor pressure than dodecane. However, this is an unlikely

explanation of the lack of decarboxylation activity because the H2 partial pressure in the

reactor is initially lowered to effectively the same level by unsaturated FA hydrogenation.

Significant catalyst inhibition was observed for CDFA deoxygenation under 10% H2 in the

50-ml autoclave. Therefore, it is reasonable that the CDFA product decreases the activity of

the decarboxylation pathway. The catalyst was heated and cooled in the CDFA product

during the reduction process—ample time for catalyst poisons to adsorb to the catalyst

surface.

4. Conclusion

In the 50-ml stirred autoclave reactor, OA hydrogenation occurred prior to

deoxygenation. H2 consumption as reactor temperature increased was attributed to the

hydrogenation of the OA double bond. OA was mostly converted to SA when

decarboxylation began. High selectivity toward decarboxylation were observed, showing that

unsaturated FAs react efficiently under 10% H2. Initial decarboxylation rate decreased with

increasing PA deoxygenation; however, specific productivity approached a maximum value

near 24 wt.% PA. With increasing PA concentration, H2 consumption increased and CO2

selectivity decreased. Deoxygenation switched over from favoring decarboxylation to

decarbonylation for 48 wt.% PA deoxygenation. The QMS traces for CDFA and canola

surrogate mixture deoxygenation show close agreement under 5% H2 purge. However, under

10% H2, CDFA deoxygenation shows signs of catalyst deactivation.

69

In the 600-ml stirred autoclave reactor, the highest specific productivity was observed

for SA deoxygenation. Lower specific productivities were observed for unsaturated feed-

stocks. Initial hydrogenation did not reach completion before deoxygenation began for any

of the unsaturated feedstocks. The decarboxylation pathway was less active for all

unsaturated FAs. Prolonged hydrogenation which significantly lowered specific productivity

was observed with lower H2 percentages in the purge gas. OA and CDFA showed similar

CO2, CO and H2 traces for deoxygenation under 10% H2. Pre-hydrogenation of unsaturated

FAs is recommended to increase catalyst productivity. Specific productivities and CO2

selectivities were comparatively lower in the 600-ml stirred due to high FA concentration.

Maintaining lower FA concentration by fed-batch operation is recommended to increase CO2

selectivity and specific productivity.

70

References

1. Subramaniam, R.; Dufreche, S.; Zappi, M.; Bajpai, R., Journal of Industrial

Microbiology & Biotechnology 2010, 37 (12), 1271-1287 2. Tyson, K. S., Biodiesel Handling and Use Guidelines (3rd Ed.). 3 ed.; DIANE: 2006; p

61. 3. Pinto, A. C.; Guarieiro, L. L. N.; Rezende, M. J. C., Journal of the Brazilian Chemical

Society 2005, 16 (6B), 1313-1330. 4. Snare, M.; Kubickova, I.; Maki-Arvela, P.; Eranen, K.; Murzin, D. Y., Industrial &

Engineering Chemistry Research 2006, 45 (16), 5708-5715.

5. Snare, M.; Kubickova, I.; Maki-Arvela, P.; Chichova, D.; Eranen, K.; Murzin, D. Y., Fuel 2008, 87 (6), 933-945.

6. Maki-Arvela, P.; Kubickova, I.; Snare, M.; Eranen, K.; Murzin, D. Y., Energy & Fuels

2007, 21 (1), 30-41. 7. Simakova, I.; Simakova, O.; Maki-Arvela, P.; Simakov, A.; Estrada, M.; Murzin, D. Y.,

Applied Catalysis a-General 2009, 355 (1-2), 100-108. 8. Simakova, I.; Simakova, O.; Maki-Arvela, P.; Murzin, D. Y., Catalysis Today 2010, 150

(1-2), 28-31. 9. Kubickova, I.; Snare, M.; Eranen, K.; Maki-Arvela, P.; Murzin, D. Y., Catalysis Today

2005, 106 (1-4), 197-200. 10. Lestari, S.; Simakova, I.; Tokarev, A.; Maki-Arvela, P.; Eranen, K.; Murzin, D. Y.,

Catalysis Letters 2008, 122 (3-4), 247-251. 11. Lestari, S.; Maki-Arvela, P.; Simakova, I.; Beltramini, J.; Lu, G. Q. M.; Murzin, D. Y.,

Catalysis Letters 2009, 130 (1-2), 48-51. 12. Lestari, S.; Maki-Arvela, P.; Eranen, K.; Beltramini, J.; Lu, G. Q. M.; Murzin, D. Y.,

Catalysis Letters 2010, 134 (3-4), 250-257. 13. Immer, J. G.; Lamb, H. H., Energy & Fuels 2010, 24, 5291-5299. 14. Immer, J. G.; Kelly, M. J.; Lamb, H. H., Applied Catalysis a-General 2010, 375 (1), 134-

139.

71

15. Provided by Wilson Wang, PhD candidate of the North Carolina State University

Mechanical and Aerospace Engineering Department. 16. Ford, J.P. and Lamb, H.H., A Comparison of Supported Pd Catalysts for Liquid-Phase

Deoxygenation of Fatty Acids, Chapter 2.

72

Table 1: Composition (mole percent) of biologically derived FAs as determined by GC-FID analysis.

FA Lard-Derived FAs Canola-Derived FAs

Myristic (C14:0) 1.8 % -

Palmitic (C16:0) 32.6 % 4.7 %

Stearic (C18:0) 9.9 % 2.5 %

Oleic (C18:1) 47.2 % 75.5 %

Linoleic (C18:2) 8.5 % 17.3 %

a

73

Table 2: Results for PA deoxygenation over 5 wt.% Pd/C in a 50-ml stirred autoclave reactor.a

wt. % PA n-C15 Yield X

CO2 Selectivity

Initial Decarboxylation Rateb

H2 Consumptionc

Specific Productivity (mmol/gcat·h)

6 0.996 0.983 0.933 1.00 0.113 23.5

12 1.04 0.929 0.934 0.54 0.094 31.1

24 1.04 0.900 0.893 0.38 0.147 43.1

48 0.864 0.792 0.524 0.28 0.528 19.4

a) Reaction conditions: 336 mg catalyst, 23.94 g total mass of dodecane and PA, 5%

H2(He), 60 ml/min purge flow rate, 15 atm, 300 °C, and 5 h reaction time. The total mass of solvent and PA was held constant for each reaction.

b) Normalized to 1.44 g PA deoxygenation. c) Moles H2 consumed per mol FA converted.

74

Table 3: Deoxygenation results for canola-derived FAs and a canola FA surrogate mixture in a 50-ml stirred autoclave reactor.a

Feedstock

% H2

n-

Alkane Yield X

CO2 Selectivity

Relative Initial Decarboxylation

Rateb H2

Consumptionc

Specific Productivity (mmol/gcat·h)

Canola FA Surrogate 5 0.888 1.01 0.933 1.00 0.139

Canola FA Surrogate 10 0.845 1.04 0.840 0.15 0.307

Canola FAs 5 0.977 0.964 0.902 0.70 0.140

Canola FAs 10 0.950 0.937 0.740 0.076 0.592

a) Reaction conditions: 5.6 mmol canola derived FA or canola surrogate mixture, 336 mg catalyst, 22.5 g dodecane solvent, 60 ml/min purge flow rate, 15 atm, 300 °C, and 4 h reaction time. b) Normalized to 5% H2 canola surrogate mixture initial decarboxylation rate. c) Moles H2 consumed per mol FA converted.

75

Table 4: Deoxygenation results for model compounds and biomass-derived FAs in a 600-ml stirred autoclave reactor.a

Feedstock Solvent Purge Gas

n-Alkane Yield X

CO2 Selectivity

H2 Consumptionb

Productivity (mmol/gcat·h)

SA Dodecane 10% H2 1.01 1.01 0.424 0.66 12.1

OA Dodecane 10% H2 ---- 0.948 0.238 1.55 4.45

Lard Dodecane 12% H2 0.940 0.937 0.266 1.31 6.34

Lard Dodecane 6% H2 0.806 0.823 0.141 1.23 2.84

Canola Dodecane 10% H2 0.861 0.851 0.251 1.78 4.77

Canola

Canola HC

Product 10% H2 0.918 0.914 0.125 2.01 4.05

a) Reaction conditions: 300 °C, 600 ml/min purge flow rate, 15 atm, 200 g FA, 100 g solvent, and 8.375 g Pd/C. b) Moles H2 consumed per mol FA converted.

76

Figure 1: GC-FID chromatograms of lard-derived fatty acids (a) and canola-derived fatty acids (b).

0.00

0.02

0.04

0.06

0.08

0.10

0.12

0.14

0.16

0.18

0 5 10 15 20 25 30 35 40 45 50

Retention T ime (min)

FID Signal (a.u.)

0.00

0.02

0.04

0.06

0.08

0.10

0.12

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0.16

0.18

30 32 34 36

Reten tion Time (m in)

Myristic C14:0

Stearic C18:0

Oleic C18:1

Linoleic C18:2

Palmitic C16:0

a

0.00

0.05

0.10

0.15

0.20

0.25

0 5 10 15 20 25 30 35 40 45 50

Retention Time (min)

FID Signal (a.u.)

0.00

0.05

0.10

0.15

0.20

0.25

30 32 34 36

Stearic C18:0

Oleic C18:1

Linoleic C18:2

Palmitic C16:0

b

77

0

0.05

0.1

0.15

0.2

0.25

0.3

0

2

4

6

8

10

12

0 1 2 3 4

CO2

CO

H2

Flow Rate (mmol/min)

Efflu

ent %

H2

Time (h)

Figure 2: CO2 and CO molar flow rates and effluent mol% H2 for OA deoxygenation at 300 °C for 3 h under 10% H2(He) purge in dodecane.

78

0

0.1

0.2

0.3

0.4

0.5

0.6

50

100

150

200

250

300

0 0.5 1 1.5 2 2.5 3

6% PA

12% PA

24% PA

T

CO


T (

oC)

Time (h)

a

0

0.02

0.04

0.06

0.08

0.1

0.12

0.14

0

1

2

3

4

5

6

7

0 1 2 3

6% PA

12% PA

24% PACO Flow Rate (mmol/min)

Efflu

ent %

H2

Time (h)

b

Figure 3: CO2 (a) and CO (b) production rates in mmol/min and effluent mol% H2 (b) for PA deoxygenation at 300 °C for 5 h in dodecane under 5% H2(He) 60 ml/min at 15 atm in a 50-ml autoclave. The total mass of PA and solvent was held constant at 23.94 g.

79

0

0.05

0.1

0.15

0.2

0

1

2

3

4

5

0 1 2 3 4 5 6 7

CO2

CO

H2


Efflu

ent %

H2

Time (h)

Figure 4: CO2 and CO production rates and effluent mol% H2 for 48 wt.% PA deoxygenation at 300 °C for 5 h in dodecane under 5% H2(He) 60 ml/min at 15 atm in a 50-ml autoclave. The initial mass of PA added was 11.52 g. The total mass of PA and solvent was 23.94 g.

80

0

0.05

0.1

0.15

0.2

0.25

0.3

0.35

0

1

2

3

4

5

6

0 1 2 3 4

CO2

CO

H2


Efflu

ent %

H2

Time (h)

a

0

0.05

0.1

0.15

0.2

0.25

0.3

0.35

0

1

2

3

4

5

6

0 1 2 3 4

CO2

CO

H2


Efflu

ent %

H2

Time (h)

b

Figure 5: CO2 and CO molar production rates and effluent mol% H2 for (a) canola surrogate mixture (b) canola-derived FAs. Reaction conditions: 300 °C for 4 h in dodecane under 5% H2(He) 60 ml/min at 15 atm in a 50-ml stirred autoclave. 5.6 mmol of FA was added to the reactor.

81

0

0.02

0.04

0.06

0.08

0.1

0

2

4

6

8

10

0 1 2 3 4 5

CO2

CO

H2


Efflu

ent %

H2

Time (h)

a

0

0.02

0.04

0.06

0.08

0.1

0

2

4

6

8

10

0 1 2 3 4 5

CO2

CO

H2


Efflu

ent %

H2

Time (h)

Figure 6: CO2 and CO production effluent mol.% H2 for deoxygenation of (a) canola FA surrogate mixture (b) canola-derived FAs. Reaction conditions: 300 °C in dodecane under 10% H2(He) 60 ml/min at 15 atm in a 50-ml stirred autoclave. 5.6 mmol of FA was added to the reactor.

82

0

0.5

1

1.5

2

2.5

3

3.5

4

0

5

10

0 2 4 6 8

CO2

CO

H2


Efflu

ent %

H2

Time (h)

Figure 7: CO2 and CO molar production rates and H2 conversion for SA deoxygenation at 300 °C in dodecane under 10% H2(He) 600 ml/min at 15 atm in a 600-ml stirred autoclave. 200 g SA, 100 g dodecane, and 8.375 g Pd/C were added to the reactor. The reactor reached operating temperature (300 °C) at t = 1 h.

83

0

0.5

1

1.5

2

2.5

3

3.5

4

0

5

10

0 5 10 15 20

CO2

CO

H2


Efflu

ent %

H2

Time (h)

Figure 8: CO2 and CO molar production rates and H2 conversion for OA deoxygenation at 300 °C in dodecane under 10% H2(He) 600 ml/min at 15 atm in a 600-ml stirred autoclave. 200 g OA, 100 g dodecane, and 8.375 g Pd/C were added to the reactor. The reactor reached operating temperature (300 °C) at t = 1 h.

84

0

1

2

3

4

0

6

12

0 5 10 15

CO2

CO

H2


Efflu

ent %

H2

Time (h)

Figure 9: CO2 and CO molar production rates and H2 conversion for lard-derived FA deoxygenation at 300 °C in dodecane under 12% H2(He) 600 ml/min at 15 atm in a 600-ml stirred autoclave. 200 g lard-derived FA, 100 g dodecane, and 8.375 g Pd/C were added to the reactor. The reactor reached operating temperature (300 °C) at t = 1 h.

85

0

0.5

1

1.5

2

2.5

3

3.5

4

0

3

6

0 5 10 15 20 25

CO2

CO

H2


Efflu

ent %

H2

Time (h)

Figure 10: CO2 and CO molar production rates and H2 conversion for lard-derived FA deoxygenation at 300 °C in dodecane under 6% H2(He) 600 ml/min at 15 atm in a 600-ml stirred autoclave. 200 g lard-derived FA, 100 g dodecane, and 8.375 g Pd/C were added to the reactor. The reactor reached operating temperature (300 °C) at t = 1 h.

86

0

0.5

1

1.5

2

2.5

3

3.5

4

0

5

10

0 5 10 15

CO2

CO

H2


Efflu

ent %

H2

Time (h)

Figure 11: CO2 and CO molar production rates and H2 conversion for canola-derived FA deoxygenation at 300 °C in dodecane under 10% H2(He) 600 ml/min at 15 atm in a 600-ml stirred autoclave. 198 g canola-derived FA, 100 g dodecane, and 8.375 g Pd/C were added to the reactor. The reactor reached operating temperature (300 °C) at t = 1 h.

87

0

0.5

1

1.5

2

2.5

3

3.5

4

0

5

10

0 5 10 15 20

CO2

CO

H2


Efflu

ent %

H2

Time (h)

Figure 12: CO2 and CO molar production rates and H2 conversion for canola-derived FA deoxygenation at 300 °C in dodecane under 10% H2(He) 600 ml/min at 15 atm in a 600-ml stirred autoclave. 198 g canola-derived FA, 100 g canola-derived FA deoxygenation product, and 8.375 g Pd/C were added to the reactor. The reactor reached operating temperature (300 °C) at t = 1 h.

ABSTRACT FORD, JEFFREY PETER. Semi-Batch Catalytic ...

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