SIMULATION STUDY OF DISTILLATION,...

SIMULATION STUDY OF DISTILLATION, STRIPPING, AND FLASH

TECHNOLOGY FOR AN ENERGY EFFICIENT METHANOL RECOVERY

UNIT IN BIODIESEL PRODUCTION PROCESSES

A Thesis

Submitted to the Faculty of Graduate Studies and Research

In Partial Fulfillment of the Requirements for the

Degree of Master of Applied Science

in

Environmental Systems Engineering

University of Regina

By

Firuz Alam Philip

Regina, Saskatchewan

November, 2013

Copyright 2013: F. A. Philip

UNIVERSITY OF REGINA

FACULTY OF GRADUATE STUDIES AND RESEARCH

SUPERVISORY AND EXAMINING COMMITTEE

Firuz Alam Philip, candidate for the degree of Master of Applied Science in Enviornmental Systems Engineering, has presented a thesis titled, Simulation Study of Distillation, Stripping, and Flash Technology for an Energy Efficient Methanol Recovery Unit in Biodiesel Production Processes, in an oral examination held on August 23, 2013. The following committee members have found the thesis acceptable in form and content, and that the candidate demonstrated satisfactory knowledge of the subject material. External Examiner: Dr. Fanhua Zeng, Petroleum Systems Engineering

Co-Supervisor: *Dr. Amornvadee Veawab, Environmental Systems Engineering

Co-Supervisor: Dr. Adisorn Aroonwilas, Industrial Systems Engineering

Committee Member: Dr. Kelvin Ng, Environmental Systems Engineering

Committee Member: *Dr. Raphael Idem, Industrial Systems Engineering

Chair of Defense: Prof. Wes Pearce, Faculty of Fine Arts *Not present at defense

i

ABSTRACT

Biodiesel is an important alternative renewable energy source currently produced

by transesterification reaction of oil or fat with methanol. To improve the conversion,

excess methanol is required, which must be recovered from the product stream and

recycled back into the process for further biodiesel production. The intensive energy

requirements for methanol recovery are an important issue that directly impacts the

production costs of biodiesel. To reduce the cost of biodiesel production, an energy

efficient methanol recovery unit (MRU) is crucial.

This work focuses on energy requirement reduction by distillation, flash-based

recovery, and newly-introduced stripping-based methanol recovery units. Four different

continuous methanol recovery units were simulated using Aspen Plus. Energy

requirements with respect to process parameters including percentage of methanol

recovery, operating pressure, and methanol-to-oil ratio for all methanol recovery units

were analyzed. Units were compared in terms of energy requirement and purity of

recovered methanol product.

The simulation results show that energy requirement for methanol recovery units

increases with increase in % methanol recovery and reflux ratio (for distillation), but

decreases with decrease in operating pressure and increase in methanol-to-oil ratio. The

recovered methanol is pure for distillation and stripping-based MRUs. However, for

flash-based MRUs, the purity of recovered methanol degrades at the high heat duty

supplied. Consequently, the single- and double-flash-based MRUs have narrow ranges of

operation. Moreover, double-flash-based MRUs have no significant advantages over

ii

single-flash-based MRUs in terms of heat duty. Comparison of heat duty among

distillation, stripping, and single-flash reveals that the single-flash-based MRU is the

most energy efficient followed by stripping and distillation-based MRUs.

iii

ACKNOWLEDGEMENTS

I would like to express my sincere gratitude to my supervisors, Dr. Amornvadee

Veawab and Dr. Adisorn Aroonwilas for giving me the opportunity to carry out this

interesting research under their enthusiastic supervision. Their enormous financial and

technical support, valuable guidance, and encouragement were a great source of

inspiration and the driving force throughout the entire course of this research.

I would like to thank the Natural Sciences and Engineering Research Council of

Canada (NSERC), the City of Regina, and the Faculty of Graduate Studies and Research

(FGSR) for their financial support. I would also like to thank the Faculty of Engineering

and Applied Science at the University of Regina for their help and support.

I am thankful to Kazi Sumon for his help and encouragement. I would also like to

thank my research group and URBSA.

Finally, I am sincerely thankful and grateful to my parents, sisters, wife and all

other family members for their unconditional love, prayers, and support to fulfill my

dream.

iv

TABLE OF CONTENTS

ABSTRACT i

ACKNOWLEDGEMENTS iii

TABLE OF CONTENTS iv

LIST OF TABLES viii

LIST OF FIGURES ix

NOMENCLATURE xiii

1. INTRODUCTION 1

1.1 Use of biodiesel as an alternative fuel 1

1.2 Significance of methanol recovery during biodiesel production 2

1.3 Alcohol recovery technologies, commercial applications, and

research works

8

1.4 Research motivation, objectives, and scope of work 12

2. BACKGROUND AND LITERATURE REVIEW 14

2.1 Important reactions associated with biodiesel production 14

2.1.1 Transesterification reaction 14

2.1.2 Esterification reaction 15

2.1.3 Soap formation 15

2.2 Biodiesel production processes 16

2.2.1 Base-catalyzed transesterification process 16

2.2.2 Two-step process 18

2.2.3 Acid-catalyzed process 20

v

2.2.4 Biox process 20

2.2.5 Supercritical process 22

2.3 Literature review on methanol recovery 24

3. METHODOLOGY 26

3.1 Process simulation procedures 26

3.2 Process flow schemes of simulated methanol recovery units 27

3.3 Chemical components 31

3.4 Process simulation 31

3.4.1 Simulation basis 31

3.4.2 Simulation framework 32

3.4.3 Simulation input 32

3.4.3.1 Operating parameters 32

3.4.3.2 Physical property parameters of chemical components 34

3.4.3.3 Type of thermodynamic model 38

3.4.4 Model calculations 41

3.4.4.1 Methanol recovery unit model 41

A) Distillation unit 43

B) Flash unit 45

3.4.4.2 Vapor-liquid equilibrium model 47

3.4.5 Simulation Outputs 48

4. SIMULATION RESULTS AND DISCUSSION 50

4.1 Distillation-based MRU 50

4.1.1 Parametric effect on heat duty 51

vi

A) Percentage of MeOH recovery 51

B) Operating pressure of distillation unit 53

C) Reflux ratio 53

D) MeOH-to-oil ratio 56

4.1.2 Quality of recovered MeOH 58

4.2 Stripping-based MRU 58

4.2.1 Parametric effects on heat duty 60


B) Operating pressure of stripping 60

C) MeOH-to-oil ratio 63


4.3 Single-flash-based MRU 66

4.3.1 Parametric effects on heat duty 66


B) Operating pressure 66



4.4 Double-flash-based MRU 72

4.4.1 Parametric effect on heat duty 72


B) Operating pressure 73



vii

4.5 Overall comparison of MRUs 79

4.5.1 Heat duty requirement 79

4.5.2 Product quality and quantity of recovered MeOH 83

5. CONCLUSIONS AND FUTURE WORK 84

5.1 Conclusions 84

5.2 Recommendations for future work 85

REFFERENCES 86

viii

LIST OF TABLES

Table 1.1 Top 10 countries that have great biodiesel production potential. 3

Table 1.2 Contents of fatty acid in various feedstocks 4

Table 1.3 Molar ratio of methanol to oil used for transesterification of

different types of oil feedstock

6

Table 1.4 Industrial practice of methanol recovery 9

Table 1.5 Previous work on methanol recovery 10

Table 3.1 Operating parameters for transesterification unit 35

Table 3.2 Process parameters for methanol recovery units 36

Table 3.3 Scalar property of methanol, glycerol, and FAME 37

Table 3.4 Properties of triolein found in the literature 39

Table 3.5 Physical properties of triolein predicted by Gani method 39

Table 3.6 The thermodynamic models used by various authors for

biodiesel process simulations

40

Table 3.7 UNIFAC DMD group assignment for this study 42

ix

LIST OF FIGURES

Figure 2.1 Base-catalyzed transesterification process 17

Figure 2.2 Two-step process 19

Figure 2.3 Acid-catalyzed process 21

Figure 2.4 Biox co-solvent process 21

Figure 2.5 Supercritical transesterification process 23

Figure 3.1 Process flow schemes of transesterification and methanol

recovery units (a) distillation-based (b) stripping-based (c)

single flash-based (d) double flash-based.

29

Figure 3.2 Simulation framework for methanol recovery by methanol

recovery units

33

Figure 3.3 Schematic diagram of Distillation column (a) full column (b)

tray of the column

44

Figure 3.4 Schematic diagram of a single-stage flash unit 46

Figure 4.1 Effect of % MeOH recovery on reboiler heat duty for the

distillation-based MRU at the reflux ratio of (a) 2 (b) 3 (c) 4

(column pressure = 0.1-1.0 atm, total stage = 7, feed stage = 4,

MeOH to oil ratio = 6, feed temperature = 60°C and feed

pressure = 4 atm).

52

Figure 4.2 Effect of pressure on reboiler heat duty for the distillation-

based MRU at the reflux ratio of (a) 2 (b) 3 (c) 4 (total stage =

7, feed stage = 4, MeOH to oil ratio = 6, feed temperature =

54

x

60°C and feed pressure = 4 atm).

Figure 4.3 Effect of reflux ratio on reboiler heat duty for the distillation-

based MRU at the column pressure of (a) 0.1 atm (b) 0.2 atm

(c) 0.3 atm (d) 0.5 atm (e) 1.0 atm (total stage = 7, feed stage

= 4, MeOH to oil ratio = 6, feed temperature = 60°C and feed

pressure = 4 atm).

55

Figure 4.4 Effect of MeOH-to-oil ratio on reboiler heat duty for the

distillation-based MRU at the reflux ratio of (a) 2 (b) 3 (c) 4

(column pressure = 0.2 atm, total stage = 7, feed stage = 4,

feed temperature = 60°C and feed pressure = 4 atm).

57

Figure 4.5 Quality of recovered MeOH from the distillation-based MRU

at the reflux ratio (a) 2 (b) 3 (c) 4 (column pressure = 0.2 atm,

total stage = 7, feed stage = 4, MeOH to oil ratio = 6, feed

temperature = 60°C and feed pressure = 4 atm).

59

Figure 4.6 Effect of % MeOH recovery on reboiler heat duty for the

stripping-based MRU (column pressure = 0.1-1.0 atm, total

stage = 7, feed stage = 1(top), MeOH to oil ratio = 6, feed


61

Figure 4.7 Effect of operating pressure on reboiler heat duty for the

stripping-based MRU (total stage = 7, feed stage = 1 (top),

MeOH to oil ratio = 6, feed temperature = 60°C and feed

pressure = 4 atm).

62

Figure 4.8 Effect of MeOH-to-oil ratio on reboiler heat duty for the 64

xi

stripping-based MRU at the pressure of (a) 0.2 atm (b) 0.5 atm

(c) 1.0 atm (total stage = 7, feed stage = 1 (top), feed


Figure 4.9 Quality of recovered MeOH from the stripping-based MRU at

the pressure of (a) 0.2 atm (b) 0.5 atm (c) 1.0 atm, (total stage

= 7, feed stage = 1 (top), MeOH to oil ratio = 6, feed


64

Figure 4.10 Effect of % MeOH recovery on heat duty for the single-flash-

based MRU (operating pressure = 0.1atm - 1.0 atm, MeOH to

oil ratio = 6, feed temperature = 60°C and feed pressure = 4.0

atm).

67

Figure 4.11 Effect of pressure on heat duty for the single-flash-based

MRU (MeOH to oil ratio = 6, feed temperature = 60°C and

feed pressure = 4.0 atm).

68

Figure 4.12 Effect of MeOH-to-oil ratio on heat duty for the single-flash-

based MRU at the pressure of (a) 0.2 atm (b) 0.5 atm (c) 1.0

atm (feed temperature = 60°C and feed pressure = 4.0 atm).

70

Figure 4.13 Quality of recovered MeOH by the single-flash-based MRU at

the pressures of (a) 0.2 atm (b) 0.5 atm (c) 1.0 atm (MeOH to

oil ratio = 6, feed temperature = 60°C and feed pressure = 4.0

atm).

71

Figure 4.14 Effect of percentage of MeOH recovery on heat duty for the

double-flash-based MRU with the first-stage pressure of 1.0

74

xii

atm and the second-stage pressure of 0.1, 0.2, 0.3, and 0.5

atm. (MeOH to oil ratio = 6, feed temperature = 60°C and

feed pressure = 4.0 atm).

Figure 4.15 Effect of first–stage flash pressure on heat duty for the

double-flash-based MRU with the second–stage flash pressure

of (a) 0.1 atm and (b) 0.2 atm. (MeOH to oil ratio = 6, feed

temperature = 60°C and feed pressure = 4.0 atm).

75

Figure 4.16 Effect of MeOH-to-oil ratio on heat duty for the double-flash-

based MRU at the first-stage pressure of 1.0 atm and the

second-stage pressure of 0.5 atm. (feed temperature = 60°C

and feed pressure = 4.0 atm).

76

Figure 4.17 Quality of recovered MeOH by the double-flash-based MRU

at the first/second pressures of (a) 1.0/0.1 atm (b) 1.0/0.2 atm

(c) 1.0/0.5 atm (MeOH to oil ratio = 6, feed temperature =

60°C and feed pressure = 4.0 atm)

78

Figure 4.18 Comparison between single-flash and double-flash operations

at the final pressure of (a) 0.1 atm (b) 0.2 atm. (MeOH to oil

ratio = 6, feed temperature = 60°C and feed pressure = 4.0

atm)

80

Figure 4.19 Comparison of heat duty among distillation, stripping and

single-flash at the pressure of (a) 0.2 atm, (b) 0.5 atm and (c)

1.0 atm. (Reflux ratio = 2 (for distillation), feed pressure = 4

atm, feed temperature = 60°C and MeOH-to-oil ratio = 6)

82

xiii

NOMENCLATURE

CH3OH Methanol

C57H104O6 Triolein

C19H36O2 Oleic acid methyl ester

C3H8O3 Glycerol

DG Diglyceride

GL Glyceride

FFA Free fatty acid

FAME Fatty acid mono alkyl ester

HCl Hydrochloric acid

H2O Water

H2SO4 Sulfuric acid

KOH Potassium hydroxide

�� Equilibrium constant

MeOH Methanol

MRU Methanol recovery unit

ML Million liter

MG Monoglyceride

NaOH Sodium hydroxide

NaOCH3 Sodium methoxide

NRTL Non-Random Two-Liquid

R – OH Alkyl alcohol

xiv

R – COOH Fatty acid

R – COOCH3 Methyl ester

TG Triglyceride

UNIQUAC Universal quasi chemical

UNIFAC UNIQUAC functional activity coefficient

UNIFAC-DMD UNIFAC Dortmund modified

VLE Vapor-liquid-equilibrium

Greek Letters

�� Vapor fugacity coefficient for component i

�� Liquid activity coefficient for component i

1

1. INTRODUCTION

1.1 Use of biodiesel as an alternative fuel

World energy demand is increasing significantly with increases in population,

globalization, heavy industrialization, and advancement of technologies. A major source

of the energy demand is from nonrenewable fossil fuels. According to Atadashi et al.

(2011), nonrenewable fossil fuels contribute to 86% of the world energy consumption.

However, reserves of fossil fuels are limited. With political pressure and environmental

concerns on the use of fossil fuels, an alternative energy source to the fossil fuels that is

renewable is necessary to overcome dependence on fossil fuels.

Such alternative renewable energy sources can be from solar, wind,

hydroelectricity, and bio-fuels including biodiesel and bio-ethanol. Among these

renewable energy sources, biodiesel is of promise as it has advantages over petroleum

diesel in aspects of environmental friendliness, renewability, higher combustion

efficiency, and lower sulfur and aromatic hydrocarbon contents. The biodiesel is

biodegradable and nontoxic, and its combustion leads to low emissions of greenhouse

gases and hydrocarbons compared to petroleum diesel (Demirbas, 2007).

Biodiesel has great potential to be a part of a sustainable energy mix in the future.

Many countries have been producing and consuming biodiesel as an alternative fuel to

petroleum diesel. Around the world, annual biodiesel production has increased from 15

thousand barrels per day in 2000 to 289 thousand barrels per day in 2008 (Atabani et al.,

2012). In European countries, the annual biodiesel production has increased significantly

as, for example, from 50 kiloton in 1993 to 2850 kiloton in 2005 (Demirbas, 2009). As

2

shown in Table 1.1 (Balat et al., 2010), Malaysia has the highest production potential,

followed by Indonesia, Argentina, USA, Brazil, Netherlands, Germany, Philippines,

Belgium, and Spain. The production cost of biodiesel is in the range of 0.53-1.71 US$/L

as reported in 2010.

1.2 Significance of methanol recovery during biodiesel production

The term biodiesel was derived from bio meaning life and diesel referring to the

processed fuel derived from biological sources (Demirbas, 2009). In technical terms,

biodiesel is a diesel engine fuel composed of monoalkyl esters of long-chain fatty acids

derived from biological feedstock, which can be categorized into three groups: vegetable

oils (edible and non-edible oils), animal fats, and used oil or grease. Regardless of

feedstock types, the main component of any oil or fat is triglyceride (TG), which is the

ester of glycerol and higher fatty acid. Different types of feedstock contain different types

of fatty acids. As seen in Table 1.2 (Lotero et al., 2005), the main fatty acid chains

present in the feedstock are palmitic acid, stearic acid, oleic acid, and linoleic acid.

Biodiesel is produced by a chemical reaction known as transesterification. In

transesterification reactions (Reaction 1.1), triglycerides of oil or fat break down into

fatty acid, which reacts with alcohol in the presence of a catalyst (usually a base) to

produce fatty acid mono alkyl ester (FAME), known as biodiesel, and glycerol as a

byproduct (Demirbas, 2009). Both ethanol and methanol can be used as the alcohol for

transesterification. However, methanol is most commonly used due to its lower cost.

3

Table 1.1: Top 10 countries that have great biodiesel production potential (Balat et al.,

2010).

Country Volume potential (MLa/year)

Production cost (US$/L)

Malaysia 14540 0.53 Indonesia 7595 0.49 Argentina 5255 0.62

USA 3212 0.70 Brazil 2567 0.62

Netherlands 2496 0.75 Germany 2024 0.79

Philippines 1234 0.53 Belgium 1213 0.78

Spain 1073 1.71 a Million liter

4

Table 1.2: Contents of fatty acid in various feedstocks (Lotero et al., 2005)

Feedstock Content of fatty acid chains (wt %)

Palmitic acid

Palmitoleic acid

Stearic Acid

Oleic acid

Linoleic acid

Linolenic acid

Rapeseed oil 3.5 N/A 0.9 64.4 22.3 8.2

Virgin olive oil 9.2 0.8 3.4 80.4 4.5 0.6

Sunflower oil 6.0 N/A 4.2 18.7 69.3 N/A

Safflower oil 5.2 N/A 2.2 76.3 16.2 N/A

Soybean 10.6 N/A 4.8 22.5 52.3 8.2

Palm oil 47.9 N/A 4.2 37 9.1 0.3

Choice white grease

23.3 3.5 11.0 47.1 11 1.0

Poultry fat 22.2 8.4 5.1 42.3 19.3 1.0

Lard 17.3 1.9 15.6 42.5 9.2 0.4

Edible tallow 28.4 N/A 14.8 44.6 2.7 N/A

Yellow grease 23.2 3.8 13.0 44.3 7.0 0.7

Brown grease 22.8 3.1 12.5 42.4 12.1 0.8

5

O

O

O

O

O

O

R1

R2

R3

+ R OHCatalyst

R2

O

O R

R3

O

O R

R1

O

O R

+

OH

OH

OH

3 (1.1)

Triglyceride

(Oil/Fat)

Alcohol Esters

(Biodiesel)

Glycerol

The transesterification reaction can be carried out through two types of processes,

i.e., catalytic and non-catalytic processes. The catalytic process can be referred to as a

homogeneous catalyzed process, heterogeneous catalyzed process, or enzyme catalyzed

process. In the homogeneous catalyzed process, the transesterification can be achieved by

either alkali or acid catalyst. For the non-catalytic process, also known as a supercritical

process, reactants are elevated to super critical state to form a single phase and reaction

temperature ranges from 350-400°C, and the reaction pressure is maintained above 80

atm (Gerpen et al., 2004).

The transesterification reaction of oil or fat is a reversible reaction. Stoichiometry

of the reaction indicates the molar ratio of alcohol to oil is 3:1, i.e., 3 mole of alcohol is

required to react with 1 mole of oil. Since the reaction is reversible, it does not continue

to completion. To shift the reaction equilibrium to produce a higher product yield, a

higher ratio of alcohol to oil is commonly used. The amount of alcohol used varies from

process to process and with the type of feedstock used. As seen from Table 1.3, all

processes employ higher ratios of alcohol to oil than 3:1. For example, the homogeneous

base catalyzed process requires the least alcohol-to-oil ratio, i.e., 6:1, while the

6

Table 1.3: Molar ratio of methanol to oil used for transesterification of different types of

oil feedstock

Oil Transesterification

process Methanol to oil ratio (mol:mol)

References

Sunflower oil Homogeneous base catalyzed 6:1 Umer et al. (2008)

Rapeseed oil Homogeneous base catalyzed 6:1 Umer and Farooq (2008)

Karanja oil Homogeneous base catalyzed 6:1 Meher et al. (2006)

Waste cooking oil Homogeneous acid catalyzed 20:1 Wang et al. (2006)

Soybean oil Homogeneous acid catalyzed 30:1 Narasimharao et al. (2007)

Sunflower oil Heterogeneous catalyzed 53:1 Babu et al. (2008)

Soybean oil Heterogeneous catalyzed 20:1 Gercia et al. (2008)

Palm oil Heterogeneous catalyzed 12:1 Bo et al. (2007)

Waste cooking oil Heterogeneous catalyzed 18:1 Jacobson et al. (2008)

Sunflower oil Supercritical 40:1 Madras et al. (2004)

Rapeseed oil Supercritical 42:1 Saka and Kusdiana, (2001)

7

heterogeneous catalyzed and supercritical processes require a much higher alcohol-to-oil

ratio, which is up to 53:1.

As discussed above, the biodiesel production typically requires a large quantity of

alcohol and leads to excess alcohol (after participating in conversion reactions) leaving

the reactor with the product. The excess alcohol must be recovered from the product

stream in alcohol recovery units and recycled back to the reactor for further biodiesel

production. The alcohol recovery is required during biodiesel production due to a number

of reasons. First, the biodiesel product must be purified to contain not more than 0.2%

alcohol (as methanol) (Knothe et al., 2005) to meet the biodiesel product specification

(ASTM D6751 or EN 14214 standard). Second, the alcohol recovery helps prevent

certain operational difficulties, including: 1) emulsification, which might take place due

to the presence of a hydroxyl group in the alcohol, rendering severe difficulties in

separation of biodiesel layer from water (Sharma et al., 2008); 2) lengthy time for

separation of biodiesel and glycerol when using gravity settling technique due to the

excess alcohol acting as a stabilizer (Gerpen et al., 2004); 3) the alcohol recovery

contributes to cost savings due to the reduced amount of alcohol required for biodiesel

production and due to the reduced flow rate of the stream in the product recovery units,

which reduces pumping requirements and loads in distillation used for final purification

of biodiesel product (Harding et al., 2007); and 4) the alcohol recovery leads to a

reduction in alcohol discharge to the environment, reducing health risks to communities.

It is also dangerous to handle and store biodiesel if it contains excess alcohol as it is

highly flammable (Gerpen et al., 2004).

8

1.3 Alcohol recovery technologies, commercial applications, and research works

The alcohol recovery unit has become an integral part of the biodiesel production

process. Common technologies used for alcohol recovery are distillation and flash. These

two technologies have been used in practice in a number of commercial biodiesel

production plants. For example, from Table 1.4, SRS Engineering Corporation provides

methanol recovery systems using distillation SRXC and ASV series for the

transesterification-based biodiesel production plants of Kyoto Fuels Corporations in

Alberta (Canada). Wintek Corporation provides methanol recovery systems using

vacuum flash towers, which can be single stage or multistage and capable of more than

99% methanol recovery. This technology has been used by a number of commercial

transesterification-based biodiesel production plants, such as Milligan Biotech (Canada),

Keystone Biofuels (USA), Middletown Biofuel (USA), and Innovation Fuels (USA).

These two technologies have also been used and studied by a number of

researchers. Table 1.5 lists the research works that have a direct emphasis on the study of

methanol recovery. As seen from the table, Dhar and Kirtania (2009), Baroutian et al.

(2010), and Kiss and Ignat (2012) carried out performance evaluations of distillation-

based methanol recovery units while Tang et al. (2010) and Wang et al. (2011) carried

out performance evaluations of flash-based methanol recovery units.

For the study of distillation technology for methanol recovery, Dhar and Kirtania

(2009) performed a simulation study using Aspen Plus to investigate parametric effects of

alcohol-to-oil ratio, reflux ratio and distillation column pressure on both methanol

recovery performance and energy consumption. Baroutian et al. (2010) carried out an

experimental study on parametric effects of heating temperature, permeate flow rate, and

9

Table 1.4: Industrial practice of methanol recovery

Reference Biodiesel production process

Methanol recovery process Methanol recovery and

purity

Existing plant Energy consumption Capacity Company name and

location

SRS Engineering Corporation (http://www.srsbiodiesel.com)

Transesterification Distillation column

The SRXC-Series – Removes Methanol from the Glycerine stream The ASV-Series - Recovers excess Methanol for cleaner unwashed fuel

> 99.9 % recovery and > 99.9 % purity

66 MMly Kyoto Fuels corporations Lethbridge, Alberta, Canada

-

Incbio (http://www.incbio.com)

Base catalysed transesterification

Vacuum evaporation module > 99.9% purity - - -

Acid catalyzed esterification

Fractional distillation column > 99.9% purity - - -

Wintek Corporation (http://www.wintek-corp.com)

Transesterification Vacuum Flash tower ( 1, 2 or 3-stages)

> 99 % recovery

- - 12 gpm 20 gpm

Milligan Biotech, Canada Keystone Biofuels, Camp Hill, PA, USA Middletown Biofuel, Middletown, PA,USA Innovation Fuels Newark, NJ, USA

-

10

Table 1.5: Previous work on methanol recovery

Reference Biodiesel production

process

Method of methanol recovery

Methodology Purpose of work Conditions (range) Finding(s)

Dhar and Kirtania, 2009

Acid catalyzed Esterification followed by alkali catalyzed Transesterification (Two step process)

Distillation column (Total stage 10)

Simulation (Aspen Plus)

To investigate effect of alcohol to oil ratio, reflux ratio and column pressure on methanol recovery and energy consumption

Alcohol to oil ratio = 6:1 – 50:1 Reflux ratio = 1 – 4 Column Pressure = 0.5, 0.75 and 1 atm.

The energy requirement increases with the increase in methanol to oil ratio used Reboiler heat duty reduced under vacuum

Baroutian et al., 2010

Homogeneous base catalyzed Transesterification

Distillation Experimental study using a three neck round bottom flask, oil bath and condenser.

To investigate influences of heating temperature, permeate flow rate and methanol to oil volume ratio on recovery of methanol with time

Temperature = 80- 130 °C Permeate (feed flow rate) = 2.4-12.3 ml/min Reactant ratio (volume) used = 1:1, 1.5:1 and 2:1

At higher parameter value methanol recovery was high.

Tang et al., 2010

Supercritical Process (methanol: oil = 30:1)

Flashing (one and two stage)

Simulation (Using mathematical model)

To investigate effect of feed temperature, feed pressure and flash pressure on methanol concentration and recovery To compare one stage and two stage flashing

Feed pressure = 15 and 25 MPa Flashing pressure = 0.4 – 5 MPa Feed temperature = 210 – 260 °C

85% recovery of methanol with 99% purity was possible at feed pressure 15 – 30 MPa and flashing pressure 0.4 MPa. The recovery of methanol for one-stage and two- stage was close

11

Wang et al., 2011

Supercritical Process

Flashing evaporation

Experimental study

To investigate effect of reactor (feed to flash unit) pressure, flash pressure and reactor temperature on methanol concentration and recovery

Feed pressure = 9 – 15 MPa Flash pressure = 0.2 – 2.0 MPa Feed temperature = 240 – 300 °C

Feed pressure and temperature has effect on methanol recovery and purity but effect of flashing pressure was most. 85% recovery of methanol with 99% purity was possible at feed pressure 15 MPa, feed temperature 300 °C and flashing pressure 0.4 MPa.

Kiss and Ignat, 2012

Transesterification Distillation (Dividing-wall column i.e. DWC)

Simulation (Aspen Plus)

To optimize conventional distillation column and proposed DWC for methanol recovery and hence compare in terms of cost and energy.

Feed flow 2900 kg/hr Feed Temperature 60 °C Feed Pressure 1.2 bar Column Pressure 0.5 bar

DWC reduced energy requirement by 27% than conventional column for methanol recovery

12

methanol-to-oil ratio on methanol recovery performance. Energy requirement was not

included in their work. Kiss and Ignat (2012) proposed a new configuration for the

distillation column, namely the dividing-wall column (DWC). They carried out an Aspen

Plus process simulation to determine methanol recovery performance and energy

requirements for the DWC and compared these with a conventional distillation column.

For the study of flash technology for methanol recovery, Tang et al. (2010)

performed a process simulation to investigate the effects of feed temperature, feed

pressure, and flash pressure on methanol recovery performance for a supercritical-based

biodiesel production process. Later in 2011, a similar parametric study was carried out by

Wang et al. (2011) through an experimental study of flash evaporation. Note that both

studies do not evaluate parametric effects on energy requirement of flash for methanol

recovery.

1.4 Research motivation, objectives, and scope of work

Due to the considerable quantities of alcohol required for biodiesel productions

that remain in the product stream, intensive energy requirements for alcohol recovery is

an important issue that directly impacts the production costs of biodiesel. As stated by

Atadashi et al. (2011), the cost of a downstream process of the biodiesel production

typically accounts for 60-80% of the process cost. Therefore, to reduce the cost of

biodiesel production, it is necessary to optimize the process and its energy requirement

for alcohol recovery. However, the knowledge of methanol recovery, particularly in the

aspect of energy requirement, which is required for the process and energy optimization,

is limited. Several works on parametric studies on energy requirement (as seen in Table

13

1.5) were reported for distillation, but none was found for flash and other possible

methanol recovery techniques such as stripping.

The objectives of this work are, therefore, to extend the knowledge of distillation-

, stripping-, and flash-based methanol recovery technologies and to recommend the most

energy-efficient technology for biodiesel production. To achieve such objectives, a series

of computational process simulations using Aspen Plus was carried out to evaluate

energy requirements of distillation, stripping, and flash processes used for methanol

recovery in the transesterification-based biodiesel production process. The data on energy

requirements were generated with respect to process parameters including percent

methanol recovery, pressure, reflux ratio, and methanol-to-oil ratio for each methanol

recovery unit (MRU). The energy data produced helped gain more understanding of the

effects of the process parameters on energy requirement. The energy requirements of

distillation-, stripping-, and flash-based MRUs were subsequently compared to determine

the most energy-efficient methanol recovery technology.

14

2. BACKGROUND AND LITERATURE REVIEW

2.1 Important reactions associated with biodiesel production

2.1.1 Transesterification reaction

Transesterification is defined as a class of organic reactions wherein one ester is

transformed into another ester by exchanging the alkoxy group of an ester by alcohol

(Demirbas, 2008). The transesterification of triglyceride to produce biodiesel occurs in

three consecutive, reversible reactions (Reactions 2.1-2.3) (Ma and Hanna, 1999;

Marchetti and Errazu, 2007). The triglyceride is first converted to diglyceride, then

monoglyceride, and eventually glycerine. In all three steps, esters (biodiesel) are

produced. The stoichiometric relation between alcohol and oil is 3:1, but excess alcohol

is usually supplied to ensure the completion of reactions towards the desired products

(Gerpen et al., 2004).

Triglyceride (TG) + R – OH ↔ Diglyceride (DG) + R – OOC – R1 (2.1)

Diglyceride (DG) + R – OH ↔ Monoglyceride (MG) + R – OOC – R2 (2.2)

Monoglyceride (MG) + R – OH ↔ Glycerol (GL) + R – OOC – R3 (2.3)

Catalysts are used for transesterification to accelerate chemical reactions by

reducing the activation energy, which is the energy needed to initiate the reaction. The

catalysts are classified into two types: homogeneous catalyst and heterogeneous catalyst.

The homogeneous catalysts can be base or acid catalyst. Common base catalysts are

sodium hydroxide (NaOH), potassium hydroxide (KOH), and sodium methoxide

(NaOCH3) while common acid catalysts are sulphuric acid (H2SO4) and hydrochloric

acid (HCl) (Vicente et al., 2004). The heterogeneous catalysts include enzymes, titanium

15

silicates, alkaline-earth metal compounds, anion exchange resins, etc. (Vicente et al.,

2004). Among these catalysts, base catalysts are most preferable for transesterification

using high quality oil containing negligible free fatty acids (FFAs) due to fast

transesterification rates, thereby offering high conversion rates to biodiesel. They,

however, perform poorly when the low-quality oil containing high FFA contents is used.

2.1.2 Esterification reaction

Low quality feedstock such as unrefined vegetables oil, waste cooking oil, and

grease contains FFAs. The FFAs in the presence of an acid catalyst react with alcohol to

produce alkyl ester and water. This reaction is known as esterification (Reaction 2.4)

(Gerpen et al., 2004).

R-COOH + CH3OH �� R-COOCH3 + H2O (2.4)

where R-COOH and R-COOCH3 represent free fatty acid and alkyl ester respectively.

The difference between transesterification and esterification is that transesterification can

be catalyzed by either an acid or base catalyst, but esterification can only be catalyzed by

an acid catalyst. Alkyl ester is formed in both esterification and transesterification

reactions in the presence of acid catalyst, which makes acid catalyzed systems insensitive

to free fatty acid.

2.1.3 Soap formation

If a base catalyst is used for transesterification, free fatty acids can react with the

catalyst in a neutralization reaction to form soap and water. Hence, in this reaction,

catalyst is consumed as the reactant as shown in Reaction 2.5 (Gerpen et al., 2004).

16

R-COOH + NaOH → R-COO-Na + H2O (2.5)

where R-COOH and R-COO-Na represent free fatty acid and soap, respectively. The

formation of soap causes a challenge for the separation of reaction products in biodiesel

production. This means a base-catalyzed process is sensitive to FFAs.

2.2 Biodiesel production processes

Biodiesel production processes can be classified into five categories according to

the way in which transesterification and esterification reactions are incorporated. Details

of these processes are given below.

2.2.1 Base-catalyzed transesterification process

Base-catalyzed transesterification is the most commonly used biodiesel

production process in which triglyceride is converted into alkyl esters (biodiesel) in a

single step using a base catalyst. For this process, high-quality feedstock containing free

fatty acid of less than 0.5% is necessary to avoid undesirable soap formation. Methanol

and ethanol are the typical alcohols used in the process, but methanol is preferable as it is

less expensive.

Figure 2.1 illustrates a simplified block diagram of the base-catalyzed

transesterification process. This process begins with the introduction of feedstock and a

mixture of alcohol and a catalyst to a transesterification reactor. The catalyst

concentration ranges from 0.3-1.5%. The amount of makeup alcohol fed into the process

is regulated to maintain the alcohol-to-triglyceride ratio between 4:1 and 20:1 (Gerpen et

al., 2004). The temperature of the reactor is maintained at 60 to 65°C. It takes approx.

17

Figure 2.1: Base-catalyzed transesterification process (modified from Gerpen et al., 2004

and “BIODIESEL IN CHEMCAD”, www.camstations.com)

Transesterification Reactor

Alcohol Recovery

Neutralization

Separator

Alcohol and Biodiesel distillation

Alcohol and Glycerol

distillation

Triglyceride

Alcohol + Base catalyst

Alcohol

Acid

Biodiesel

Glycerol

Alcohol

Alcohol

18

6 to10 minutes to convert triglyceride into biodiesel and glycerol. After reaction, the

product mixture is passed to an alcohol recovery unit to remove excess alcohol and

recycle it back. The product stream is then fed to the neutralization unit where the base

catalyst is neutralized by acid prior to the separation of biodiesel and glycerol in a

separator. After separation, both biodiesel and glycol product streams are purified in the

downstream units to remove the remaining alcohol.

2.2.2 Two-step process

Two-step process is suitable for feedstock containing free fatty acids (FFAs)

higher than 0.5%. The general concept of this process is to produce biodiesel via both

esterification and transesterification reactions in two sequential steps, i.e. esterification

and then transesterification. As illustrated in Figure 2.2, in the first step, an esterification

reaction is used to convert FFAs into alkyl esters, keeping the FFA content low prior to

the transesterification reaction. The esterification reaction in the first reactor occurs in the

presence of an acid catalyst, such as H2SO4, which is usually added into the process

through the incoming alcohol stream. The reactor is operated at 60-65°C with a residence

time of approximately one hour. After reaction, the acid catalyst is neutralized by a base,

and the water (by-product) is removed through a dryer to prevent poor conversion in the

transesterification step. After drying, the esterification product is fed to the second reactor

where an additional amount of alcohol may be added to the reactor to maintain a certain

ratio of alcohol to triglyceride. A base catalyst is fed to the reactor as a mixture of alcohol

and catalyst. The transesterification reaction is then carried out. All subsequent units in

this transesterification process are similar to those described previously in the base-

catalyzed transesterification process.

19

Figure 2.2: Two-step process (modified from Gerpen et al., 2004 and “BIODIESEL IN

CHEMCAD”, www.camstations.com)

Transesterification Reactor

Alcohol Recovery

Neutralization

Separator

Alcohol and Biodiesel distillation

Alcohol and Glycerol

distillation

Triglyceride

Alcohol + Acid catalyst

Acid

Biodiesel

Glycerol

Alcohol

Alcohol

Esterification Reactor

Neutralization

Dryer

Base

Alcohol + Base catalyst

Alcohol

20

2.2.3 Acid-catalyzed process

In this process, both esterification and transesterification reactions are promoted

to occur at the same time in a single step. Similar to the two-step process, the

esterification reaction is activated by an acid catalyst. Unlike the first two processes, the

transesterification of triglyceride in this process is driven by the same acid catalyst used

for esterification. This simply makes possible the parallel action of the two reactions.

This one-step process is suitable for feedstock that contains a large amount of FFAs.

Sulfuric acid, nitric acid, and phosphoric acid are the commonly used catalysts. The

alcohol-to-FFA ratio ranges from 20:1 to as high as 40:1 (Gerpen et al., 2004). Reaction

temperature is maintained at 60-65°C. As shown in Figure 2.3, after the reaction occurs,

the product stream is transferred to an alcohol recovery unit for the removal of excess

alcohol and then transferred to a neutralization unit. After neutralization, the product is

washed by water and transferred to a phase separator for the removal of wash water.

Finally, the biodiesel product is purified in an evaporation unit.

2.2.4 Biox process

The rate of biodiesel production is dependent upon the contact between the two

reactants (alcohol and oil feedstock). Greater contact results in faster production. In the

Biox process (Figure 2.4), better contact between the two reactants is achievable through

the use of another chemical referred to as a “co-solvent,” which is capable of dissolving

both alcohol and oil feedstock. The use of a co-solvent allows the two reactants to be

present in a homogeneous phase, thereby requiring no catalyst for biodiesel conversion.

In this process, the typical co-solvent known as tetrahydrofuran is added directly into the

21

Figure 2.3: Acid-catalyzed process (redrawn from Gerpen et al., 2004)

Figure 2.4: Biox co-solvent process (redrawn from Gerpen et al., 2004)

Triglyceride

Alcohol + Acid catalyst

Esterification Reactor

Neutralization

Base Alcohol

Alcohol Recovery

Separator Biodiesel

Washing

Evaporation

Waste water

Reactor Recovery

Unit

Separation Unit

Triglyceride

Alcohol

Glycerol

Biodiesel (Ester)

Co-solvent

Alcohol, Co-solvent

22

reactor where the reaction temperature is kept at 30°C and the residence time is

approximately 5-10 minutes. Because its boiling point is close to the boiling point of

alcohol, tetrahydrofuran can be recovered from the reaction product together with the

excess alcohol in a single step. Without the use of a catalyst, this process requires no

neutralization section. Despite its simplicity, this non-catalyzed process is subject to one

major shortcoming, i.e., the commonly used co-solvent is highly toxic. The cost of this

simple process would be considerably high due to the requirement for special equipment

and great care to prevent leakage of the co-solvent as well as to completely remove it

from product streams.

2.2.5 Supercritical process

In this process, reaction between triglyceride and alcohol occurs above the critical

point of the feedstock. The supercritical condition allows both reactants to form a single

phase of combined fluid, promoting biodiesel conversion without the presence of a

catalyst. The typical reaction temperature ranges from 350-400°C, and the reaction

pressure is maintained above 80 atm (Gerpen et al., 2004). This process requires a high

alcohol-to-oil ratio (about 42:1). Under proper operating conditions, the complete

conversion can be achieved within 3-5 minutes. To prevent product decomposition at a

high temperature and pressure, the product stream derived from the reactor must be

quenched very rapidly. This makes this non-catalyzed process energy intensive, thereby

requiring high capital and operating costs. A simplified block diagram of the supercritical

process is shown in Figure 2.5.

23

Figure 2.5: Supercritical transesterification process (redrawn from Gerpen et al., 2004)

Supercritical reactor

Alcohol recovery

unit

Separation unit

Triglyceride (Oil)

Alcohol (Methanol)

Alcohol

Biodiesel (Ester)

Glycerol

24

2.3 Literature review on methanol recovery

A number of previous works on methanol recovery from biodiesel production

processes were published in the literature and are summarized in Table 1.5. In 2009, Dhar

and Kirtonia carried out a simulation study of excess methanol recovery using a 10-stage

distillation unit in Aspen Plus. UNIFAC was chosen as the thermodynamic model. They

investigated the effects of alcohol-to-oil ratio, reflux ratio, and column pressure on

methanol recovery as well as energy consumption. The alcohol-to-oil ratio was varied

from 6:1 to 50:1 while the reflux ratio was varied from 1 to 4. The column pressures were

0.5, 0.75, and 1 atm. They concluded that energy requirement increases with methanol-

to-oil ratio and reboiler heat duty reduced under vacuum. Their study was conducted in

narrow ranges of parameters (pressure and reflux ratio). Only three specific pressures

were used at the reflux ratio of 1. The results could not be used to project how the

pressures (lower than 0.5 atm) and reflux ratios other than 1 will affect reboiler duty and

to what extent the pressure can be lowered to reduce heat duty. No results of purity of

recovered methanol were presented.

Baroutian et al. (2010) performed an experimental study of methanol recovery

where biodiesel was produced by base-catalyzed transesterification in a membrane

reactor. Distillation to recover methanol was carried out using three necks round bottom

flask, oil bath, and condenser. Influences of heating temperature, permeate flow rate, and

methanol-to-oil volume ratio on recovery of methanol with time were investigated.

Temperature was varied from 80 to 130°C, permeate (feed flow rate) was varied form 2.4

– 12.3 ml/min, and oil-to-alcohol ratios were 1:1, 1:1.5, and 1:2. They found that these

parameters had a significant effect on rate of methanol recovery. However, no

25

investigation was done to determine the effect of these parameters on energy

consumption.

Tang et al. (2010) carried out a process simulation study (using a mathematical

model) of methanol recovery by flash process in biodiesel production with the

supercritical method. Effects of feed temperature, feed pressure, and flash pressure on

methanol recovery were investigated. It was found that more than 85% methanol

recovery with a purity of 99% was possible at a feed pressure of 15 – 30 MPa and flash

pressure of 0.4 MPa. Methanol recovery performance for one-stage was close to two-

stage flash under the same feed conditions.

Wang et al. (2011) carried out an experimental study of methanol recovery by

flash process. They investigated effects of feed temperature, feed pressure, and flash

pressure on methanol recovery. They showed that at reaction pressure of 9–15 MPa and

reaction temperature of 240–300°C, flash pressure had a significant influence on

methanol recovery and methanol content in gas phase. The reaction temperature and

reaction pressure also affected methanol recovery and methanol content in gas phase. An

85% methanol recovery with 99% purity was possible at a feed pressure of 15 MPa, feed

temperature of 300°C, and flash pressure of 0.4 MPa. Note that the effects on energy

consumption were not investigated.

Kiss and Ignat (2012) used a new column configuration, namely the dividing wall

distillation column (DWC), for methanol recovery in the transesterification-based

biodiesel process. Aspen Plus and Aspen Dynamics were used to perform rigorous

steady state and dynamic simulations. They concluded that the proposed DWC required

27% less energy than a conventional column for methanol recovery.

26

3. METHODOLOGY

3.1 Process simulation procedures

Evaluation of methanol recovery units was carried out using process simulation

software Aspen Plus V7.1 developed by Aspen Technology Inc. (USA). Aspen Plus is a

computer-aided simulation software which uses the underlying physical relationships

(e.g., material and energy balances, thermodynamic equilibrium, rate equations) to

predict process performance (e.g., stream properties, operating conditions, and equipment

sizes). It can handle very complex processes, including multiple-column separation

systems, chemical reactors, distillation of chemically reactive compounds. Aspen Plus

can be used to perform sensitivity analyses, estimate and regress physical properties and

optimize processes. The following are the steps taken to implement process simulation:

• Draw complete graphical simulation process flow sheets, which involve placing

and labeling all unit operation models in the flowsheet and connecting all the

units using labeled streams.

• Specify all the required chemical components involved in the processes.

• Select appropriate thermodynamic models for all unit blocks to represent the

physical properties of the components and mixtures in the process.

• Provide thermodynamic parameters that can be retrieved from the Aspen Plus

database or can be input from other sources.

• Specify the operating conditions of all unit operations.

• Perform the simulation and model analysis.

27

3.2 Process flow schemes of simulated methanol recovery units

In this work, four types of methanol recovery units (MRUs), i.e., distillation, stripping,

single-flash, and double-flash, were evaluated. As illustrated in Figure 3.1, their process

flow schemes used during simulation consisted of two main process units:

transesterification unit and methanol recovery units (MRUs). Note that the flow scheme

of the transesterification unit follows Zhang et al. (2003) and Glisic and Skala (2009),

and was identical for all four process schemes. The product purification processes after

the MRU were not included in this simulation.

The transesterification unit starts with mixing a stream of fresh methanol (stream

101) with a catalyst (sodium hydroxide; NaOH) in Mixer 1. The mixture is then fed to

Mixer 2 by Pump 1 to combine with a stream of recycled methanol from the MRU. After

mixing, the mixture of methanol, NaOH, and recycled methanol is introduced to a reactor

where it mixes with a stream of vegetable oil (stream 102), which is preheated in a heater.

In the reactor, transesterification takes place and produces biodiesel and glycerol. The

reactor was set to operate at 333 K and 4 atm with 95% conversion of oil to biodiesel

(Zhang et al., 2003 and Glisic and Skala, 2009). The product stream (stream 104)

containing a mixture of biodiesel, glycerol, unreacted oil, and unreacted methanol is

subsequently introduced to the MRU.

Figure 3.1 (a) illustrates the process flow scheme of the distillation-based MRU.

The product stream (steam 104) from the transesterification reactor is fed to the distill

where the methanol is separated from the product stream by means of heat. The

recovered methanol (stream 105) leaves the distill from the top and passes through a

reflux condenser and Pump 3 before being sent to Mixer 2 for reuse in the

28

(a)

(b)

MIXER1

PUMP1

MIXER2

PUMP2

HEATER

REACTOR

DISTILL

PUMP3NAOH

101

101B

103

102

102C

104

105

106

RECYCLE

101A

102A

TO FURTHER PURIFICATION

MIXER1

PUMP1

MIXER2

PUMP2

HEATER

REACTOR

PUMP3

STRIPPER

COOLER

NAOH

101

103

102

102C

104

107

106

RECYCLE

101A

102A

105

101B


Transesterification unit MRU

MRU Transesterification unit

29

(c)

(d)

Figure 3.1: Process flow schemes of transesterification and methanol recovery units (a)

distillation-based (b) stripping-based (c) single flash-based (d) double flash-

based.

MIXER1

PUMP1

MIXER2

PUMP2

HEATER

REACTOR

PUMP3

COOLER

FLASH

NAOH

101

103

102

102C

104

107

106

RECYCLE

101A

102A

105

101B


MIXER1

PUMP1

MIXER2

PUMP2

HEATER

REACTOR

PUMP3

COOLER

NAOH

101

103

102

102C104

110

106

RECYCLE

101A

102A

105

101B


FLASH1

FLASH2

B3

107

108

109

MRU

Transesterification unit


MRU

30

transesterification unit. A mixture of biodiesel, glycerol, oil, and unrecovered methanol

(stream 106) leaves the bottom of the distill for further purification processes.

Figure 3.1 (b) illustrates the process flow scheme of the stripping-based MRU.

The product stream (stream 104) from the transesterification reactor is introduced to a

stripper where the methanol is separated from the product stream by means of heat. The

recovered methanol (stream 105) exits the stripper top, passes through a cooler for

temperature adjustment, and is eventually sent back to the transesterification unit. The

un-stripped mixture containing biodiesel, glycerol, oil, and unrecovered methanol (stream

106) leaves the bottom of the stripper for further purification processes.

Figure 3.1 (c) illustrates the process flow scheme of the single-flash unit. The

product stream (stream 104) from the transesterification reactor is fed to a flash drum

where a certain amount of the methanol is flashed out. This stream of the recovered

methanol (stream 105) leaves the flash drum from the top and passes through a cooler

prior to being recycled back to the transesterification unit. The un-flashed mixture

containing biodiesel, glycerol, oil, and unrecovered methanol (stream 106) leaves the

bottom of the flash for further purification processes.

Figure 3.1 (d) illustrates the process flow scheme of the double-flash unit. The

product stream (stream 104) from the transesterification reactor is fed to the first flash

drum (Flash 1) where the methanol is flashed out. The un-flashed mixture containing

biodiesel, glycerol, oil, and unrecovered methanol (stream 106) leaves Flash 1 and enters

the second flash drum (Flash 2) for further methanol recovery. The streams of recovered

methanol from Flash 1 and Flash 2 are then sent to a mixer (Mixer 3) and a cooler prior

to being recycled back to the transesterification unit.

31

3.3 Chemical components

Triolein (C57H104O6) was used in this study to represent vegetable oil. Methanol

(CH3OH) was taken as alcohol. The product biodiesel (FAME) was represented by oleic

acid methyl ester (C19H36O2). Glycerol (C3H8O3) was taken as by product.

Triolein (triglycerides) was chosen to represent vegetables oil because it is the

ester of major fatty acid (oleic acid) present in a variety of vegetable oils such as canola,

rapeseed, olive, palm, and peanut oil (Harding et al., 2007). It was also previously used in

the literature to represent vegetables oil. Zhang et al. (2003) used it to represent canola

oil, which is the major vegetable oil in Canada. It was also used to represent soybean oil,

which is the major domestic oil crop of the United States by Wang (2008) and Myint et

al. (2009). Glisic et al. (2009) and Lee et al. (2011) used triolein as the vegetable oil in

their simulation.

3.4 Process simulation

3.4.1 Simulation basis

Simulation was done based on 1 kmol/hr feed oil. Feed oil was taken as virgin

vegetable oil containing no free fatty acid (FFA). Throughout the simulation, the feed oil

flow rate was kept constant while the feed of fresh methanol was regulated to maintain

the desired methanol-to-oil ratio. It was assumed that no esterification reaction occurred,

and, hence, no water was produced as the product. For simplicity, it was also assumed

that the concentration of catalyst (NaOH) was zero during simulation.

32

3.4.2 Simulation framework

The simulation framework shown in Figure 3.2 consists of three main parts:

simulation inputs, calculation, and outputs. The simulation input requires three types of

information: operating parameters, physical property parameters, and type of

thermodynamic model. The operating parameters include temperature, pressure, reflux

ratio, distillate rate, feed stage, and total stage while the physical properties of the

components are molecular weight, boiling point, critical temperature, and critical

pressure. Type of thermodynamic model (UNIFAC-DMD) was also chosen here as an

input.

The model calculations involve a series of nonlinear equations representing

vapor-liquid equilibrium and unit operation models. The vapor-liquid equilibrium

calculation resulted in equilibrium constant (��), liquid activity coefficient ( ��), and

vapor fugacity coefficient (��). These are required for the calculation of mass and energy

balances for MRUs.

The simulation outputs provide information on flow rates and composition of

chemical components for all process streams and process units. Ultimately, these outputs

were used to analyze percent methanol recovery and heat duty per unit mass of methanol

recovered.

3.4.3 Simulation input

3.4.3.1 Operating parameters

Operating parameters are the variables that need to be specified to define the unit

model used in simulation. The required operating parameters for transesterification unit

33

Figure 3.2: Simulation framework for methanol recovery by methanol recovery units

Physical property parameter • Molecular weight ( MW) • Boiling point (TB) • Critical temperature (TC) • Critical pressure (PC) • Acentric factor (ω)

Operating parameter

• Temperature • Pressure • Reflux ratio • Distillate rate • Feed stage and Total stage

Type of Thermodynamic model

UNIFAC-DMD

Methanol recovery units (MRUs)

Vapor-liquid equilibrium model

• Liquid activity coefficient ( ��) • Vapor fugacity coefficient (��) • Equilibrium constant (��)

Output

• Flow rate of chemical components • composition of chemical

components • Energy requirement or Heat duty

of process units

Post analysis

• % methanol recovery • Heat duty per kg methanol

recovered

Input

Calculation

Output

Unit operation model

• Mass balance • Energy balance • equilibrium relation

• Summation


34

and methanol recovery units (MRUs) are listed in Tables 3.1 and 3.2, respectively. The

operating pressure and temperature for all process components in the transesterification

unit were specified at 4 atm and 333 K, respectively. These conditions are similar to the

operating conditions used in the works of Zhang et al. (2003) and Glisic and Skala

(2009). The molar ratio of methanol to oil ranges from 6:1 to 15:1.

For the MRUs the type of process parameters required as the model inputs is

dependent upon the type of methanol recovery units (MRUs). As listed in Table 3.2, the

distillation- and stripping-based MRUs require a similar set of operating parameters, i.e.,

column pressure, distillate rate, total stage, and feed stage, except that the distillation unit

requires the specification of condenser reflux ratio whereas the stripping unit does not.

The pressures of distillation and stripping columns are specified to be in the range of 0.1

to 1.0 atm. Such vacuum pressure was used to avoid decomposition of biodiesel and

glycerol (Zhang et al., 2003 and Morais et al., 2010). The single and double flash units

require the inputs of pressure and temperature in the range of 0.1 – 1.0 atm and 333 – 393

K, respectively.

3.4.3.2 Physical property parameters of chemical components

The required physical property parameters are molecular weight (MW), normal

boiling point (TB), critical temperature (TC), critical pressure (PC), acentric factor (ω),

liquid density (��) or liquid molar volume (VL), ideal gas heat capacity (CP,IG), and heat

of vaporization (∆Hvap). Their values for methanol, glycerol, and FAME were available in

the data bank of Aspen Plus, which integrates the database of the National Institute of

Science and Technology Thermo-Data Engine (NIST TDE). They are listed in Table 3.3.

35

Table 3.1: Operating parameters for transesterification unit

Process component Operating parameters Value

Pump 1, 2 and 3 Pressure (atm) 4

Mixer Pressure (atm) 4

Heater Temperature (K)

Pressure (atm)

333

4

Reactor Temperature (K)

Pressure (atm)

Conversion of oil into biodiesel (%)

Methanol to oil ratio (molar)

333

4

95

6:1 – 15:1

36

Table 3.2: Process parameters for methanol recovery units

Methanol recovery unit

(MRU)

Operating parameters Value

Distillation Column Pressure (atm) 0.1 – 1.0

Distillate rate (kmol/hr) 2.0 – 3.1a

Reflux ratio 1 – 4

Total stage 7

Feed stage 4b

Stripping Column Pressure (atm) 0.1 – 1.0

Distillate rate (kmol/hr) 2.0 – 3.1a

Total stage 7

Feed stage 1a

Single flash Pressure (atm) 0.1 – 1.0

Temperature (K) 333 – 393

Double flash Pressure (atm) 0.1 – 1.0

Temperature (K) 333 – 393

a This range of distillate rate was based on the methanol to oil molar ratio of 6:1. b Stage number was counted from top of the column.

37

Table 3.3: Scalar property of methanol, glycerol, and FAME (Aspen Plus databank V

7.1)

Parameter

Unit

Chemical component

Methanol Glycerol FAME

Molecular weight (MW) 32.042 92.094 296.493

Acentric factor (ω) 0.565 0.512 1.0494

Critical Pressure (PC) atm 79.782 74.019 12.632

Boiling Point (TB) K 337.85 561 617

Critical Temperature (TC) K 512.5 850 764

38

For triolein, its physical properties were not taken from the Aspen Plus data bank

but were predicted from the methods used in the literature. This is because the values of

the physical properties in Aspen Plus considerably differ from those found in the

literature. As shown in Table 3.4, the values of TB, TC, and PC from Aspen Plus are 1120

K, 1640 K, and 470 kPa, respectively while those from the literature are in the ranges of

824 – 879 K, 841 – 977 K, and 327 – 369 kPa, respectively.

To predict the values of TB, TC, and PC, the methods of Gani, Jobac, Mani, and

Ambrose are available in the Aspen Plus property estimation tools. The Gani method was

chosen for this study because it was reported to give reliable results compared to

experimental data (Chang and liu 2010). This method was also used in other previous

works including the works of Glisic et al. (2007) and Tang et al. (2010). The detailed

calculation of the Gani method can be found in Poling et al. (2001), and the predicted

data are tabulated in Table 3.5.

3.4.3.3 Type of thermodynamic model

A number of thermodynamic models to determine liquid activity coefficient (��) are

available as options in the process simulator. The ones found in the literature for the

simulation of biodiesel production processes are NRTL, UNIQUAC, and UNIFAC as

listed in Table 3.6. However UNIFAC was used for the prediction of missing binary

interaction parameter by some authors even though they used NRTL or UNIQUAC. The

Dortmund modified UNIFAC (UNIFAC-DMD) model was chosen in this study because

it can be used to predict missing binary interaction parameters in Aspen databank for

methanol-biodisel-glycerol systems, it was proven to provide reliable prediction results

39

Table 3.4: Properties of triolein found in the literature

Parameter Tang et al.

(2006)

Glisic et al.

(2007)

Chang and Liu.

(2010)

Aspen

Plus

TB (K) 879a 827.4b 824c 1120

TC (K) 954.1a 977.8b 841c 1640

PC (KPa) 360.2a 334b 327.02c 470

a Estimated by the method of Dohrn and Brunner (1991, 1994) b Estimated by the method of Constantinou and Gani (1994) c Estimated by NIST TDE integrated with Aspen

Table 3.5: Physical properties of triolein predicted by Gani method

Parameter Predicted value

TB (K) 822.5

TC (K) 943.3

PC (KPa) 322.4

40

Table 3.6: The thermodynamic models used by various authors for biodiesel process

simulations

Reference Thermodynamic model Software

Zhang et al. (2003) NRTL and UNIQUAC Hysys

Kasteren and Nissworo (2007) UNIQUAC Aspen Plus

West et al. (2007) NRTL Aspen Plus

West et al. (2008) NRTL and UNIFAC Hysys

Glisic and Skala (2009) UNIQUAC and UNNIFAC-LL Aspen Plus

Gutiérrez et al. (2009) NRTL and UNIFAC Aspen Plus

Kiwjaroun et al. (2009) NRTL and UNIQUAC Hysys

Morais et al. (2010) NRTL and UNIQUAC Aspen Plus

Santana et al. (2010) NRTL and UNIFAC Hysys

Sotoft et al. (2010) UNIFAC -DMD Aspen Plus

41

compared to experimental data, and it was also recommended by Kuramochi et al.

(2009). Negi et al. (2006) carried out experiments to determine liquid-liquid phase

equilibrium data for the glycerol-methanol-methyl oleate (biodiesel) ternary systems

at333 and 408 K and, then, compared the results with the prediction results of UNIFAC

and UNFAC-DMD models. They showed that both the predictive models were in good

agreement with the experimental data at 333 K, but there was deviation at 408 K. In

2009, Kuramochi et al. reported the UNFAC-DMD model was appropriate to represent

VLE of the methanol-biodiesel-glycerol systems.

As the UNIFAC-DMD model uses the group contribution method to predict

liquid activity coefficient, chemical components are divided into different functional

groups. The activity coefficient of components is then calculated by the interaction of the

functional groups (Fredenslund et al., 1975). Detailed equations to calculate the activity

coefficient by the UNIFAC-DMD model are given in Weidlich and Gmehling, (1987).

The functional groups used for this study are shown in Table 3.7.

3.4.4 Model calculations

3.4.4.1 Methanol recovery unit model

The following paragraphs provide concepts and calculation equations involved in

modeling of the studied MRUs, i.e., distillation-, stripping-, and flash-based MRUs. Note

that the concepts and equations for the stripping process are not given here. This is

because they are similar to those of the distillation process. The only difference in the

aspect of process features is that the distillation is equipped with a condenser at the

column top, but the stripper is not. The concepts and equations for the flash unit were

applied for both single and double flash processes.

42

Table 3.7: UNIFAC DMD group assignment for this study (Negi et al., 2006 and

Kuramochi et al., 2009)

Component UNIFAC DMD group assignment

Biodiesel 2 × CH3, 13 × CH2, 1 × CH=CH, CH2COO

Glycerol 1 × CH, 2 × CH2, 2 × OH(p), 1 × OH(s)

Methanol CH3OH

Triolein 3 × CH3, 41 × CH2, 1 × CH, 3 × CH2COO, 3 × CH=CH

p: primary alcohol, s: secondary alcohol

43

A) Distillation unit

The distillation unit used for methanol recovery in this study is the multi-component

distillation column with a cascade of trays starting from the top with a vapor condenser

and a reboiler for vapor boilup at the bottom, which itself acts as an equilibrium stage

(Figure 3.3). As illustrated in Figure 3.3(b), for each tray j, vapor and liquid phases are in

equilibrium with mole fractions of �� and ��, respectively. Each tray has vapor and

liquid flowing from it (�� and��) and is connected to streams above and below (�� and

��! ). The following are the distillation equations for each tray (Biegler et al., 1997):

Mass balance:

"�#�� + �� ,�� + ��! ��,�! − �� − �� = 0 (3. 1)

)ℎ+,+ � = 1, … / 012 3 = 1, … … 4

Equilibrium relation:

�� = �� (3. 2)

�� = � 56�, 7� , ��8 (3.3)

Summation equations:

∑ �� = 1 (3.4)

∑ �� = 1 (3.5)

)ℎ+,+ 3 = 1, … … 4

Heat balance:

:"�� + �� :�,�� + ��! :�,�! − ��:�� − ��:�� = 0 (3.6)

where "� is the feed flow rate at tray j, #�� is the feed composition, �� is the liquid flow

rate from the tray above j, ��,�� is the liquid mole fraction of component i in �� ,

44

(a)

(b)

Figure 3.3: Schematic diagram of distillation column (a) full column (b) tray of the

column (modified from Biegler et al. (1997) and Fredenslund et al. (1977)).

F j Tj Pj

L j V j + 1

L j – 1 V j

45

��! is the vapor flow rate from the tray below j, ��,�! is the mole fraction of component

i in ��! , �� is the liquid flowing out from tray j, �� is the mole fraction of component i

in ��, �� is the vapor flow rate from the tray j, �� is the mole the mole fraction of

component i in ��, �� is the equilibrium constant, 6� and 7� are column temperature and

pressure, respectively, and :� and :� are enthalpy of liquid and vapor at corresponding

tray temperature and pressure. These Mass, Equilibrium, Summation and Heat (MESH)

equations form the standard model for a tray distillation unit. There have been many

algorithms to solve the MESH system of equation. However a standard (Inside–out)

algorithm was chosen for simulation using Aspen Plus.

B) Flash unit

In the flash unit, a single inlet stream, is brought to the condition such that a liquid

phase and a vapor phase are developed and approach equilibrium in the vessel commonly

known as the flash drum or the separator. The multi-component, single stage, vapor-

liquid equilibrium flash is depicted in Figure 3.4. Similar to the distillation unit, the

following mass balance, phase equilibrium, and energy balance equations are required to

describe a single stage flash unit:

Mass balance:

� + � = " (3.7)

�� + �� = #�" (3.8)

Phase equilibrium:

�� = �� (3.9)

46

Figure 3.4: Schematic diagram of a single-stage flash unit (redrawn from Biegler et al. (1997))

47

Enthalpy balance:

:�� + :�� = :;" + < (3.10)

Summation:

∑ �� = ∑ �� (3.11)

where � is the liquid flow rate, � is the vapor flow rate, " is the feed flow rate, �� is the

mole fraction of component i in liquid phase, �� is the mole fraction of component i in

vapor phase, #� is the mole fraction of component i in the feed stream, �� is the

equilibrium constant, Q is the heat duty, and :�, :�, and :;are the enthalpy of liquid

vapor and feed streams, respectively.

If the feed quantities F,zi and HF are known, and �� and �� are to be calculated,

then there remain five quantities (T, P, Q, L, and V) that must be fixed in order to

completely describe this system. Since L and V are not independent, it can be assumed

that L is always a calculated quantity. Hence, any two of the variables may be specified

arbitrarily, and the other two are determined to satisfy the equations. Various numerical

algorithms are available for solving the described equation. Aspen Plus uses the Inside-

out algorithm proposed by Boston and Britt (1978).

3.4.4.2 Vapor-liquid equilibrium model

To model the methanol recovery units, which are essentially equilibrium

separation units, proper vapor-liquid equilibrium relation is vital. For the equilibrium

systems, the equilibrium relation between vapor and liquid phases can be expressed as:

�� = �� (3.12)

48

where �� is the mole fraction of component i in vapor phase, �� is the mole fraction of

component i in liquid phase, and �� is the equilibrium constant, which can be expressed

using standard thermodynamics as:

�� = =>;>?,@

A>B C (3.13)

where �� is the liquid activity coefficient of component i, D�E,� is the liquid fugacity of pure

component i at mixture temperature and pressure, �� is the fugacity coefficient at vapor

phase of component i, and P is the total pressure. For the condensable components, D�E,�is

calculated using the following:

D�E,� = 7�E (6) ��(�� = 1, 7�E , 6) exp IJ K>@ (L,C)

MLC

C>? 2NO (3.14)

where 7�E is the saturation (vapor) pressure of component i, �� is the fugacity coefficient

at saturation, and �� is the molar liquid volume at temperature of T. Only pure

component data are required to calculate D�E,�.

3.4.5 Simulation Outputs

The simulation of the transesterification unit connected with the MRUs results in

a series of data including flow rate and composition of chemical components in all

streams and energy requirement or heat duty for the MRUs as well as other process units.

These data were subsequently used in post analysis to determine percent methanol

recovered and heat duty per kg methanol recovered for the MRUs. They were also

analyzed in terms of how heat duty depends on process operating parameters. The

equations used are given below:

49

% methanol recovery = \] �� ^_�]`_^_ a_ b�c]� (de/b^)\] �� a_ b�c]� g__ ] hij� (de/b^) × 100 (3.15)

Heat duty (MJ/hr) = q_� r � (qsj\t)×�� × uvww �

b^ × x. zx { �� × h{

w|{ (3.16)

heat duty per kg methanol recovered �� = q_� r � (h{/b^)

\] �� ^_�]`_^_ a_ b�c]� (de/b^) (3.17)

50

4. SIMULATION RESULTS AND DISCUSSION

This chapter provides simulation results of the energy requirement (also referred

to as heat duty) of four types of methanol (MeOH) recovery units (MRUs), i.e.,

distillation, stripping, single-flash, and double-flash. The heat duty for each MRU was

analyzed in terms of parametric effects reported as a function of process parameters

including percentage of MeOH recovery, operating pressure of MRU, methanol-to-oil

ratio, and reflux ratio (in the case of the distillation-based MRU). The quality of

recovered MeOH from biodiesel products for each MRU was also analyzed and reported

here. The heat duties of all four MRUs were subsequently compared.

4.1 Distillation-based MRU

Distillation is a unit operation designed to separate a liquid mixture based on

difference in volatility or boiling point of components. In general, a distillation unit is

composed of three important components: a distillation column separating lighter

components from heavier, a reboiler supplying heat for operation, and a reflux condenser

converting distilled vapor into liquid phase (referred to as “top product”). In distillation

operations, a portion of the top product is reintroduced back to the top of the distillation

column to promote product rectification. The heat duty data for distillation were derived

from a 7-stage column where the feed at 60°C and MeOH-to-oil ratio of 6 was introduced

on the 4th stage (the middle of the column).

51

4.1.1 Parametric effect on heat duty

A) Percentage of MeOH recovery

Figure 4.1 presents reboiler heat duty at five different operating pressures and

three reflux ratios as a function of percentage of MeOH recovery. In general, an increase

in the degree of MeOH recovery requires more energy to separate one kilogram of MeOH

from the feed mixture. The increasing energy requirement could be divided into two

regions: moderate recovery (less than 90%) and high recovery (greater than 90%). In the

moderate recovery region, a change in reboiler heat duty due to an increase in MeOH

recovery is relatively small. Regardless of reflux ratio and operation pressure, raising the

recovery from 75 to 90% results in an increase in heat duty of only 0.5 MJ/kg MeOH. In

the high recovery region, the change in reboiler heat duty is much more significant, and

its magnitude does depend upon the operating pressure. An increase in MeOH recovery

from 90 to 98% requires additional heat energy of 1.2, 1.5, 2.0, 3.7, and 5.9 MJ/kg for

operation pressures of 0.1, 0.2, 0.3, 0.5, and 1.0 atm, respectively. The increasing reboiler

heat duty with percentage of MeOH recovery is essentially controlled by the vapor liquid

equilibrium (VLE) feature of the MeOH system or vapor pressure of MeOH in particular.

As the percentage of MeOH recovery increases, the MeOH content of vapor phase within

the distillation column also increases. For a given operating pressure, this results in an

increase in vapor-phase partial pressure of MeOH during the distillation. The increasing

partial pressure can hinder mass transfer process that proceeds to extract MeOH from the

liquid feed mixture under specific conditions. To promote more mass transfer activity as

well as MeOH recovery, additional heat energy must be supplied through the reboiler.

52

(a)

(b)

(c)

Figure 4.1: Effect of % MeOH recovery on reboiler heat duty for the distillation-based

MRU at the reflux ratio of (a) 2 (b) 3 (c) 4 (column pressure = 0.1-1.0 atm,

total stage = 7, feed stage = 4, MeOH to oil ratio = 6, feed temperature = 60

oC and feed pressure = 4 atm).

2

4

6

8

10

12

14

70 75 80 85 90 95 100R

eboi

ler

heat

dut

y (M

J/kg

)

% MeOH recovery

0.1 atm0.2 atm0.3 atm0.5 atm1.0 atm

2

4

6

8

10

12

14

70 75 80 85 90 95 100

Reb

oile

r he

at d

uty

(MJ/

kg)

% MeOH recovery


2

4

6

8

10

12

14

70 75 80 85 90 95 100

Reb

oile

r he

at d

uty

(MJ/

kg)

% MeOH recovery


53

B) Operating pressure of distillation unit

Figure 4.2 demonstrates how the operating pressure of a distillation column has

an impact on reboiler heat duty for 80 – 98% MeOH recovery. It is apparent that the

reboiler heat duty increases with operating pressure. The effect of pressure is due to the

VLE feature of the MeOH system as described in the previous section. The change in

reboiler heat duty becomes more significant as the degree of MeOH recovery increases.

For instance, at a reflux ratio of 2 (Figure 4.2-a), raising the column pressure from 0.1 to

1.0 atm results in an increase in heat duty by only 1.5 MJ/kg in order to recover 80% of

MeOH. Such increase in heat duty could reach approximately 2 times (3.2 MJ/kg) and

more than 4 times (6.7 MJ/kg) when the MeOH recovery of 95 and 98%, respectively, are

the operation targets. This behaviour is true for all reflux ratios.

C) Reflux ratio

Figure 4.3 illustrates the effect of reflux ratio on reboiler heat duty for recovering

MeOH by distillation. It is clear that an increase in reflux ratio causes a linear increase in

reboiler heat duty regardless of percentage of MeOH recovery and operating pressure. An

increase in the reflux ratio from 1 to 4 leads to an increase in heat duty by 3.7 MJ/kg and

by 3.2 MJ/kg for vacuum operation at 0.1 atm and atmospheric operation at 1.0 atm,

respectively. Based on the general concept of process operation, raising the reflux ratio

causes more heat energy to be withdrawn from the distillation unit at the column top. To

compensate such energy loss, additional heat energy must be supplied to the column

through the reboiler. Thus, operating the distillation column at a lower reflux ratio

requires lower heat energy to recover the same amount of MeOH. However, lowering the

54

(a)

(b)

(c)

Figure 4.2: Effect of pressure on reboiler heat duty for the distillation-based MRU at the

reflux ratio of (a) 2 (b) 3 (c) 4 (total stage = 7, feed stage = 4, MeOH to oil

ratio = 6, feed temperature = 60 oC and feed pressure = 4 atm).

2

4

6

8

10

12

14

0.0 0.2 0.4 0.6 0.8 1.0R

eboi

ler

heat

dut

y (M

J/kg

)

Pressure (atm)

80% MeOH recovery85% MeOH recovery90% MeOH recovery95% MeOH recovery98% MeOH recovery

2

4

6

8

10

12

14

0.0 0.2 0.4 0.6 0.8 1.0

Reb

oile

r he

at d

uty

(MJ/

kg)

Pressure (atm)


2

4

6

8

10

12

14

0.0 0.2 0.4 0.6 0.8 1.0

Reb

oile

r he

at d

uty

(MJ/

kg)

Pressure (atm)


55

(a) (b)

(c) (d)

(e)

Figure 4.3: Effect of reflux ratio on reboiler heat duty for the distillation-based MRU at

the column pressure of (a) 0.1 atm (b) 0.2 atm (c) 0.3 atm (d) 0.5 atm (e) 1.0

atm (total stage = 7, feed stage = 4, MeOH to oil ratio = 6, feed temperature =

60 oC and feed pressure = 4 atm).

0

2

4

6

8

0.0 1.0 2.0 3.0 4.0

Reb

oile

r he

at d

uty

(MJ/

kg)

Reflux ratio

80% MeOH recovery85% MeOH recovery90% MeOH recovery95% MeOH recovery98% MeOH recovery 0

2

4

6

8

0.0 1.0 2.0 3.0 4.0

Reb

oile

r he

at d

uty

(MJ/

kg)

Reflux ratio


0

2

4

6

8

10

0.0 1.0 2.0 3.0 4.0

Reb

oile

r he

at d

uty

(MJ/

kg)

Reflux ratio


0

2

4

6

8

10

0.0 1.0 2.0 3.0 4.0

Reb

oile

r he

at d

uty

(MJ/

kg)

Reflux ratio


0

2

4

6

8

10

12

14

0.0 1.0 2.0 3.0 4.0

Reb

oile

r he

at d

uty

(MJ/

kg)

Reflux ratio


56

reflux ratio below the hydrodynamic limit could result in column-dry-up operation.

D) MeOH-to-oil ratio

MeOH-to-oil ratio is an important process parameter for biodiesel production as it

has a significant impact on the rate of biodiesel conversion. A higher MeOH-to-oil ratio

offers a rapid biodiesel reaction that could lead to a reduction in the necessary size of the

biodiesel reactor. However, it tends to require more energy for recovering the unused

MeOH after reactions. In typical biodiesel plants, an MeOH-to-oil ratio of at least 6 is

required so as to allow the conversion reaction to proceed at a reasonable rate. Therefore,

ratios ranging from 6 to 15 are of interest in this investigation.

Figure 4.4 shows the effect of MeOH-to-oil ratio on reboiler heat duty presented

in terms of energy requirement per unit mass of MeOH recovered. The reported heat duty

data were generated for three reflux ratios, i.e., 2, 3, and 4, and five different degrees of

MeOH recovery, i.e., 80, 85, 90, 95, and 98%. It is apparent that an increase in MeOH-to-

oil ratio leads to a non-linear reduction in reboiler heat duty. Increasing the ratio from 6

to 9 results in a significant reduction in heat duty, i.e., 41 – 50% drop for 98% MeOH

recovery, 28 – 40% drop for 95% recovery, and 14 – 21% drop for less than 90%

recovery. The reduction in reboiler heat duty is rather small when the MeOH-to-oil ratio

exceeds a ratio of 9. Although the higher MeOH-to-oil ratio offers the lower energy

requirement for recovering one kg of MeOH, it should be kept in mind that the total

energy consumed by the MeOH recovery process (in MJ unit) must be assessed together

with the total mass of MeOH to be recovered.

57

(a)

(b)

(c)

Figure 4.4: Effect of MeOH-to-oil ratio on reboiler heat duty for the distillation-based

MRU at the reflux ratio of (a) 2 (b) 3 (c) 4 (column pressure = 0.2 atm, total

stage = 7, feed stage = 4, feed temperature = 60 oC and feed pressure = 4

atm).

0

2

4

6

8

10

12

14

3 6 9 12 15R

eboi

ler

heat

dut

y (M

J/kg

)

MeOH to oil ratio


0

2

4

6

8

10

12

14

3 6 9 12 15

Reb

oile

r he

at d

uty

(MJ/

kg)

MeOH to oil ratio


0

2

4

6

8

10

12

14

3 6 9 12 15

Reb

oile

r he

at d

uty

(MJ/

kg)

MeOH to oil ratio


58

4.1.2 Quality of recovered MeOH

Quality of product streams is an important measure that can be used to reveal how

well the MeOH recovery unit works. It accounts for the MeOH content of both top and

bottom products derived from the distillation column. Figure 4.5 shows the mole

fractions of MeOH in the product as a function of reboiler heat duty at three different

reflux ratios. The figure shows that the mole fraction of MeOH in the top product is

constant at approximately 1.0 regardless of reboiler heat duty. This suggests the

distillation technique is capable of producing a recovered MeOH stream with high purity

even at a MeOH recovery target of 98%. Unlike the top product, the mole fraction of

MeOH in the bottom product depends on the magnitude of reboiler heat duty. A greater

amount of MeOH remains in the bottom product (or oil-phase product) when a lower heat

duty is supplied. For instance (see Figure 4.5-a), as high as 20% MeOH was found at the

column bottom when a heat duty of 3.1 MJ/kg was applied. As more heat energy was

introduced, the MeOH content of the bottom product kept reducing until it reached a

mole fraction of nearly 0.0.

4.2 Stripping-based MRU

Stripping is a physical separation process designed to remove one or more lighter

components from a liquid mixture by means of heat or the use of carrier gas. Process

components of the stripping unit are rather similar to those of the distillation unit. They

include a stripping column for component separation, a reboiler for heat supply, and a

condenser for vapor condensation. The difference between the stripping and distillation

units is that the condensed stream (or “top product”) of the stripping unit is not circulated

59

(a)

(b)

(c)

Figure 4.5: Quality of recovered MeOH from the distillation-based MRU at the reflux

ratio (a) 2 (b) 3 (c) 4 (column pressure = 0.2 atm, total stage = 7, feed stage =

4, MeOH to oil ratio = 6, feed temperature = 60 oC and feed pressure = 4

atm).

0.0

0.2

0.4

0.6

0.8

1.0

2 4 6 8M

ole

frac

tion

of M

eOH

Reboiler heat duty (MJ/kg)

at topat bottom

98% MeOH recovery

0.0

0.2

0.4

0.6

0.8

1.0

2 4 6 8

Mol

e fr

acti

on o

f MeO

H


at topat bottom

98% MeOH recovery

0.0

0.2

0.4

0.6

0.8

1.0

2 4 6 8

Mol

e fr

acti

on o

f MeO

H


at topat bottom

98% MeOH recovery

60

back to the column top. In other words, there is no reflux operation for the stripping unit.

The heat duty data for stripping in this work were derived from a 7-stage column where a

feed at 60°C and MeOH-to-oil ratio of 6 was introduced on the 1st stage (top of the

column).

4.2.1 Parametric effects on heat duty


Figure 4.6 illustrates the relationship of the reboiler heat duty of a stripping

column and the percentage of MeOH recovery. The heat duty data were obtained for five

operating pressures, i.e., 0.1, 0.2, 0.3, 0.5, and 1.0 atm. Similar to the results for the

distillation operation, the higher percentage of MeOH recovery demands more energy

supply for the stripping operation. Only small increments in reboiler heat duty are

required for increasing MeOH recovery performance up to 95%. Beyond this point, the

stripper-based MRU requires a significant increment in heat duty to achieve a recovery

target. The increasing reboiler heat duty is controlled by VLE or the vapor pressure of the

MeOH system.

B) Operating pressure of stripping

Figure 4.7 illustrates that the operating pressure of the stripping column has an

impact on reboiler heat duty. Similar to distillation operation, a higher reboiler heat duty

is required as the operating pressure increases to achieve a MeOH recovery target. The

effect of operating pressure becomes more pronounced when the degree of MeOH

recovery is increased. Changing the column pressure from 0.1 to 1.0 atm leads to an

61

Figure 4.6: Effect of % MeOH recovery on reboiler heat duty for the stripping-based

MRU (column pressure = 0.1-1.0 atm, total stage = 7, feed stage = 1(top),

MeOH to oil ratio = 6, feed temperature = 60 oC and feed pressure = 4 atm).

0

2

4

6

8

10

50 60 70 80 90 100

Reb

oile

r he

at d

uty

(MJ/

kg)

% MeOH recovery


62

Figure 4.7: Effect of operating pressure on reboiler heat duty for the stripping-based

MRU (total stage = 7, feed stage = 1 (top), MeOH to oil ratio = 6, feed

temperature = 60 oC and feed pressure = 4 atm).

0

2

4

6

8

10

0.0 0.2 0.4 0.6 0.8 1.0

Reb

oile

r he

at d

uty

(MJ/

kg)

Pressure (atm)


63

increase in reboiler heat duty by 2.0, 3.0, and 6.2 MJ/kg for MeOH recovery of 80, 95,

and 98%, respectively.

C) MeOH-to-oil ratio

Figure 4.8 illustrates the effect of MeOH-to-oil ratio on reboiler heat duty for the

stripping process as a function of operating pressure and degree of MeOH recovery. In

general, an increase in MeOH-to-oil ratio leads to a non-linear reduction in reboiler heat

duty. For instance, at 0.2 atm (Figure 4.8-a), the heat duty for 98% recovery decreases by

44% when MeOH-to-oil ratio increases from 6 to 9, and decreases by only 17% when the

ratio increases from 9 to 12. Figure 4.8 also shows that this non-linear effect becomes

more pronounced as the operating pressure increases. At the higher pressure of 1.0 atm

(Figure 4.8-c), the reduction in reboiler heat duty for 98% recovery could reach as high as

63% as MeOH-to-oil ratio increases from 6 to 9. However, the heat duty decreases by

only 30% and 17% when the MeOH-to-oil ratio increases further form 9 to12 and from

12 to 15, respectively.


Figure 4.9 shows mole fractions of MeOH in the top and the bottom products

derived from the stripping operation. Similar to the distillation, the stripping unit is

capable of producing a high purity of MeOH stream (approx. 100%) regardless of

reboiler heat duty (even at the high reboiler heat duty needed to recover 98% MeOH).

The MeOH content of the bottom product decreases (reflecting greater percent MeOH

recovery) as more energy is introduced through the reboiler. It should be noted that such

64

(a)

(b)

(c)

Figure 4.8: Effect of MeOH-to-oil ratio on reboiler heat duty for the stripping-based

MRU at the pressure of (a) 0.2 atm (b) 0.5 atm (c) 1.0 atm (total stage = 7,

feed stage = 1 (top), feed temperature = 60oC and feed pressure = 4 atm).

0

2

4

6

8

10

3 6 9 12 15R

eboi

ler

heat

dut

y (M

J/kg

)

MeOH to oil ratio


0

2

4

6

8

10

3 6 9 12 15

Reb

oile

r he

at d

uty

(MJ/

kg)

MeOH to oil ratio


0

2

4

6

8

10

3 6 9 12 15

Reb

oile

r he

at d

uty

(MJ/

kg)

MeOH to oil ratio


65

(a)

(b)

(c)

Figure 4.9: Quality of recovered MeOH from the stripping-based MRU at the pressure of

(a) 0.2 atm (b) 0.5 atm (c) 1.0 atm, (total stage = 7, feed stage = 1 (top),

MeOH to oil ratio = 6, feed temperature = 60 oC and feed pressure = 4 atm).

0.0

0.2

0.4

0.6

0.8

1.0

0 1 2 3 4M

ole

frac

tion

of M

eOH


at topat bottom

98% MeOH recovery

0.0

0.2

0.4

0.6

0.8

1.0

0 2 4 6 8

Mol

e fr

acti

on o

f MeO

H


at topat bottom

98% MeOH recovery

0.0

0.2

0.4

0.6

0.8

1.0

0 2 4 6 8 10

Mol

e fr

acti

on o

f MeO

H


at bottomat top

98% MeOH recovery

66

reduction in the MeOH content in the bottom product becomes more significant at higher

pressures.

4.3 Single-flash-based MRU

Flash is a liquid-vapor separation technique where a liquid mixture undergoes a

reduction in pressure by passing through a throttling device located inside a flash drum.

The reduction in pressure allows a portion of lighter components in the feed mixture

(MeOH in this case) to flash into a vapor, making the remaining liquid richer in the

heavier components (biodiesel product in this case). This separation technique can be

achieved through either single-step pressure reduction or multiple-step reduction.

4.3.1 Parametric effects on heat duty


Figure 4.10 shows heat duty for recovering MeOH from the liquid mixture using a

single-flash drum. Similar to the distillation and stripping operations, energy requirement

for MeOH recovery increases with the MeOH recovery target. Apparently, the change in

the heat duty is relatively small for the MeOH recovery targets below 90%, but rather

significant for high recovery targets greater than 90%. The exponential increase in heat

duty is more pronounced at higher pressures.

B) Operating pressure

Figure 4.11 illustrates the relationship of operating pressure of flash drum and

heat duty for MeOH recovery. An increase in the pressure causes the heat duty to rise.

67

Figure 4.10: Effect of % MeOH recovery on heat duty for the single-flash-based MRU

(operating pressure = 0.1atm - 1.0 atm, MeOH to oil ratio = 6, feed

temperature = 60 °C and feed pressure = 4.0 atm).

0

1

2

3

4

5

6

60 70 80 90 100

Hea

t dut

y (M

J/kg

)

% MeOH recovery


68

Figure 4.11: Effect of pressure on heat duty for the single-flash-based MRU (MeOH to

oil ratio = 6, feed temperature = 60 °C and feed pressure = 4.0 atm).

0

1

2

3

4

5

0.0 0.2 0.4 0.6 0.8 1.0

Hea

t dut

y (M

J/kg

)

Pressure (atm)

80% MeOH recovery85% MeOH recovery90% MeOH recovery95% MeOH recovery

69

Changing the operating pressure from 0.1 to 1.0 atm leads to an increase in heat duty by

1.8 MJ/kg, 2.0 MJ/kg, and 3.1 MJ/kg for MeOH recovery of 80%, 90%, and 95%,

respectively.


Figure 4.12 illustrates the behaviour of heat duty for a single-flash drum with

respect to MeOH-to-oil ratio. The heat duty data were generated as a function of

operating pressure and degree of MeOH recovery. It is apparent that an increase in the

MeOH-to-oil ratio leads to a non-linear reduction in heat duty. For instance, at 0.5 atm

(Figure 4.12-b), the heat duty for 98% recovery decreases by 53% when the MeOH-to-oil

ratio increases from 6 to 9 and decreases by only 24% when the ratio increases from 9 to

12. Figure 4.12 also shows that this non-linear effect becomes more prominent as the

operating pressure increases. At the higher pressure of 1.0 atm (Figure 4.12-c), the

decrease in reboiler heat duty for 98% recovery could reach as high as 61% as MeOH-to-

oil ratio increases from 6 to 9. However, the heat duty decreases by only 32% and 17%

when the MeOH-to-oil ratio increases further from 9 to12 and from 12 to 15,

respectively.


Figure 4.13 presents the mole fractions of MeOH in both top and bottom products

derived from the single-flash unit as a function of heat duty at three different operating

pressures. At 0.2 atm (Figure 4.13-a), mole fraction of MeOH in the top product is found

to be constant at approximately 1.0 regardless of heat duty. However, at 0.5 atm (Figure

70

(a)

(b)

(c)

Figure 4.12: Effect of MeOH-to-oil ratio on heat duty for the single-flash-based MRU at

the pressure of (a) 0.2 atm (b) 0.5 atm (c) 1.0 atm (feed temperature = 60 °C

and feed pressure = 4.0 atm).

0

2

4

6

8

10

3 6 9 12 15H

eat

du

ty (

MJ/

kg

)MeOH to oil ratio

90% MeOH recovery95% MeOH recovery98% MeOH recovery

0

2

4

6

8

10

3 6 9 12 15

Hea

t d

uty

(M

J/k

g)

MeOH to oil ratio


0

2

4

6

8

10

3 6 9 12 15

Hea

t d

uty

(M

J/k

g)

MeOH to oil ratio


71

(a)

(b)

(c)

Figure 4.13: Quality of recovered MeOH by the single-flash-based MRU at the pressures

of (a) 0.2 atm (b) 0.5 atm (c) 1.0 atm (MeOH to oil ratio = 6, feed


0.0

0.2

0.4

0.6

0.8

1.0

0 2 4 6M

ole

frac

tion

of

MeO

H

Heat duty (MJ/kg)

at topat bottom 98% MeOH

recovery

0.0

0.2

0.4

0.6

0.8

1.0

0 2 4 6

Mol

e fr

acti

on o

f M

eOH

Heat duty (MJ/kg)

at topat bottom

94% MeOH recovery

0.0

0.2

0.4

0.6

0.8

1.0

0 2 4 6

Mo

le f

ract

ion

of

MeO

H

Heat duty (MJ/kg)

at topat bottom

91% MeOHrecovery

72

4.13-b), the mole fraction of MeOH in the top product starts to decrease when the heat

duty is increased beyond approximately 2.9 MJ/kg corresponding to 94% MeOH

recovery. A similar behaviour is also observed at 1.0 atm (Figure 4.13-c). That is, the

mole fraction of MeOH in the top product drops from approximately 1.0 when the heat

duty is greater than approximately 3.3 MJ/kg (91% MeOH recovery). This suggests that

the purity of the recovered MeOH deteriorates with increasing heat duty. The increasing

heat duty not only flashes off MeOH, but also other chemical components in the bottom

products. This implies that it would be more difficult compared to distillation and

stripping for flash operations to achieve MeOH recovery of as high as 98%, especially at

higher pressure, while maintaining approximately 100% purity of recovered MeOH.

4.4 Double-flash-based MRU

Separation of liquid mixture by flash technique can be achieved either by single-

step pressure reduction or multiple-step reduction. For the multiple-step pressure

reduction, flash units can be arranged in series or cascade. In this section, the energy

requirement for MeOH recovery using two-step pressure reduction in the double-flash

unit is studied. The two flash units are connected in series, where a feed at 60°C, 4 atm,

and methanol-to-oil ratio of 6 is introduced to the first-flash unit, and the un-flashed

liquid mixture containing unrecovered methanol is then introduced to the second-flash

unit at a lower pressure.

4.4.1 Parametric effect on heat duty

73


Figure 4.14 illustrates the relationship of the heat duty of double-flash based

MRU and percentage of MeOH recovery. In general, an increase in the degree of MeOH

recovery requires more energy to separate MeOH from the product mixture. Similar to

the results of the single-flash operation, only small increments in heat duty are required

for increasing MeOH recovery performance up to 95%. Beyond this point, an increment

of heat duty becomes significant to achieve a given MeOH recovery target. Such

increment is more prominent for operation at lower pressures.

B) Operating pressure

Figure 4.15 represents the heat duty as a function of operating pressure for

methanol recovery of 95%. The flash pressure is reduced in two stages: a first stage with

an operating pressure of 0.2, 0.5, 1.0 and 2.0 atm and a second–stage with an operating

pressure of either 0.1 atm (Figure 4.15-a) or 0.2 atm (Figure 4.15-b). The figure shows

that the heat duty is not sensitive to the change in first-stage flash pressure. For instance,

at the second-stage pressure of 0.1 atm (Figure 4.15-a), the heat duty remains constant

around 1.48 MJ/kg regardless of the first-stage flash pressure.


Figure 4.16 shows the effect of MeOH-to-oil ratio on heat duty presented in terms

of energy requirement per unit mass of MeOH recovered. The reported heat duty data

were generated for three different degree of MeOH recovery i.e., 90, 95, and 98%, at a

first-stage pressure of 1.0 atm and a second stage pressure of 0.5 atm. Similar to the

74

Figure 4.14: Effect of percentage of MeOH recovery on heat duty for the double-flash-

based MRU with the first-stage pressure of 1.0 atm and the second-stage

pressure of 0.1, 0.2, 0.3, and 0.5 atm. (MeOH to oil ratio = 6, feed


0

1

2

3

4

5

6

75 80 85 90 95 100

Hea

t dut

y (M

J/kg

)

% MeOH recovery

1.0/0.5 atm1.0/0.3 atm1.0/0.2 atm1.0/0.1 atm

First/second stage pressure:

75

(a)

(b)

Figure 4.15: Effect of first–stage flash pressure on heat duty for the double-flash-based

MRU with the second–stage flash pressure of (a) 0.1 atm and (b) 0.2 atm.

(MeOH to oil ratio = 6, feed temperature = 60 °C and feed pressure = 4.0

atm).

0

1

2

3

0.0 0.5 1.0 1.5 2.0 2.5

Hea

t d

uty

(M

J/k

g)

First-stage flash pressure (atm)

95 % MeOH recovery

0

1

2

3

0.0 0.5 1.0 1.5 2.0 2.5

Hea

t d

uty

(M

J/k

g)

First-stage flash pressure (atm)

95% MeOH recovery

76

Figure 4.16: Effect of MeOH-to-oil ratio on heat duty for the double-flash-based MRU at

the first-stage pressure of 1.0 atm and the second-stage pressure of 0.5 atm.

(feed temperature = 60 °C and feed pressure = 4.0 atm).

0

1

2

3

4

5

6

3 6 9 12 15

Hea

t d

uty

(M

J/k

g)

MeOH to oil ratio


77

distillation, stripping, and single-flash operations, an increase in MeOH-to-oil ratio leads

to a non-linear reduction in heat duty. Increasing the ratio from 6 to 9 results in a

significant reduction in heat duty, i.e., 53, 42, and 38% drops for 98, 95, and 90% MeOH

recovery, respectively. Note that although higher MeOH-to-oil ratio offers lower energy

requirement for recovering one kg of MeOH, the total energy (MJ) consumed by the

MRU must be assessed together with the total mass of MeOH to be recovered.


Figure 4.17 presents the mole fractions of MeOH in both the top and the bottom

products derived from the double-flash drum as a function of heat duty at three different

operating pressures. A similar behaviour to the single-flash operation can be observed

here. At the first-stage/second-stage pressure of 1.0/0.1 atm (Figure 4.17-a), the purity of

recovered MeOH in the top product is approx. 100% regardless of heat duty. However, at

pressures of 1.0/0.2 atm and 1.0/0.5 atm (Figure 4.17-b, c), the purity of the recovered

MeOH in the top product decreases and the MeOH recovery performance cannot achieve

greater than 98% and 96% recovery when the heat duty is increased beyond

approximately 2.9 MJ/kg and 3.5 MJ/kg, respectively. This suggests that the double-

flash, similar to single-flash, has the capacity to produce recovered MeOH with high

purity (approx. 100%) but % MeOH recovery may not be great. For instance (Figure

4.17-c), at higher pressure such as 1.0/0.5 atm, it is not feasible to obtain pure recovered

MeOH while achieving MeOH recovery above 96%.

78

(a)

(b)

(c)

Figure 4.17: Quality of recovered MeOH by the double-flash-based MRU at the

first/second-stage pressures of (a) 1.0/0.1 atm (b) 1.0/0.2 atm (c) 1.0/0.5 atm (MeOH to

oil ratio = 6, feed temperature = 60 °C and feed pressure = 4.0 atm).

0.0

0.2

0.4

0.6

0.8

1.0

0 1 2 3 4 5 6M

ole

frac

tion

of M

eOH

Heat duty (MJ/kg)


recovery

0.0

0.2

0.4

0.6

0.8

1.0

0 1 2 3 4 5 6

Mol

e fr

acti

on o

f MeO

H

Heat duty (MJ/kg)


recovery

0.0

0.2

0.4

0.6

0.8

1.0

0 1 2 3 4 5 6

Mol

e fr

acti

on o

f M

eOH

Heat duty (MJ/kg)

at topat bottom

96% MeOH recovery

79

4.5 Overall comparison of MRUs

4.5.1 Heat duty requirement

Figure 4.18 presents the heat duty for single-flash- and double-flash-based MRUs

as a function of % MeOH recovery. It is clear that there is no apparent difference in the

heat duty required for the single-flash- and double-flash-based MRUs when compared at

given final pressures of 0.1 atm (Figure 4.18-a) and 0.2 atm (Figure 4.18-b). It can, thus,

be concluded that the double-flash unit has no noticeable advantage over the single-flash

unit in the aspect of heat duty, and the single-flash should be used rather than the double-

flash unit due to the capital costs of the units.

Figure 4.19 shows a comparison of heat duty required for distillation, stripping,

and single-flash-based MRUs as a function of MeOH recovery performance. Among the

three units, the distillation-based operation requires the highest heat duty, followed by

stripping and flash, for a given target of MeOH recovery. For instance, at 0.2 atm (Figure

4.19-a) operating pressure, to recover 90% MeOH, the heat duty required for the

distillation unit is 3.71 MJ/kg, whereas the heat duty required by stripping and single-

flash units are 1.80 and 1.45 MJ/kg, respectively. At any % MeOH recovery and

operating pressures, stripping-based operation results in a significant reduction in the heat

duty compared to the distillation unit, i.e., 30 – 52% drop for 90% MeOH recovery, 30 –

45% drop for 95% MeOH recovery, and 19 – 35% drop for 98% recovery. However

single-flash-based operation results in even more reduction in heat duty compared to

distillation, i.e., 40 – 61% drop for 90% MeOH recovery, 32 – 54% drop for 95% MeOH

recovery, and 21 – 43% drop for 98% recovery. This indicates that the single-flash-based

80

(a)

(b)

Figure 4.18: Comparison between single-flash and double-flash operations at the final

pressure of (a) 0.1 atm (b) 0.2 atm. (MeOH to oil ratio = 6, feed temperature

= 60 °C and feed pressure = 4.0 atm).

81

(a)

(b)

82

(c)

Figure 4.19: Comparison of heat duty among distillation, stripping and single-flash at the

pressure of (a) 0.2 atm, (b) 0.5 atm and (c) 1.0 atm. (Reflux ratio = 2 (for

distillation), feed pressure = 4 atm, feed temperature = 60 °C and MeOH-to-

oil ratio = 6)

83

MRU is the most energy-efficient and the stripping-based MRU is the second most

energy-efficient among the three MRUs.

4.5.2 Product quality and quantity of recovered MeOH

In the aspect of recovered MeOH product, the distillation- and stripping-based

MRUs show advantages over the single- and double-flash MRUs. The distillation and the

stripping units can recover pure (approx. 100%) MeOH while at the same time achieving

high % MeOH recovery (approx. 98%) regardless of level of heat duty supplied and

operating pressure. However, the single- and double-flash units can only be operated at a

limited level of heat duty at a given operating pressure to produce pure (approx. 100%)

MeOH product. Supplying excessive heat to units will degrade the quality of MeOH

product since other chemical components will also flash off with MeOH. This, thus,

makes it difficult to operate the flash units to achieve high % MeOH recovery while

maintaining pure MeOH product.

84

5. CONCLUSIONS AND FUTURE WORK

5.1 Conclusions

Biodiesel production via transesterification requires excess methanol to move the

production reaction forward. This study focuses in detail on simulation analysis of heat

duty to recover excess methanol from the product mixture by distillation, stripping,

single-flash- and double-flash-based methanol recovery units (MRU). For each MRU,

parametric effects on heat duty and quality of recovered methanol were analyzed and

subsequently the heat duties for MRUs were compared. The following conclusions were

drawn based on the simulation results and performance analysis:

• For all MRUs, heat duty is sensitive to operating parameters and has similar

behaviour.

• Heat duty increases with the increase of % MeOH recovery. Apparently, the

change in heat duty is relatively small for a moderate recovery target and

increases significantly at high recovery targets (>95%). Heat duty also increases

with increase of reflux ratio (for distillation).

• Heat duty decreases with decrease in operating pressure and increase in MeOH-

to-oil ratio.

• The double-flash unit has no noticeable advantage over the single-flash unit in the

aspect of heat duty; hence, the single-flash should be used rather than the double-

flash due to the capital costs of the units.

• The single-flash-based MRU is the most energy efficient, followed by stripping

and distillation-based MRUs.

85

• Distillation and stripping-based MRUs can produce pure MeOH at a wide range

of operating conditions and can also achieve high % MeOH recovery.

• It is more difficult for single- and double-flash-based MRUs to achieve MeOH

recovery as high as 98% while maintaining high purity of recovered MeOH.

• Based on the heat energy requirement as well as purity of recovered methanol,

single-flash is suitable for low operating pressures while stripping-based MRUs

are suitable for higher operating pressures.

5.2 Recommendations for future work

In this work, the energy required to recover excess MeOH by distillation,

stripping, and flash-based MRUs at different operating conditions was studied.

Simulation analysis was done assuming the vegetable oil feedstock is virgin, no

esterification reaction occurs, and, hence, no water is present in the system. However, oil

may contain some free fatty acids that will undergo esterification reaction, producing

water as by product and causing soap formation. Thus, it is recommended to investigate

the performance of MRUs in the presence of water.

Results show that with increasing vacuum, the energy requirement for MeOH

recovery decreases. However, generating vacuum pressure requires energy so there is a

trade off between energy and level of vacuum. Hence, it is important to analyze the costs

associated with vacuum generation needed for MeOH recovery.

86

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