reaction kinetics of biodiesel production by using low...

REACTION KINETICS OF BIODIESEL PRODUCTION BY USING LOW

QUALITY FEEDSTOCK

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

Ling Zhou

Regina, Saskatchewan

October, 2013

Copyright 2013: L. Zhou

UNIVERSITY OF REGINA

FACULTY OF GRADUATE STUDIES AND RESEARCH

SUPERVISORY AND EXAMINING COMMITTEE

Ling Zhou, candidate for the degree of Master of Applied Science in Environmental Systems Engineering, has presented a thesis titled, Reaction Kinetics of Biodiesel Production by Using Low Quality Feedstock, in an oral examination held on August 30, 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. Daoyong Yang, Petroleum Systems Engineering

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

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

Committee Member: Dr. Stephanie Young, Environmental Systems Engineering

Committee Member: *Dr. David deMontigny, Industrial Systems Engineering

Chair of Defense: Dr. Doug Durst, Faculty of Social Work *Not present at defense

I

Abstract

Biodiesel is considered to be one of the potential renewable alternatives to

petroleum since it is biodegradable, non-toxic, and has low emission profiles. The main

challenge of its commercialization is the associated high production cost due to the high

quality feedstock used. Low quality feedstocks such as waste cooking oils are much

cheaper and more widely available. However, low quality feedstocks normally contain a

large amount of free fatty acids (FFAs), which consume the alkaline catalyst in the

biodiesel production, thereby decreasing the biodiesel production rate. An acid-catalyzed

esterification process can effectively pretreat the FFAs prior to or during the biodiesel

production. Previous studies on biodiesel production processes including esterification

and transesterification were conducted in a well-mixed system, in which the

hydrodynamic effect on the reaction could not be completely defined. Therefore, the

objective of this research is to provide a better understanding of the reaction kinetics of

acid-catalyzed esterification and alkali-catalyzed transesterification for optimizing the

biodiesel production process when using low quality feedstocks.

This study developed a new reaction system of esterification reaction in an

immiscible two-phase system, which eliminates the hydrodynamic effect on the reaction.

Based on the new reaction system, a series of experiments were conducted by using oleic

acid/linoleic acid as FFA to mix with the virgin canola oil as a low quality feedstock. The

reaction rate constant and activation energy of esterification were determined at different

temperatures. The impact of different reaction variables was evaluated in terms of FFA

conversion or acid value, including: temperature, catalyst concentration, initial FFA

content, and type of FFA. Results showed that reaction temperature, catalyst

II

concentration, and initial FFA content had great impacts on the esterification. The effect

of the catalyst concentration also depends on the reaction temperature. It had a significant

impact on esterification at high temperatures of 50°C and 62°C, but little impact at the

low temperature of 35°C. Additionally, an increase in initial FFA content increased the

reaction rate instead of the reaction rate constant. The reaction performance of oleic acid

and linoleic acid were also compared in terms of reaction rate constant and activation

energy. Oleic acid and linoleic acid were found to have the same reaction behaviour

under the same reaction conditions.

The parametric effect on the alkali-catalyzed transesterification reaction was also

evaluated in terms of FAME (fatty acid methyl esters, biodiesel) content (wt.%) of the

reaction product as a function of reaction time. The experiments were carried out in a

different experimental setup by using virgin canola oil as feedstock to react with

methanol catalyzed by sodium hydroxide (NaOH). The tested reaction parameters include

reaction temperature, catalyst concentration, and initial FFA content. The biodiesel

production rate was found to increase as the reaction temperature increased regardless of

the catalyst concentration. The achieved maximum biodiesel content ranged from 86 to

90% (w/w). An increase in catalyst concentration led to a higher biodiesel production rate,

and as expected, high contents of FFA decreased the biodiesel production rate and made

the subsequent separation process difficulty due to the undesirable soap formation. Based

on the kinetics study on transesterification, the reaction kinetics were found to be

different for low temperatures (25oC and 35

oC) and high temperatures (50°C and 65°C),

which resulted in different designs for reactor volume for a given duty based on different

temperatures.

III

Acknowledgements

First and foremost, I offer my sincerest gratitude to my co-supervisors Dr.

Amornvadee Veawab and Dr. Adisorn Aroonwilas for their enormous support and

guidance and patience throughout this thesis work. Without their valuable technical

assistance on the experimental design and troubleshooting as well as instructive

suggestions on result analysis, this thesis would not have been accomplished.

I would also like to express my sincerely appreciation to the Natural Sciences and

Engineering Research Council of Canada (NSERC) for the financial support, without

which I could not have fully concentrated on my studies.

My sincerely appreciation goes to the biodiesel research supporting organizations:

City of Regina and Communities of Tomorrow for their valuable suggestions during my

experimental work and the Faculty of Engineering and Applied Science and Faculty of

Graduate Studies and Research (FGSR) of the University of Regina for supporting an

excellent and safe laboratory and research environment.

Most important of all, I would like to express my deepest gratitude to my husband

(Zheng Cui) for his constant patience and love; my parents (Shuqing Gou and Shiqing

Zhou) and parents-in-law (Chufeng Cui and Yufang Zuo) for their endless support and

encouragement during my study; my adorable children (Jiahao Cui and Jiayue Cui) for

being so smart and lovely. Without their love, I would not have had the necessary

enthusiasm and energy to work on my research.

IV

Table of Contents

Abstract ............................................................................................................................... I

Acknowledgements ......................................................................................................... III

Table of Contents ............................................................................................................ IV

List of Tables ..................................................................................................................VII

List of Figures ............................................................................................................... VIII

Nomenclature .................................................................................................................XII

Chapter 1 Introduction and Scope of Research ..............................................................1

1.1 Introduction of Biodiesel ........................................................................................... 1

1.2 Biodiesel Production ................................................................................................. 4

1.3 Research Motivation and Objective .......................................................................... 9

Chapter 2 Literature Reivew ..........................................................................................13

2.1 Liquid/liquid Heterogeneous Reaction.................................................................... 13

2.2 Esterification Process .............................................................................................. 18

2.2.1 Chemistry of Esterification ............................................................................... 19

2.2.2 Kinetic Studies on Esterification of Biodiesel Production ............................... 20

2.3 Transesterification Process ...................................................................................... 23

2.3.1 Chemistry of Transesterification ...................................................................... 23

2.3.2 Parametric Effects on Alkali-catalyzed Transesterification Process ................ 27

2.4 Process of Biodiesel Production from Low Quality Feedstocks ............................. 29

V

Chapter 3 Acid-catalyzed Esterification Reaction ........................................................33

3.1 Acid-catalyzed Esterification Experiments ............................................................. 33

3.1.1 Materials ........................................................................................................... 33

3.1.2 Experimental Setups ......................................................................................... 33

3.1.3 Experimental Procedure and Conditions .......................................................... 37

3.1.4 Analytical Methods........................................................................................... 38

3.2 Results and Discussion ............................................................................................ 41

3.2.1 Design of a New Reaction System ................................................................... 41

3.2.2 Determination of Reaction Rate Constant and Activation Energy ................... 51

3.2.3 Parametric Effects on the Esterification Reaction ............................................ 60

Chapter 4 Alkali-catalyzed Transesterification Reaction ............................................75

4.1 Alkali-catalyzed Transesterification Experiments .................................................. 75

4.1.1 Materials ........................................................................................................... 75

4.1.2 Experimental Setups ......................................................................................... 75

4.1.3 Experimental Procedure and Conditions .......................................................... 79

4.1.4 Analytical Methods........................................................................................... 80

4.2 Results and Discussion ............................................................................................ 83

4.2.1 Effect of Reaction Temperature ....................................................................... 83

4.2.2 Effect of Catalyst Concentration ...................................................................... 88

4.2.3 Effect of FFA Content ...................................................................................... 93

VI

4.2.4 Determination of Reaction Rate Constant ........................................................ 99

4.2.5 Demonstration of Reactor Design .................................................................. 102

Chapter 5 Conclusions and Recommendations ...........................................................117

5.1 Conclusions ........................................................................................................... 117

5.2 Recommendations for Future Work ...................................................................... 120

References .......................................................................................................................122

VII

List of Tables

Table 1.1: Average biodiesel emissions compared to conventional diesels ................... 3

Table 1.2: FFAs content in various biodiesel feedstocks ............................................... 5

Table 2.1: Literatures on kinetic study of the acid-catalyzed esterification

reaction using homogenous catalysts ........................................................... 24

Table 3.1: Purities and suppliers of chemicals .............................................................. 34

Table 3.2: Experimental conditions for the acid-catalyzed esterification reaction ....... 40

Table 3.3: Reaction rate constants at 3 wt.% H2SO4..................................................... 56



Table 3.6: Activation energy in the esterification reaction of oleic acid ...................... 61

Table 3.7: Activation energy of the esterification reaction using linoleic acid ............ 72

Table 4.1: Purities and suppliers of chemicals .............................................................. 76

Table 4.2: Experiment conditions for the alkali-catalyzed transesterification

reaction ......................................................................................................... 82

Table 4.3: Duration time and conversion rate for slow reaction region (200 rpm) .... 101

Table 4.4: Observed reaction rate constant for alkali-catalyzed transesterification

(200 rpm) .................................................................................................. 104

Table 4.5: Experimental data for slow reaction region (200 rmp) .............................. 110

Table 4.6: Summary of reactor design at different temperatures (200 rpm) .............. 116

VIII

List of Figures

Figure 1.1: Simplified scheme of two biodiesel production methods from low

quality feedstocks ............................................................................................. 8

Figure 1.2: Acitavation energy obtained in different studies at different catalyst

concentrations ................................................................................................ 11

Figure 2.1: Mass transfer process based on the two-film theory ...................................... 14

Figure 2.2: Mass transfer process of FFA from the oil phase to the methanol

phase............................................................................................................... 16

Figure 3.1: Schematic diagram of experimental setup for acid-catlyzed

esterification reaction ..................................................................................... 35

Figure 3.2: Photographs of the esterification experimental setup (Original in

color) .............................................................................................................. 36

Figure 3.3: Experimental procedure for the acid-catalyzed esterification ........................ 39

Figure 3.4: Effect of position of the mechanical impeller on the FFA

conversion (T=50°C, H2SO4 concentration=3 wt.%, FFA

content=36 mgKOH/g) .................................................................................. 43

Figure 3.5: Change of the interface state with increasing agitation speeds ...................... 45

Figure 3.6: Effect of agitation speed on the FFA conversion rate (T=50°C,

H2SO4 concentration=3wt.%,, FFA content=37 mg KOH/g) ........................ 47

Figure 3.7: Effect of agitation speed on the FFA conversion rate (T=62°C,

H2SO4 concentration=3 wt.%, FFA content=37 mgKOH/g) ......................... 48

Figure 3.8: Effect of agitation speed on the FFA conversion rate (T=35°C, H2SO4

concentration=3 wt.%, FFA content=37 mgKOH/g) ..................................... 49

IX

Figure 3.9: Graph of FFA

FFA

C

C 0ln as a function of time (H2SO4 concentration=3

wt.%) (a) T=35°C (b) T=50°C (c) T=62°C ................................................... 53


FFA

C


wt.%) (a)T=35°C (b) T=50°C (c) T = 62°C ................................................... 54


FFA

C


wt.%) (a) T=35°C (b) T=50°C (c) T=62°C .................................................... 55

Figure 3.12: Arrhenius plot of lnk'RX against 1/T (Esterification of oleic acid) ................ 59

Figure 3.13: Effect of temperature on the FFA conversion (H2SO4

concentration=3 wt.%) ................................................................................... 62

Figure 3.14: Effect of temperature on the reaction rate constant (H2SO4

concentration=3 wt.%) ................................................................................. 64

Figure 3.15: Effect of catalyst concentration on the FFA conversion (Initial

FFA content=35-38 mgKOH/g) .................................................................. 65

Figure 3.16: Effect of catalyst concentration on reaction rate constant (Initial

FFA content=35-38 mgKOH/g) .................................................................. 67

Figure 3.17: Change of FFA content as a function of reaction time ................................. 68

Figure 3.18: Effect of the initial FFA content on the reaction rate constant .................... 70

Figure 3.19: Arrhenius plot of lnk΄RX against 1/T (Esterification of linoleic acid) ........... 71

Figure 3.20: Comparison of the reaction rate constants by using mixed FFA with

different ratios of oleic acid versus linoleic acid (T=62°C, H2SO4

concentration=3 wt.%) ................................................................................. 74

X

Figure 4.1: Schematic diagram of the alkali-catalyzed transesterification

experimental setup ......................................................................................... 77

Figure 4.2: Photographs of the alkali-catalyzed transesterification experimental

setup (Original in color) ................................................................................. 78

Figure 4.3: Experimental procedure for the alkali-catalyzed transesterification .............. 81

Figure 4.4: Effect of temperature on the conversion profile at 0.2 wt.% NaOH

(Methanol/Canola oil=9:1(molar ratio); mixing speed=200 rpm) ................. 84





Figure 4.7: Effect of catalyst concentration on the conversion profile at 25°C

(Methanol/Canola oil=9:1 (molar ratio); mixing speed=200 rpm) ................ 89







Figure 4.11: Effect of free fatty acid content on the biodiesel conversion at 0.2

wt. % NaOH (Methanol/Canola oil=9:1 (molar ratio); mixing

speed=200 rpm; T=65°C) ............................................................................... 95

XI


wt.% NaOH (Methanol/Canola oil=9:1 (molar ratio); mixing

speed=200 rpm; T=65°C) ............................................................................ 96


wt.% NaOH (Methanol/Canola oil=9:1 (molar ratio); mixing

speed=200 rpm; T=65°C ) ........................................................................... 97

Figure 4.14: Photographs showing appearances of separation of reaction

mixtures in the separating funnel (Sample collected at reaction

time=1 hour; NaOH (wt.%) =0.6%; methanol/Canola oil=9:1

(molar ratio); mixing speed=200 rpm; T=65°C ) ....................................... 98

Figure 4.15: Plots of )1

1(ln

xvs. reaction time, t, (a) NaOH concentration =0.2

wt.% (b) NaOH concentration =0.6 wt.% (c) NaOH concentration

=1 wt.% .................................................................................................... 103

Figure 4.16: Batch reaction process ................................................................................ 105

Figure 4.17: Plug flow reaction process ......................................................................... 106

Figure 4.18: Schematic of a batch reactor ...................................................................... 108

Figure 4.19: Schematic of a plug flow reactor ................................................................ 112

Figure 4.20: Schematic of a continuous stirred tank reactor .......................................... 114

XII

Nomenclature

A frequency factor

ASTM D6751 the US Standard Specification for Biodiesel

CAL

concentration of alcohol solution in alcohol bulk

C*

FFA-i concentration of FFA at the interface in the oil film

CFFA-O concentration of FFA in the oil bulk

CH3OK

potassium methoxide

CH3ONa

sodium methoxide

CO

carbon monoxide

CO2

carbon dioxide

CSTR continuous stirred-tank reactor

Ea

activation energy(kJ mol−1

)

FAME fatty acid methyl esters

Fe2(SO4)3 ferric sulphate

FFA free fatty acid

H2O water

H2SO4 sulphuric acid

HCl hydrochloric acid

HI hydriodic acid

k reaction rate constant (min−1

)

k' pseudo-rate constant (min

−1)

kFFA-O mass transfer coefficient of FFA in the oil film (mol m

−3 min

−1)

kFFA-AL

mass transfer coefficient of FFA in the methanol film (mol m−3

XIII

min−1

)

KOH potassium hydroxide

Na2SO4 sodium sulfate

NaOH

sodium hydroxide

NOx

nitrous oxide

nPAH

nitrated polycyclic aromatic hydrocarbons

P*

FFA-i

concentration of FFA at the interface in the methanol film

PFFA-AL

concentration of FFA in the methanol bulk

PAH polycyclic aromatic hydrocarbons

PM particulate matter

r

reaction rate (mol m−3

min−1

)

rFFA mass transfer rate of FFA to the interface (mol m−3

min−1

)

FFFA

flux rate of alcohol (mol min−1

)

R universal gas constant ( 8.314 J mol−1

K−1

)

COOH-R1 FFA

OHR2 alcohol

21 COORR biodiesel

S interfacial surface (m2)

SnCl2 tin chloride

SOx sulfur oxides

t time (s)

T temperature ( K)

VRX reaction volume (m3)

Greek letters

XIV

α the reaction order with respect to FFA

β reaction order with respect to alcohol

1

Chapter 1 Introduction and Scope of Research

1.1 Introduction of Biodiesel

Energy use is considered to be the most fundamental requirement for human

existence (Refaat, Attia et al. 2008). Among different kinds of fuels, petroleum

constitutes the majority of the world’s energy supply. It plays a significant role in

industry, transportation, and agriculture, as well as to meet many other basic human

needs (Shahid and Jamal 2008). According to Johnston and Holloway (2007), the global

demand for petroleum is predicted to increase 40% by 2025. However, petroleum is a

finite and nonrenewable energy source, which has already caused serious environmental

pollution. Therefore, a sustainable, affordable, and environmentally friendly alternative to

petroleum is urgently needed.

Biodiesel is considered to be one of the most attractive alternatives to

conventional petroleum-based diesel (Santacesaria, Tesser et al. 2007). It is composed of

mono-alkyl esters of long chain fatty acids derived from a renewable lipid feedstock,

which conforms to the US Standard Specification for Biodiesel (ASTM D6751). “Bio”

indicates a source of energy that is biological and renewable; “diesel” means it can be

used only in diesel engines (Zhang, Dube et al. 2003). Biodiesel can be used directly in a

diesel engine in its neat pure form called B100 (100% biodiesel) or in a blend with

different proportions of conventional diesel fuels. Common blends include B20 (20%

biodiesel and 80% conventional diesel), which are much closer to diesel fuel properties

than B100 and B5 (5% biodiesel and 95% conventional diesel).

2

Compared to conventional diesels, biodiesel has a number of advantages as

follows:

(1) It can be directly used in a diesel engine without any modification;

(2) It is renewable, non-toxic, and biodegradable since the feedstock is originally

from plants or animals such as soybean, canola, palm, corn, and animal fat;

(3) Combustion of biodiesel does not increase current net atmospheric level of

carbon dioxide (CO2), one of the greenhouse gases, because the carbon in

biodiesel is originally removed from the air by plants; and,

(4) Use of biodiesel can reduce air pollution because biodiesel emissions have

lower levels of particulate matter (PM), carbon monoxide (CO), sulfur oxides

(SOx), hydrocarbons, soot, and other byproducts, as shown in Table 1.1.

and biodiesel has several disadvantages as well:

(1) The high cost of biodiesel, which is about one and a half times that of

conventional petroleum-based diesel, is the main challenge to its

commercialization (Zhang, Dube et al. 2003). The cost of biodiesel depends

on a number of factors such as the cost of feedstocks and reactants, the nature

of its purification, its storage, and so on. Of all of these, the cost of feedstock

accounts for 70-95% of the total biodiesel production cost (Krawczyk 1996;

JConnemann 1998). Therefore, using more economic feedstocks such as

waste cooking oils and fats can significantly decrease the biodiesel cost;

(2) Biodiesel may become a gel in cold weather since its cloud point is generally

higher than conventional diesels. Addition of cold flow additives can prevent

it from gelling at a low temperature;

3

Table 1.1: Average biodiesel emissions compared to conventional diesels

Emission type B20 B100

Total unburned hydrocarbons -20% -67%

Carbon monoxide (CO) -12% -48%

Carbon dioxide (CO2)—life cycle production -16% -79%

Particulate matter (PM) -12% -47%

Nitrogen oxides (NOx) +2% +10%

Sulfur oxides (SOx) -20% -100%

Polycyclic aromatic hydrocarbons (PAH) -13% -80%

Nitrated PAH (nPAH) -50% -90%

Source: (Sheehan, Camobreco et al. 1998)

4

(3) Biodiesel solvent can cause degradation of rubbers and elastomers. Using

low-percentage biodiesel blends can mitigate the degradation, but the

compromise will somewhat lessen the positive environmental benefits; and,

(4) Biodiesel emissions contain more smog-forming nitrous oxide (NOx) than

conventional diesels, which may delay the injection timing of engines.

1.2 Biodiesel Production

There are four well-established biodiesel production methods: direct use and

blending, micro-emulsions, thermal cracking (pyrolysis), and transesterification (Ma and

Hanna 1999). Among these methods, transesterification is one of the most commonly

used methods in the biodiesel production industry, which uses vegetable oils or animal

fats as feedstock to react with a short chain alcohol (methanol, ethanol, butanol, or amyl

alcohol) to yield biodiesel as the main product and glycerol as a by-product. The

transesterification reaction is shown in Equation 1.1

CH2-OOC-R1

CH2-OOC-R3

CH1-OOC-R2+ 3R'OH

CatalystR1-COO-R'

R2-COO-R'

R3-COO-R'

+

CH2-OH

CH1-OH

CH2-OH

Glyeride Alcohol Esters Glycerol

[1.1]

where R1, R2, R3, R' = alkyl groups.

Feedstock quality is a significant factor affecting biodiesel production. A wide

variety of materials including fats, oils, or other grease sources can be used for biodiesel

production. According to the content of FFAs in the feedstock, feedstocks can be

categorized as high quality feedstocks and low quality feedstocks. As shown in Table 1.2,

5

Table 1.2: FFAs content in various biodiesel feedstocks

Feedstock FFAs

Refined vegetable oils < 0.05 %

Crude vegetable oil 0.3 – 0.7%

Restaurant waste grease 2 – 7%

Animal fat 5 – 30%

Trap grease 40 – 100%

Source: (Van Gerpen, Shanks et al. 2004)

6

low quality feedstocks usually have higher content of FFAs than high quality feedstocks.

For example the content of FFAs in high quality feedstocks such as refined vegetable oils

and crude vegetable oils is less than 1%. However in the low quality feedstocks such as

restaurant waste grease and animal fats, the content of FFAs is higher than 2% and can

even reach 100% in the trap grease.

Industrial biodiesel production from a high quality feedstock normally includes a

transesterification process and a purification process. The homogenous alkaline catalyst

is commercially used to catalyze the transesterification process (Abbaszaadeh, Ghobadian

et al. 2012). Homogeneous alkaline catalyst, such as alkaline metal alkoxides and

hydroxides as well as sodium or potassium carbonates, are used. It has high catalytic

activity and is widely available and economical, while it also requires modest operational

conditions and achieves high conversion in a minimal time. After the transesterification

process, the production residues and impurities left in the crude biodiesel must be

removed by a separation process since they may damage the engine combustion systems.

The production impurities through the transesterification process include glycerol,

unreacted alcohol, catalyst, and other side reaction byproducts such as water, soap, etc.

Since the cost of a high quality feedstock typically accounts for 70-95% of total

biodiesel production cost. Low quality feedstocks are much cheaper than high quality

feedstocks. Thus, using low quality feedstocks such as waste cooking oils or non-edible

oils instead of high quality feedstocks will significantly reduce the biodiesel production

cost. However, low quality feedstocks have high content of FFAs, which can react with

the alkaline catalyst and produce soaps. This side reaction in the alkali-catalyzed

transesterification process will reduce the catalyst efficiency and the biodiesel conversion

7

rate. Additionally, the formation of soaps will make the later purification process difficult.

As a result, the undesired side reactions caused by FFAs will increase the cost of

biodiesel production. Therefore, when using low quality feedstocks for biodiesel

production, the content of FFAs must be reduced to an acceptable level (typically below

1% according to Freedman, Pryde et al. (1984); Liu (1994); and Ma, Clements et al.

(1998)) before the alkali-catalyzed transesterification process. One efficient method for

removing the FFAs from feedstocks is esterification. As shown in Equation 1.2, in the

esterification reaction the FFAs react with a low molecular weight alcohol, such as

methanol, ethanol, isopropanol or butyl, to produce biodiesel.

biodiesel alcohol FFAs

OHCOORR OHR COOH-R 22121 [1.2]

where R1= a linear chain of 11–17 carbon atoms containing a variable number of

unsaturations depending on the particular origin of the raw material; R2 = a methyl radical.

However, the esterification reaction is extremely slow, taking several days to reach

equilibrium at typical reaction conditions (Liu, Lotero et al. 2006). A variety of catalysts

can effectively increase the reaction rate, including: homogenous mineral acids (sulfuric

acid (H2SO4), hydrochloric acid (HCl) or hydriodic acid (HI)), and heterogeneous solid

acids (various sulfonic resins).

As shown in Figure 1.1, there are normally two methods for producing biodiesel

from low quality feedstocks. Method Ι has two consecutive reaction steps. The first one is

a pretreatment step, in which an acid-catalyzed esterification reaction occurs to reduce

the FFAs content to an acceptable level. Then the product of the esterification reaction is

8

Figure 1.1: Simplified scheme of two biodiesel production methods from low quality

feedstocks. Source: (Van Gerpen, Shanks et al. 2004)

Method І:

Method П:

Transesterification Purification

Acid Catalyst Alkaline Catalyst

Esterification

Esterification

+

Transesterification

Purification

Acid Catalyst

9

sent to the second reaction system. In the second step, an alkali-catalyzed

transesterification occurs to convert the oils/fats to crude biodiesel. Method Ι can produce

biodiesel from a low quality feedstock quickly and effectively. However, the production

cost is increased due to an additional pretreatment step. Method Π has one acid-catalyzed

reaction step, in which the esterification reaction and transesterification reaction take

place in one unit at the same time. Since the acid-catalyst is insensitive to the FFAs, an

acid catalyst is added to the system for accelerating both esterification and

transesterification reactions. Method Π is more efficient and economic than method I

since it has only one reaction step (Zhang, Dube et al. 2003). However the reaction rate

of Method Π is very low compared to that in Method Ι, normally taking several days to

complete.

1.3 Research Motivation and Objective

The esterification and transesterification reactions are essentially heterogeneous

because the nonpolar oil phase and the polar alcohol phase are immiscible with each

other. Therefore, their overall reaction rates mainly depend on two important factors: the

hydrodynamic effect between these two phases and the chemical reaction kinetics. In

order to optimize the biodiesel production process and design a high performance

reaction system, the hydrodynamic effect and chemical reaction kinetics must be

completely understood.

Previous kinetic studies on the esterification reaction were mostly carried out in a

pseudo-homogenous reaction system. Sufficient mixing was provided in these systems in

order to eliminate the hydrodynamic effect on the overall reaction rate. The previous

10

results on activation energy are plotted as a function of catalyst concentration in Figure

1.2. The result shows a large deviation and an inconclusive trend as the catalyst

concentration increases. According to Boocock, Konar et al. (1996), the hydrodynamic

effect was significant when a heterogeneous reaction system was vigorously agitated.

Von Blottnitz, Sadat-Rezai et al. (2004) found that even in a homogeneous reaction

system, the hydrodynamic effect still existed when a co-solvent was used. Therefore, the

inconsistent results from previous studies indicate that the hydrodynamic effect may exist

in their systems and affect the overall reaction rate.

In an agitated system, improved mixing can help enhance the mass transfer

coefficient and also increase the interfacial surface area available for the reaction.

According the results of Fernandes and Sharma (1967), the mass transfer coefficient and

interfacial area increase with the increasing mixing speed until an equilibrium stage is

reached when there is no significant increase in both. When an equilibrium stage is

reached, even if the agitation speed increases, the mass transfer rate keeps constant.

Therefore, in the previous studies, the change of the mass transfer rate does not affect the

reaction rate at an equilibrium stage. However, the hydrodynamic effect may still exist in

the reaction system. Furthermore, except for the speed of the agitator, there are other

variables affecting the hydrodynamic effect including the ratio of the agitator diameter to

the vessel diameter, position of the agitator in the reactor, the liquid level, dispersed

phase hold-up, etc. Thus, the hydrodynamic effect on the reaction still varies at the

equilibrium stage in this way.

11

Figure 1.2: Acitavation energy obtained in different studies at different catalyst

concentrations

0

10

20

30

40

50

60

70

0 2 4 6 8 10 12

Acti

va

tio

n E

nerg

y k

J/m

ol

H2SO4 wt.%

Sendzikiene, Makareviciene et al. (2004)

Berrios, Siles et al. (2007)

Aranda, Santos et al. (2008)

Supardan (2008)

Thiruvengadaravi, Nandagopal et al. (2009)

Praveen K.S Yadav.et al (2010)

12

The objectives of this study are: i) to develop a new reaction system of the

esterification reaction that takes into account the effects of both hydrodynamic and

chemical kinetics,ii) to evaluate the parametric effect on the FFA conversion rate of the

esterification reaction and obtain the reaction rate constant and activation energy, and iii)

to evaluate the reaction kinetics of alkali-catalyzed transesterification reaction, and

discuss the parametric effects on the reaction conversion rate. Virgin canola oil was used

as a high quality feedstock. Mixtures of virgin canola oil and pure oleic acid or pure

linoleic acid were used as substrates for low quality feedstocks in this study.

This thesis consists of five chapters. Chapter 1 introduces the general background

of biodiesel production technologies and the research motivation and objectives. Chapter

2 provides a comprehensive literature review on biodiesel production. Chapter 3

describes the details of experiments for esterification reaction, including materials, setups,

conditions, procedures, and sample analysis, and also provides the experimental results

and discussion. Chapter 4 describes the details of the experiments for transesterification,

and also discusses the parameters effects on the reaction rate and different reactor designs

for a given duty under certain conditions. Finally, Chapter 5 summarizes the research

results and provides recommendations for future work.

13

Chapter 2 Literature Review

2.1 Liquid/liquid Heterogeneous Reaction

Since the polar alcohol and nonpolar oil are two immiscible phases, one must

diffuse into the other before the heterogeneous reaction between them can happen. Thus,

both a mass transfer process of the reactant(s) from one phase to the other phase and a

chemical reaction take place in the heterogeneous reaction. The overall reaction rate

expression should consist of the mass transfer rate and the chemical reaction rate.

Mass transfer is a process in which one component diffuses from one phase to

another phase or the same phase because of a concentration difference (Strigle 1987). It

occurs in various industrial operations such as distillation, absorption, evaporation,

adsorption, and liquid/liquid extraction. The theories that can be used to describe the

mass transfer process include the two-film theory, surface renewal theory, and boundary

layer theory (Geankoplis 1993). The two-film theory was adopted in most studies

because it is the simplest theory and leads to a very similar result to the others.

For convenience in notation, the two-film theory is discussed by using a

gas/liquid reaction as an example. A liquid/liquid heterogeneous reaction has a similar

mass transfer process as a gas/liquid reaction. According to the two-film theory, as

illustrated in Figure 2.1, the gas/liquid phases are separated by an interface. There is one

film in either phase that adheres to the interface. For a fast gas/liquid reaction, mass

transfer occurs through the following three consecutive steps:

(1) The component A in the gas bulk diffuses through the gas film. There is a gas

phase mass transfer resistance in the gas film.

14

Figure 2.1: Mass transfer process based on the two-film theory (Redrawn from Astaria,

Savage et al. (1983))

Distance from interface

Reaction zone

Interface Direction of mass transfer

Gas film

Co

nce

ntr

ati

on

of

solu

te A

Liquid film

yA,G

yA,i

CA,i

CA,L

15

(2) The component A diffuses through the gas/liquid interface. It is assumed that

no mass transfer resistance exists at the interface.

(3) The component A diffuses through the liquid film. If the chemical reaction is

fast, it takes place in the liquid film. There is a liquid phase mass transfer

resistance in the liquid film.

The esterification reaction of biodiesel production is a heterogeneous reaction

between the alcohol and FFAs, which is presented as follows:

[2.1]

Since the majority of the alcohol is expected to be in the polar alcohol phase (Ataya,

Dubé et al. 2007) and the catalyst is located only in the methanol phase, the esterification

reaction mostly completes in the methanol phase. Thus, FFA must enter the methanol

phase first in order to react with methanol. Figure 2.2 illustrates the complete mass

transfer process of FFA from the oil phase to the methanol phase. Firstly, FFA diffuses

through the oil film from the oil bulk phase, then through the oil/methanol interface, and

finally through the methanol film to the methanol bulk. The mass transfer rates of FFA

are given by the rate expressions of Equations 2.2 and 2.3:

In the oil film:

)( *iFFAOFFAOFFAFFA CCakr

[2.2]

where FFAr = the mass transfer rate of FFA to the interface; OFFAk = the mass transfer

coefficient of FFA in the oil film; OFFAC = the concentration of FFA in the oil bulk;

*iFFAC = the concentration of FFA at the interface in the oil film; a= the interfacial area

in per unit volume.

Alcohol + FFA

Catalyst

Ester + H2O

16

Figure 2.2: Mass transfer process of FFA from the oil phase to the methanol phase

PFFA-AL

Oil Film Methanol Film

Interface

CFFA-O

C*

FFA-i

P*

FFA-i Oil bulk Methanol

bulk

b Direction of mass transfer

b

Co

nce

ntr

ati

on

of

FF

A

17

In the methanol film:

)( *ALFFAiFFAALFFAFFA PPakr

[2.3]

where ALFFAk = the mass transfer coefficient of FFA in the methanol film; *iFFAP = the

concentration of FFA at the interface in the methanol film; ALFFAP = the concentration of

FFA in the methanol bulk phase.

According to Ataya, Dubé et al. (2007), the esterification is almost an instaneouse

reaction, so it can be assumed that the reaction takes place only at the interface. Then, the

overall mass transfer resistance of FFA only exists in the oil film. From Equation 2.2, the

flow rate of FFA is obtained as follows:

RXiFFAOFFAOFFARXFFAFFA VCCakVrF )( *

[2.4]

and the equation for the overall reaction rate at the interface is:

iALiFFARXRX CCkr *

[2.5]

where VRX = the reaction volume; kRX = the reaction rate constant; α= the reaction order

with respect to FFA; β = the reaction order with respect to alcohol. iALC = the

concentration of methanol at the interface.

Since the methanol used in this study is pure and in a large quantity, the

concentration of methanol remains constant in the methanol phase during the

esterification reaction. iALC becomes a constant and can combine with kRX. Therefore,

the overall reaction rate can be expressed with respect to the concentration of FFA:

*' iFFARXRX Ckr

[2.6]

where k'RX = the pseudo-rate constant.

18

The flow rate of FFA at the interface is:

RXiFFARXRXFFAFFA VCkVrF*'

[2.7]

According to the previous studies from Sendzikiene, Makareviciene et al. (2004);

Kocsisová, Cvengroš et al. (2005); Cardoso, Neves et al. (2008); Aranda, Santos et al.

(2008); Thiruvengadaravi, Nandagopal et al. (2009), the esterification reaction follows a

first-order kinetic law with respect to the concentration of FFA. By combining Equation

2.4 and 2.7, the following equation is obtained:

7.2*

4.2* )()( EqiFFAEqiFFAOFFAOFFA CCCC =

)'

11(

' RXOFFARX

FFA

RXRX

FFA

RXOFFA

FFA

kakV

F

Vk

F

aVk

F

[2.8]

The overall reaction rate of the esterification reaction is obtained by rearranging Equation

2.8:

OFFA

RXOFFA

RX

FFA C

kak

rV

F

'

11

1 [2.9]

2.2 Esterification Process

Biodiesel is normally made from high quality feedstocks, such as edible oils.

However, there is a large amount of low quality feedstocks that can be converted to

biodiesel. The challenge of using low quality feedstocks for biodiesel production is that

the low quality feedstock contains a large amount of FFAs, which can have a side

reaction with the alkali-catalyst used in the transesterification process to produce

undesirable soaps, inhibiting the separation of biodiesel from glycerol. Soap formation

19

can also produce water that will hydrolyze the triglycerides and aggravate the soap

formation. This undesirable side reaction will add a fixed cost due to the use of an

additional unit for removing soaps and also lead to a reduction of the yield.

When using a low quality feedstock for biodiesel production, a pretreatment step,

i.e., esterification, is required. In the esterification process, the FFAs are converted into

biodiesel without forming soaps, which increases the final yield. It can take place without

any catalyst due to the weak acidity of carboxylic acids, but the reaction is extremely

slow and requires several days to complete at typical reaction conditions. Previous

research results showed that either homogenous mineral acids, such as H2SO4, HCl, or HI,

or heterogeneous solid acids, such as various sulfonic resins, can effectively catalyze the

esterification reaction. The homogenous catalyst is more effective than the heterogeneous

catalyst in the esterification reaction, and the reaction kinetics using heterogeneous

catalysts are more complicated than those using homogenous catalysts since the

restriction of both absorption and dis-absorption rates in the pore of the catalyst needs to

be considered in the overall reaction rate.

2.2.1 Chemistry of Esterification

The mechanism of esterification reaction involves a process related to

nucleophilic substitution. It can be illustrated in the following scheme:

Step 1: the carboxylic acid is protonated initially by the strong inorganic acid

catalyst (typically H2SO4):

[2.10]

20

Step 2: the alcohol nucleophile (two lone pairs on the oxygen) adds the sp2 carbon

and the alcohol proton is lost:

[2.11]

Step 3: the new ester bond between the carboxyl group carbon and the alcohol

oxygen is formed:

[2.12]

Step 4: H2O is eliminated at one site or the other:

[2.13]

Step 5: the excess proton leaves, regenerating the inorganic acid catalyst:

[2.14]

2.2.2 Kinetic Studies on Esterification of Biodiesel Production

There are very few studies reported on the kinetic study of the esterification

reaction of biodiesel production. Most of them were limited to their particular reaction

conditions.

21

Berrios, Siles et al. (2007) carried out a kinetic study on the esterification of

sunflower oils with an anhydrous methanol. Their kinetic model was developed based on

several assumptions: i): the reaction under the operating conditions was controlled by a

chemical reaction; ii): the non-catalyzed reaction was negligible; iii) the esterification

reaction occurs in the oil phase; iv): the methanol concentration was constant throughout

the reaction; and v): the reaction system was pseudo-homogeneous, first-order in the

forward direction, and second-order in the reverse direction. The reaction rate is

determined by the forward reaction rate and reverse reaction rate as shown in Equation

2.15:

[2.15]

where [A]= the concentration of FFA in mgKOH/g oil; [C]= the concentration of FAME,

which is assumed to be zero (t=0); [D]= the concentration of water, which is assumed to

be zero (t=0); K1= the reaction constant of the forward reaction; K2= the reaction constant

of the reverse reaction. [A0]= the initial concentration of FFA; [E]= the removed acidity.

Since the concentration of FFA in the system is determined by its initial concentration,

and the removed acidity, Equation 2.15 is rearranged as shown in the following:

[2.16]

By integrating Equation 2.16, the kinetic model is obtained as follows:

[2.17]

where

22

[2.18]

[2.19]

[2.20]

The activation energy was calculated by using the Arrhenius equation. The same kinetic

model was adopted by Supardan (2008) and Thiruvengadaravi, Nandagopal et al. (2009).

Sendzikiene, Makareviciene et al. (2004) used a mixture of rapeseed oil and oleic

acid as a low qualitity feedstock. It reacted with anhydrous methanol, and sulfuric acid

was added to the system to catalyze the reaction. The mixing speed was selected at a

constant of 850 rpm. During the reaction, it was found that diffusion restrictions are

characteristic for the entire ranges of FFA concentrations and reaction times, since the

reaction rate constant changed during the reaction time.

Aranda, Santos et al. (2008) studied the esterification of palm fatty acids with an

anhydrous methanol and ethanol by using homogeneous catalysts. The reaction happened

in a 600 mL stainless steel batch reactor. The agitation speed was kept constant (500 rpm).

The reaction rate constants and reaction orders were estimated using the following model:

[2.21]

where FA= fatty acid; ALC= alcohol.

23

Yadav, Singh et al. (2010) studied the reaction kinetics of the esterification of the

palm fatty acid. Though the hydrodynamic effect on the overall reaction rate was not

considered in their model.

Kinetic study of esterification of oleic acid in soybean oil using ethanol was

evaluated by Cardoso, Neves et al. (2008). The kintic model was expressed as Equation

2.22. The resulting data fits a first order kinetic behaviour. However, the hydrodynamic

effect on the overall reaction rate was still not considered in the model.

[2.22]

The important kinetics findings and results from previous studies are summarized

in Table 2.1.

2.3 Transesterification Process

Transesterification, also called alcoholysis, is a traditional technology to produce

biodiesel. It is the most effective process to transform the big triglyceride molecules into

small and straight-chain molecules of fatty acid esters. It can reduce the molecular weight

to one-third that of the oil and the viscosity by a factor of eight, and it can increase the

volatility.

2.3.1 Chemistry of Transesterification

In the biodiesel transesterification process, triglycerides react with an alcohol in

the presence of some catalyst to produce esters (biodiesel) and another alcohol (glycerol).

As shown in Equation 2.23, Equation 2.24, and Equation 2.25, the transesterification

reaction is reversible and includes three consecutive steps: conversion of triglycerides to

24

Table 2.1: Literatures on kinetic study of the acid-catalyzed esterification reaction using homogenous catalysts

References Research Objectives Test Conditions Important Findings

Yadav, Singh et al.

(2010)

Optimized the conditions for

production of palm fatty acid methyl

esters

Reactants: Palm fatty acid(FFA=93 wt.%)

Alcohol: methanol

Catalyst: H2SO4

500 mL three neck flask, stirrer

Pseudo first-order kinetics for

esterification.

Et=15.31 kJ mol-1

Thiruvengadaravi,

Nandagopal et al.

(2009)

Optimized the pretreatment process.

Undertook kinetic and

thermodynamic studies of

esterification

FFA: FFA in Pongamia

Alcohol: methanol

Catalyst: H2SO4

Bath reactor, mechanical stirrer, speed

(N/A)

Pseudo first-order kinetics for

esterification.

Rate constants and activation energy

were determined.

Optimum conditions: methanol to oil

ratio=9:1, H2SO4=1 wt.%,

temperature=60oC

Ea=280.1J/mol at H2SO4=1 wt.%,

Cardoso, Neves et al.

(2008)

Evaluated the use of SnCl2·2H2O as

catalyst for the ethanolysis of oleic

acid (pure and added to soybean oil)

Investigated key parameters of

reaction

FFA: oleic acid in soybean oil

Alcohol: ethanol

Catalyst: H2SO4, tin chloride (SnCl2)

50 mL three-necked glass flask, magnetic

stirrer speed N/A

SnCl2 is a potential catalyst for the low

quality raw materials.

A first order dependence for both

esterification reaction catalyzed by

H2SO4 and SnCl2.

Aranda, Santos et al.

(2008)

Studied the esterification of palm

fatty acids, by-products of edible

palm Oil production, to produce

biodiesel, using homogeneous acid

catalysts.

FFA: palmitic and oleic acids in palm oil

Alcohol: methanol, ethanol

Catalyst: H2SO4 (98%); phosphoric acid

(85%); trichloroacetic acid (98%) and

methanesulfonic acid (95%).

Stainless steel 600 mL batch reactor (PARR

842), stirring peed=500 rpm

First order with respect to fatty acid and

zero order with respect to alcohol.

Ea=15.046 Kcal/mol at H2SO4=0.01

wt.%, Ea=10.054 Kcal/mol at

H2SO4=0.0 3 wt.%, Ea=6.528 Kcal/mol

at H2SO4=0.05 wt.%.

Yalçinyuva, Deligöz

et al. (2008)

Studied the esterification kinetics of

myristic acid with isopropyl alcohol

with both homogeneously and

heterogeneously catalyzed systems

FFA: myristic acid

Alcohol: isopropyl alcohol

Catalyst: ρ-toluene sulfonic acid, amberlyst-

15 and Degussa (acidic cation exchange

resin)

250 mL round bottomed reactor, magnetic

stirrer, mixing speed=450 rpm

Second-order kinetics for the

homogeneous catalyst.

No pore diffusion when using

heterogeneously catalyst limitation.

Supardan (2008)

Studied the effect of operational

variables on the esterification of

FFA in low grade CPO;

Studied the influence of operational

variables on the kinetics.

FFA: low grade CPO with FFA content of

5.6% and 33.3%.

Alcohol: methanol

Catalyst: H2SO4

Mechanical agitation =464 rpm

The esterification reaction of FFA in

low grade CPO is irreversible.

A first-order kinetic law for the

reaction. Ea=30.4 kJ/mol, A=305

25

Table 2.1: Literatures on kinetic study of the acid-catalyzed esterification reaction using homogenous catalysts (cont’d)

Berrios, Siles et al.

(2007)

Examined the influence of

operational variables on the

kinetics

FFA: Fatty acids in sunflower oil

Alcohol: methanol

Catalyst: H2SO4

Bath reactor, magnetic agitation speed 600

rpm


forward reaction and a second-order

for the reverse reaction.

Ea=50.745 kJ/mol at H2SO4=5 wt.%,

Ea=44.559kJ/mol at H2SO4=10 wt.%

Kocsisová, Cvengroš et

al. (2005)

Studied the reaction of ester

preparation in short reaction time

with small excess of alcohol, low

catalyst, but higher conversion to

esters

FFA: Commercial mixture of fatty acids,

mixture of ME and FFA with different acid

values

Alcohol: methanol

Catalyst: ρ -Toluene-sulfonic acid

Ambient pressure, temperature above the

boiling point of MeOH, continual flow of

liquid MeOH into the reaction mixture.

mechenical stirrer(speed N/A)


reactions.

The reaction rate is two to three

times higher than at the

temperatures close to the boiling

point of MeOH

Sendzikiene,

Makareviciene et al.

(2004)

Determined the optimal

conditions of free fatty acid

esterification by methanol using

acid catalyst;

Calculated the kinetic parameters

of this process.

FFA: Oliec acid

Alcohol: Anhydrous methanol

Catalyst: Concentrated H2SO4

500 mL 3-neck distillation flask mechanical

stirrer(speed 800 min-1)

First order of the reactions after

excluding agent diffusion on the

reaction rate.

Et ≈ 13.3 kJ/mol under the

experimental conditions

Diffusion restrictions are

characteristic for the entire range of

concentrations and reaction times

studied

26

diglycerides; conversion of diglycerides to monoglycerides; and conversion of

monoglycerides to glycerol:

Triglyceride (TG) +R′OH Diglyceride (DG) + R′COOR1 [2. 23]

Diglyceride (DG) +R′OH Monoglyceride (MG) + R′COOR2 [2. 24]

Monoglyceride (MG) +R′OH Glycerol (GL) + 3R′COOR3 [2. 25]

where R1, R2, R3, R' = alkyl groups.

The overall reaction of transesterification is expressed as follows:

Triglyceride (TG) +3 R′OH Glycerol (GL) + 3R′COOR3 [2.26]

The alcohols used in the transesterification process can be methanol, ethanol,

propanol, butanol, or amyl alcohol. Methanol and ethanol are used most frequently.

However, methanol is usually preferred since it is relatively inexpensive and has small

molecular mass. In addition, a lower amount of methanol is needed than ethanol and it

can react with triglycerides quickly.

Since alcohol and triglycerides are immiscible, a catalyst is needed to accelerate

the transesterification reaction rate and the specific yield. Several different types of alkali

and acid catalysts are normally used, such as NaOH, potassium hydroxide (KOH), H2SO4,

ion exchange resins, lipases, and supercritical fluids. The most commonly used catalysts

are strong alkaline catalysts. Acid catalysts are normally used for the esterification of

FFA when using a low quality feedstock.

27

2.3.2 Parametric Effects on Alkali-catalyzed Transesterification Process

2.3.2.1 Effect of the FFA and moisture contents

For the alkali-catalyzed transesterification, the feedstock is very sensitive to the

FFA content, and all materials should be substantially anhydrous (Wright, Segur et al.

1944). The presence of FFA and water can cause an undesired side reaction with the

catalyst and produce soaps. Therefore， the effectiveness of catalyst is reduced and the

formed soaps increase the viscosity of the reaction mixture. High viscosity will lead to

the formation of gels, which make the latter separation of glycerol difficult. Meher, Vidya

Sagar et al. (2006) indicated that the FFA and moisture contents are key parameters for

determining the feasibility of the transesterification process. It is suggested that the FFA

content of the feedstock used in the transesterification process should be as low as

possible, typically below 1% (acid value less than 2 mgKOH/g). Canakci and Van

Gerpen (2001) reduced the recommended acidity to below 0.5%.

2.3.2.2 Effect of the molar ratio of alcohol to triglyceride

The molar ratio of alcohol to triglyceride is another important factor affecting the

yield of biodiesel. According to Equation 2.26, the transesterification reaction requires

three moles of alcohol and one mole of triglyceride to yield three moles of fatty acid alkyl

esters and one mole of glycerol. However, transesterification is an equilibrium reaction so

a large amount of excess alcohol is required to advance the reaction. Freedman, Pryde et

al. (1984) studied the effect of different molar ratios of alcohol to triglyceride from 1:1 to

6:1 on the transesterification reaction by using different vegetable oils including soybean,

sunflower, peanut, and cotton seed oils. For all the tested oils, the highest conversions (93%

28

-98%) were achieved at a 6:1 molar ratio of alcohol to oil. Rashid and Anwar (2008)

found the optimum yield (98%) of biodiesel was obtained at a 6:1 molar ratio of alcohol

to oil. In the case that molar ratios of alcohol to oil are higher than 6:1, separation of

esters from glycerol will be difficult. The excess methanol can hinder the gravity

decantation, and a portion of the glycerol will remain in the biodiesel phase.

2.3.2.3 Effect of the catalyst type and concentration

Catalysts used for the transesterification are classified as alkali, acid, or enzyme,

among which alkali catalysts are more effective (Freedman, Pryde et al. 1984). Alkaline

metal hydroxides, such as KOH and NaOH, and metal alkoxides, such as sodium

methoxide (CH3ONa), can be used as catalyst to accelerate the transesterification reaction.

Alkaline metal hydroxides are cheaper and less active than metal alkoxides, but they can

achieve the same conversions of vegetable oils just by increasing their concentrations to 1

or 2 mol% (Schuchardt, Sercheli et al. 1998). Currently, they are being widely used in

industrial biodiesel production. KOH was used by Vicente, Martínez et al. (2006) at 25°C

and 45°C and they found the reaction rate increased as the KOH concentration increased.

The same behaviour was also observed in other temperatures. Leung and Guo (2006)

found that the maximum content of biodiesel was reached when the catalyst

concentrations of NaOH, CH3ONa, or KOH are 1.1, 1.3, or 1.5 wt.%, respectively.

Moreover, the biodiesel yields, by using NaOH and KOH as catalyst, were lower than

that of CH3ONa. Meka, Tripathi et al. (2007) studied the effect of NaOH concentration

on the reaction time at two temperatures of 50°C and 60°C when using safflower oils as

feedstock. The reaction time decreased proportionally as an increase in NaOH

concentration from 1 to 2 wt.%, but soaps were formed when the NaOH concentration

29

was above 2 wt.%. Rashid and Anwar (2008) evaluated the effect of KOH, NaOH,

potassium methoxide (CH3OK), and CH3ONa and their concentrations on the

transesterification of safflower oils. In their study, CH3ONa exhibited the highest yield of

methyl esters.

2.3.2.4 Effect of the reaction temperature

The reaction rate of the transesterification process is strongly influenced by the

reaction temperature. Increasing temperature can enhance the solubility between two

miscible phases and create much interfacial surface area for the transesterification

reaction. Generally, the transesterification was conducted near the boiling point of

alcohol at atmospheric pressure. Freedman, Pryde et al. (1984) found that at temperatures

of 60°C, 75°C, and 114°C, ester conversions of 96% to 98% were obtained by

transesterifying refined oils with methanol, ethanol, and butanol for one hour using 0.5%

CH3ONa as catalyst. However, Leung and Guo (2006) found a higher temperature can

decrease the viscosities of feedstock oils and increase the reaction rate of

transesterification. In addition, higher temperature will accelerate the side saponification

reaction of triglycerides. Rashid and Anwar (2008) recommended that the optimum

temperature for methanolysis of safflower oils is 60°C, and a conversion of 98% was

achieved after 120 min.

2.4 Process of Biodiesel Production from Low Quality Feedstocks

As discussed before, the cost of high quality feedstocks accounts 70-95% of the

total biodiesel production cost. Therefore, an alternative economic approach for reducing

the biodiesel production cost is to use another affordable feedstock such as low quality

30

feedstocks. A number of low quality feedstocks can be used for biodiesel production, for

example, spoiled soybeans, beef and pork tallow, recycled restaurant frying oils, and by-

products such as soap stock from other processes involving vegetable oils.

Canakci and Van Gerpen (2001) developed a two-step process to produce fuel–

quality biodiesel by using low quality feedstocks. The process includes an acid

esterification followed by an alkaline transesterification. Their results showed the

esterification as a pre-treatment process could successfully decrease the acid value of

yellow and brown grease to less than 2 mgKOH/g, but a higher molar ratio and longer

reaction time were needed than in those using simulated low quality feedstocks.

Al-Widyan and Al-Shyoukh (2002) studied the transesterification of waste

vegetable oils by using acid catalysts (HCl and H2SO4). Their results showed that by

using a high catalyst concentration (1.5-2.25 M), biodiesel could be produced in a shorter

reaction time and had a lower specific gravity than that using a low catalyst concentration.

They concluded that the optimum reaction condition in the transesterification process was

2.25 M H2SO4 with 100% excess ethanol.

Zhang, Dube et al. (2003) carried out a simulation process for comparing the

alkali-catalyzed and acid-catalyzed biodiesel production processes when using waste

cooking oil. The alkali-catalyzed process reduced the raw material cost, but it was a very

complex process with a great number of equipment pieces due to the pre-treatment of

FFA. The acid-catalyzed process had less equipment pieces but required a large amount

of methanol.

31

Ghadge and Raheman (2005) developed a two-step pre-treatment process

including esterification followed by an alkali-transesterification to produce biodiesel

from mahua oils (Madhuca indica), which contain 19% FFA by weight of oil. A yield of

98% mahua biodiesel was obtained, which has comparable fuel properties with diesel and

meets the American and European standards of biodiesel.

Ramadhas, Jayaraj et al. (2005) also developed a two-step process (an acid

esterification followed by an alkali transesterification) for biodiesel production from

rubber seed oils containing a high level of FFA. Their results showed that the first step

(acid-catalyzed esterification) could reduce the FFA content to less than 2%. The alkali-

catalyzed transesterification process converted the products of the first step to mono-

esters and glycerol.

Zullaikah, Lai et al. (2005) employed a two-step acid-catalyzed methanolysis

process to convert rice bran oils into fatty acid methyl ester. A H2SO4 solution (1-5 wt.%)

was used as acid catalyst. The first step was carried out at 60°C and more than 98% FFA

and less than 35% of oil was reacted in 2 hours. The organic phase of the first step

reaction product was used as a substrate for a second acid-catalyzed methanolysis at

100°C. Through the two-step methanolysis process, more than 98% FAME in the product

was obtained in less than 8 hours.

Zheng, Kates et al. (2006) studied the reaction kinetics of the acid-catalyzed

transesterification of waste frying oils in excess methanol. Their results showed the acid-

catalyzed transesterification reaction of waste frying oils in methanol is essentially a

pseudo-first-order reaction, provided that the methanol/oil molar ratio is close to 250:1 at

32

70°C or in the range of 74:1 and 250:1 at 80°C. Under these conditions, the biodiesel

production could reach 99 ± 1%.

Wang, Liu et al. (2007) designed a new two-step catalysis process for biodiesel

production. In their process, ferric sulphate (Fe2(SO4)3) was utilized to catalyze the

esterification reaction, and then, KOH was added to catalyze the transesterification

reaction. The lowest acid value of waste cooking oils pretreated by Fe2(SO4)3 was 2.10 ±

0.036 mg KOH/g. Their results showed the conversion of FFA in the waste cooking oil

could reach 97.22% in the first step.

33

Chapter 3 Acid-catalyzed Esterification Reaction

3.1 Acid-catalyzed Esterification Experiments

A large number of bench-scale experiments were conducted to investigate the

reaction kinetics of the Acid-catalyzed esterification reaction. A mixture of virgin canola

oil and pure oleic acid /pure linoleic acid was used as a substrate of a low quality

feedstock. This chapter provides details of the experimental apparatuses, experimental

procedures, sample analyses, and data analyses.

3.1.1 Materials

The vegetable oil used in the experiments was the “No Name” brand Canola oil

purchased from local grocery store. The oil tested had a FFA content of less than 0.015

wt.%. Methanol (purity: 99.98%) and H2SO4 (purity: 95-98%) for reaction were

purchased from Fisher Scientific (Ottawa, Ontario). Free fatty acids of oleic acid (purity

≥ 90%) and linoleic acid (purity ≥ 90%) were bought from Sigma-Aldrich (Oakville,

Ontario). The purities and suppliers of chemicals used in the experiments are listed in

Table 3.1.

3.1.2 Experimental Setups

Figures 3.1 and 3.2 show a schematic diagram and photographs of the

experimental setup designed for the esterification reaction. The esterification reaction

system consisted of

(1) one 500 mL bench-scale reactor;

34

Table 3.1: Purities and suppliers of chemicals

Chemical name Supplier Purity

H2SO4 Fisher Scientific 95-98%

Linoleic acid Sigma-Aldrich 90%

Methanol Fisher Scientific 99.98%

Oleic acid Sigma-Aldrich 90%

35

Figure 3.1: Schematic diagram of experimental setup for acid-catalyzed esterification

reaction

Water Jacket

Mechanical Stirrer

Water Bath

Thermostatic

Temperature

Indicator

Condenser

Impeller

36

Figure 3.2: Photographs of the esterification experimental setup (Original in color)

Reactor Water bath

37

(2) one reflux condenser, which was connected to the reactor in order to prevent

material loss from vaporization;

(3) one warmer warming jacket to maintain the reaction temperature with an

accuracy of ±1°C;

(4) one water bath to adjust to a desired temperature;

(5) one mechanical stirrer to provide a desired mixing intensity;

(6) one temperature couple to monitor the reaction temperature at the

liquid/liquid interface; and

(7) one stopper for sample collections.

3.1.3 Experimental Procedure and Conditions

In the esterification experiment, pure oleic or linoleic acid was used as a

representative of FFA. The virgin canola oil was mixed with pure oleic acid or linoleic

acid as a substrate of a low quality feedstock containing different levels of FFA. The

esterification of low quality feedstock was performed in a 500 mL bench-scale reactor.

Prior to the reaction, 250 mL low quality oil was added into the reactor. An impeller was

placed in the middle of the oil phase and set at a particular mixing speed in order to keep

the interface between the two phases undisturbed. Meanwhile, a known amount of H2SO4

(catalyst) was mixed with 93 mL methanol. With this amount of methanol, the

concentration of methanol was excessively larger than that of FFA (over 40 times), in

order to drive the reversible reaction equilibrium towards the formation of ester and

eliminate the impact of the concentration of methanol on the reaction rate. The

catalyst/methanol mixture was heated to the reaction temperature in a water bath. In order

to keep the two-phase interface undisturbed, a separating funnel was used to smoothly

add the preheated catalyst/methanol mixture into the reactor. The reaction temperature

38

was controlled by the water bath. The reaction was timed until it reached its equilibrium.

During the experiment, samples were collected from the oil phase using a 15 mL syringe

at different time intervals, transferred into 15 mL test tubes, and then immersed in cold

water at 4°C to quench the reaction immediately. For better separation of the final

mixture, the samples were centrifuged for 5 min at 3000 rpm, and, then, the top layer

sample was collected and sent for analysis. Figure 3.3 illustrates the esterification

experimental procedure in steps. Except for the sample analysis, all the experiments were

conducted in the fume hood for safety purposes. Experimental conditions for the acid-

catalyzed esterification are listed in Table 3.2 .

3.1.4 Analytical Methods

FFA content was determined by colour-indicator titration from the Standard Test

Method for Acid and Base Number (ASTM D 974).

p-naphtholbenzein was used as an indicator in an isopropanol/toluene mixture.

The sample was titrated against a 0.1 mol/L potassium hydroxide (KOH) solution. The

titration endpoint was determined when the colour of the sample changed from orange to

green. The acid number and FFA conversion were calculated as follows:

(g) weight sample

(mg/mmol) 56.1(mmol/mL) )(

sample g

KOH mg valueAcid

KOHNmLKOHvolume

[3.1]

100(%)conversionFFA

i

ti

A

AA [3.2]

where Ai= the initial acid value; and At= the acid value at a certain reaction time.

39

Figure 3.3: Experimental procedure for the acid-catalyzed esterification

Titration

Dried

biodiesel

Centrifuge

Reaction

mixture

Quench

Reaction

Reaction

mixture

Methanol/H2SO4

Esterification

reaction

Simulated low quality oil

40

Table 3.2: Experimental conditions for the acid-catalyzed esterification reaction

Experimental Parameter Condition

Acid value range (mg KOH/g) 4-38

Methanol (mL) 93

Oil + oleic acid /linoleic acid (mL) 250

H2SO4 (wt.%) 1, 2, 3

Reaction temperature (°C)* 35, 50, 62

* Accuracy = 1.0°C

41

3.2 Results and Discussion

In a pseudo-homogenous reaction system, the hydrodynamic effect was normally

ignored. However, the hydrodynamic effect was still found to be significant and affect

the overall reaction rate in the heterogeneous reaction system. The esterification reaction

is a heterogeneous reaction of two immiscible phases so its overall reaction rate is

affected by the hydrodynamic effect and chemical reaction. In this study, a new reaction

system of esterification reaction was developed considering both hydrodynamic effect

and chemical reaction. Based on this reaction system, a new experimental setup was

designed to achieve a chemical kinetically-controlled reaction system, in which the

hydrodynamic effect was eliminated under particular experimental conditions. By using

this reaction system, a number of experiments were carried out. The parametric effects

including temperature, catalyst concentration, and initial FFA concentration on the FFA

conversion were examined. In addition, the chemical reaction rate constant and activation

energy were also determined.

3.2.1 Design of a New Reaction System

According to the two-film theory, two factors contribute to the hydrodynamic

effect on the heterogeneous reaction: one is the mass transfer resistance between the

reactant bulk and the reaction interface and the other is the interfacial surface area

available for the reaction. Providing improved mixing in the reaction system can not only

enhance the interchange of the reactant between the interface and the bulk, but also

increase the number of droplets and decrease their dimensions, thereby increasing the

interfacial surface area. In previous esterification reaction systems, the reaction kinetics

were only studied under high agitation speeds, in which the hydrodynamic limitation was

42

considered negligible. However, the correlation between the agitation speed and the

hydrodynamic effect in the systems was not fully ascertained. Since it is difficult to

quantify the mass transfer resistance and the interfacial surface area in a vigorously

mixed reaction system, the new reaction system was controlled in a gentle mixed system

(undisturbed interface), in which the interfacial surface area is fixed and equal to the

undisturbed two-phase interface. Since the interfacial surface area is fixed, the

hydrodynamic effect on the esterification reaction is only dependent on the mass transfer

resistance, which changes with the agitation in the system.

Three modes of agitation, including no agitation in the system, agitation in the oil

phase, and agitation in the methanol phase, were evaluated by using the esterification

experimental setup as shown in Figure 3.1. The agitation speed was controlled to keep the

two-phase interface undisturbed. As shown in Figure 3.4, the FFA conversion slightly

increased up to 3% in 60 min in the cases of agitation in the methanol phase and no

agitating in the system.

However, in the case of agitation in the oil phase, the FFA conversion increased

up to 40% in 60 min. The above results show that the FFA conversion did not increase

due to the agitation in the methanol phase compared to that of no agitation in the system.

This indicates that mass transfer resistance does not exist in the methanol phase, since it

is almost an instantaneous reaction, so the esterification reaction takes place at the mass

transfer interface, i.e., the two-phase interface. Mass transfer resistance exists in the oil

phase because the FFA conversion greatly increased due to the agitation

43

Figure 3.4: Effect of position of the mechanical impeller on the FFA conversion

(T=50°C, H2SO4 concentration =3 wt.%, FFA content =36 mgKOH/g)

0

20

40

60

80

100

0 10 20 30 40 50 60

FF

A C

on

versi

on

(%

)

Reaction Time (min)

no agitation at two phase

agitate oil phase 50 rpm

agitate methanol phase at 50 rpm

44

in the oil phase. When the oil phase was agitated, the mass transfer resistance of FFA in

the oil phase was reduced due to the enhanced interchange of FFA between the interface

and the oil bulk, resulting in a much higher increase in FFA conversion with time

compared to that of no agitation in the system.

Under a particular agitation speed limit in the oil phase, the interfacial surface

area available for the reaction remains constant as the agitation speed increases and is

equal to the surface area of the two-phase interface as long as the interface remains

undisturbed. When the interface is disturbed as the agitation speed increases, the nonpolar

oil phase in a larger volume will become a continuous phase and the methanol phase will

become droplets of dispersed phase. The interfacial surface area for the reaction will

change and be determined by the droplet size of the dispersed phase. Figure 3.5 illustrates

the change of the interface state between the oil and methanol phases as the agitation

speed increases.

Since the mass transfer resistance of FFA only exists in the oil phase and the

interfacial surface area equals the fixed surface area of the two-phase interface at

particular agitation speeds, the mass transfer effect on the overall reaction rate can be

easily evaluated in the esterification reaction system. A number of experiments were

carried out to study the FFA conversion at different agitation speeds. The agitation speed

was controlled in a range of 0 to 200 rpm, within which the state of the two-phase

interface in the system shifted from undisturbed to disturbed by visual observation.

As shown in Figure 3.6, three different reaction stages are found as the agitation

speed increases at 50°C. In stage I (0 to 80 rpm), the two-phase interface was not

disturbed; the interfacial surface area equaled the surface area of the interface.

45

Figure 3.5: Change of the interface state with increasing agitation speeds

Low High Agitation Speed

46

When increasing the agitation speed, the film thickness of the oil phase decreased,

thereby decreasing the mass transfer resistance of FFA. As a result, the FFA conversion

increased as the agitation speed increased. In stage II (80 to 115 rpm), the interface was

still not disturbed; the interfacial surface area equaled the surface area of the interface.

However, the FFA conversion remained constant as the agitation speed increased. This is

because the film thickness in the oil phase became negligible as the agitation speed

increased. Thus, the mass transfer resistance of FFA in the oil phase had no impact on the

FFA conversion. The esterification reaction was totally controlled by chemical reaction.

In stage III (larger than 115 rpm), the interface was disturbed and the FFA conversion

increased again as the agitation speed increased. The reaction system became a

vigorously mixed system in which both the interfacial surface area and mass transfer

resistance of FFA in the mixed system imposed great impacts on the FFA conversion.

The same results were obtained at 62°C and 35°C as shown in Figures 3.7 and 3.8. The

FFA conversion remained constant (stage II) when the agitation speeds were in the

ranges of 80-100 rpm and 80-123 rpm, respectively.

A new reaction system was developed for esterification reaction, considering both

hydrodynamic effect and chemical reaction. In order to eliminate the hydrodynamic

effect in the reaction system, the new system needed to meet the following conditions:

(1) no agitation exists in the methanol phase because mass transfer resistance is

negligible in the methanol phase;

47


concentration=3wt.%, FFA content=37 mgKOH/g)

0

20

40

60

80

100

0 20 40 60 80 100 120 140 160 180 200 220

FF

A C

ov

nersi

on

(%

)

Agitation Speed (rpm)

Reaction Time

20 min

40 min

60 min

I

II

III

48


concentration=3 wt.%, FFA content=37 mgKOH/g)

0

20

40

60

80

100

0 20 40 60 80 100 120

FF

A C

on

versi

on

(%

)


Reaction Time

20 min

40 min

60 min

II

49


concentration=3 wt.%, FFA content=37 mgKOH/g)

0

20

40

60

80

100

0 20 40 60 80 100 120 140

FF

A C

on

versi

on

(%

)


Reaction Time

0 min

20 min

40 min

60 min

III

50

(2) agitation exists in the oil phase because mass transfer resistance of FFA exists

in the oil phase.

(3) interface between the oil phase and methanol phase remains undisturbed

when mixed, i.e., the two phases are separate, and

(4) agitation speed in the oil phase is controlled at 80 rpm in this study, in which

the FFA conversion does not change as the agitation speed increases. The

mass transfer resistance of FFA in the oil phase is negligible.

and also two assumptions need to be made

(1) all reacted FFA is converted to biodiesel; and

(2) the reverse reaction of esterification is negligible because the excessive

amount of methanol can drive the reaction forward.

In Chapter 2, the overall reaction rate of esterification is obtained with Equation

2.9:

OFFA

RXOFFA

RX

FFA C

kak

rV

F

'

11

1 [2.9]

In this equation, the overall reaction rate is affected by the concentration of FFA in the oil

bulk, the mass transfer resistance in the oil film, and the chemical reaction rate. In this

study, the hydrodynamic effect is negligible and can be ignored. The chemical reaction

rate of the esterification is then:

FFARX Ckr ' [3.3]

51

(Note: for a simplified notation, we will use CFFA as the concentration of the FFA in the

oil phase instead of CFFA-o hereafter.)

By integrating Equation 3.3, the pseudo-first order rate constant can be obtained

from Equation 3.4, which is the slope of the graph, by plotting FFA

FFA

C

C 0ln against time.

tkC

CRX

FFA

FFA 'ln 0 [3.4]

where CFFA0 = initial concentration of FFA; CFFA = concentration of FFA at time t.

The Arrhenius equation was used to study the influence of temperature on specific

reaction rates. Once the RXk' value is determined at different temperatures, the activation

energy for the esterification can be estimated by using the Arrhenius equation:

RT

Ea

RX Aek ' [3.5]

where A=frequency factor; Ea=activation energy; R=universal gas constant; T=

temperature.

3.2.2 Determination of Reaction Rate Constant and Activation Energy

Based on the previous discussions and results, an experimental setup was

designed for studying the reaction kinetics of esterification. As shown in Figure 3.1, an

impeller was placed in the oil phase. According to previous results in this chapter, the

agitation speed was controlled at 80 rpm to eliminate the hydrodynamic effect on the

reaction. The kinetics study was carried out under different experimental conditions

including temperature, catalyst concentration, and initial concentration of FFA. The

reaction rate constant and activation energy were estimated.

52

According to Equation 3.4, the relationship between

FFA

FFA

C

C 0ln and reaction time (t)

is linear. The value of the rate constant (k´RX) equals the slope of the linear regression

trendline. Thus, FFA

FFA

C

C 0ln is plotted against t in different experimental conditions. As

shown in Figures 3.9-3.11, the resulting data fits pseudo-first order kinetic behaviour.

The high correlation coefficients (R2) of the liner equation indicate that there is a first

order dependence of the esterification reaction catalyzed by H2SO4. The rate constants

(k´RX) under different experimental conditions, including temperature, catalyst

concentration, and initial concentration of FFA, were calculated and are shown in Tables

3.3-3.5.

The activation energy and frequency factor of the esterification reaction were

estimated using Equation 3.5. By taking the natural logarithm of both sides of Equation

3.5, Equation 3.6 is obtained:

RT

EaAk RX ln'ln [3.6]

Equation 3.6 shows lnk'RX and 1/T is a linear relationship with a slope of -Ea/RT

and an intercept of lnA. Since the values of k'RX at different temperatures were determined

in Table 3.3-3.5, the Arrhenius plot of lnk'RX versus 1/T is made in Figure 3.12 at three

H2SO4 concentrations (1, 2, and 3 wt.%). By performing a linear regression of lnk'RX

versus 1/T, the activation energy and frequency factor are determined from the

53

(a)

(b)

(c)


FFA

C

C 0ln as a function of time (H2SO4 concentration =3 wt.%)

(a)T=35°C (b) T=50°C (c) T=62°C

0

0.5

1

1.5

2

2.5

3

0 60 120 180 240 300 360ln

(CF

FA

0/C

FF

A)

Reaction Time (min)

Initial FFA Content

35.47 mgKOH/g

13.32 mgKOH/g

5.90 mgKOH/g

0

0.5

1

1.5

2

2.5

3

0 60 120 180 240

ln(C

FF

A0/C

FF

A)

Reaction Time (min)

Initial FFA Content

36.39 mgKOH/g

14.14 mgKOH/g

4.86 mgKOH/g

0

0.5

1

1.5

2

2.5

3

0 60 120 180 240

ln(C

FF

A0/C

FF

A)

Reaction Time (min)

Initial FFA Content

35.62 mgKOH/g

15.30 mgKOH/g

4.74 mgKOH/g

54

(a)

(b)

(c)


FFA

C


(a)T=35°C (b) T=50°C (c) T=62°C

0

0.5

1

1.5

2

2.5

3

0 60 120 180 240 300 360

ln(C

FF

A0/C

FF

A)

Reaction Time (min)

Intial FFA Content

37.33 mgKOH/g

5.20 mgKOH/g

0

0.5

1

1.5

2

2.5

3

0 60 120 180 240 300

ln(C

FF

A0/C

FF

A)

Reaction Time (min)

Intial FFA Content

37.33 mgKOH/g

4.72 mgKOH/g

0

0.5

1

1.5

2

2.5

3

0 60 120 180 240 300

ln C

FF

AO

/CF

FA

Reaction Time (min)

Intial FFA Content

37.33 mgKOH/g

4.72 mgKOH/g

55

(a)

(b)

(c)


FFA

C


(a)T=35°C (b) T=50°C (c) T=62°C

0

0.5

1

1.5

2

2.5

3

0 60 120 180 240 300 360 420

ln (C

FF

A0/C

FF

A)

Reaction Time (min)

Initial FFA Content

35.87 mgKOH/g

14.60 mgKOH/g

5.27 mgKOH/g

0

0.5

1

1.5

2

2.5

3

0 60 120 180 240 300 360

ln(C

FF

A0/C

FF

A)

Reaction Time (min)

Initial FFA Content

36.09 mgKOH/g

4.74 mgKOH/g

0

0.5

1

1.5

2

2.5

3

0 60 120 180 240 300

ln(C

FF

A0/C

FF

A)

Reaction Time (min)

Initial FFA Content

36.09 mgKOH/g

4.74 mgKOH/g

56

Table 3.3: Reaction rate constants at 3 wt.% H2SO4

Initial FFA content

(mgKOH/g)

Temperature

(°C)

Reaction rate constant

k´RX (min-1

) ×102

Correlation

coefficients

(R2)

35.47

35

0.53 0.9376

13.32 0.53 0.9910

5.90 0.47 0.9559

36.39

1.07 0.9583

14.14 50 1.12 0.9970

4.86

0.95 0.9920

35.62

62

1.63 0.9801

15.30 1.83 0.9910

4.74 1.77 0.9862

57


Initial FFA content

(mgKOH/g)

Temperature

(°C)


k´RX (min-1

) ×102

Correlation

coefficients

(R2)

37.33 35

0.59 0.9855

5.20 0.54 0.9968

37.33 50

0.97 0.9889

4.72 0.93 0.9941

37.33 62

1.48 0.9885

4.72 1.61 0.9876

58


Initial FFA content

(mgKOH/g)

Temperature

(°C)


k´RX (min-1

) ×102

Correlation

coefficients

(R2)

35.47 35

0.49 0.9537

14.6 0.50 0.9384

5.90

0.37 0.9884

36.09 50

0.76 0.9819

4.86 0.80 0.9914

36.09 62

1.25 0.9836

4.74 1.29 0.9987

59

Figure 3.12: Arrhenius plot of lnk'RX against 1/T (Esterification of oleic acid)

-6

-5

-4

-3

-2

-1

0

0.00295 0.003 0.00305 0.0031 0.00315 0.0032 0.00325 0.0033

lnk

' RX

1/T

H2SO4 Concentration

1wt.%

2wt.%

3wt.%

60

slope and intercept of the regression trendline, respectively. Results including the

activation energy, frequency factor, and the correlation coefficient (R2) are shown in

Table 3.6. The correlation coefficient is very close to one, which indicates a very good

linear relationship between lnk'RX and 1/T. Additionally, the frequency factor increases as

the H2SO4 concentration increases. The high frequency factor, which is a measure of

collisions between reactants, indicates that the esterification reaction is more favoured at

3 wt.% H2SO4 than those at 2 wt. % and 1 wt.% H2SO4.

3.2.3 Parametric Effects on the Esterification Reaction

3.2.3.1 Effect of Temperature

The reaction temperature is an important operating parameter affecting the

reaction rate. The effect of temperature on the esterification reaction was studied by using

a 3 wt.% H2SO4. Figure 3.13 shows the FFA conversion as a function of time at three

different temperatures including 35°C, 50°C, and 62°C. The initial FFA contents used in

the experiment are 35-38 mg KOH/g, as in Figure 3.13 (a), and 13-16 mg KOH/g, as in

Figure 3.13 (b). The results show that the reaction temperature has a great impact on FFA

conversion. An increase of temperature caused FFA conversion to increase until the

reaction reached equilibrium. It took less time to reach the same conversion at a high

temperature than at a relatively low temperature. For example, when the initial FFA

concentration was 35-38 mg KOH/g, it took approximate 500 min at 35°C, 200 min at

50°C, and less than 100 min at 62°C to get 80% FFA conversion. This indicates that the

esterification reaction rate increased as the reaction temperature increased, and the

61

Table 3.6: Activation energy in the esterification reaction of oleic acid

H2SO4 concentration

(wt.%)

Ea

(kJ/mol) A

Correlation coefficients

(R2)

1 32.48 1436.55 0.9961

2 31.96 1450.99 0.9956

3 39.14 22136.87 0.9999

62

(a)

(b)

Figure 3.13: Effect of temperature on the FFA conversion (H2SO4 concentration=3 wt.%)

(a) Initial FFA content=35-38 mg KOH/g (b) Initial FFA content=3-16 mg

KOH/g

0

20

40

60

80

100

0 200 400 600 800

FF

A C

on

versi

on

(%

)

Reaction Time (min)

Temperature

35℃

50℃

62℃

0

20

40

60

80

100

0 100 200 300 400 500

FF

A C

on

versi

on

(%

)

Reaction Time (min)

Temperature

35℃

50℃

62℃

63

reaction rate decreased with time. This can be explained by the reaction rate constants

calculated at three different temperatures. As seen in Figure 3.14, an increase of

temperature leads to an increased reaction rate constant in a proportional manner. Since,

at a higher temperature, the FFA molecules and alcohol molecules have more thermal

energy and the collision frequency between them is increased with the elevated

temperature, an increase of temperature causes the reaction rate constant to increase,

leading to an increase of the reaction rate. Therefore, it is preferable that the esterification

reaction proceed at a relatively high temperature in order to obtain a high reaction rate.

3.2.3.2 Effect of Catalyst Concentration

Figure 3.15 shows the FFA conversion profiles of the esterification reaction using

three different concentrations of H2SO4: 1 wt.%, 2 wt.%, and 3 wt.% at three reaction

temperatures of 35°C, 50°C, and 62°C. The results show that at relatively high

temperatures, such as 50°C or 62°C, the FFA conversion increased as the catalyst

concentration increased over the same time until the reaction reached equilibrium. For

example, at 62°C, only 30% FFA conversion was obtained in 60 min when the

concentration of H2SO4 was 1%, but when the concentration of H2SO4 was increased to 2

wt.%, the FFA conversion increased to 50% in 60 min. Also, when the concentration of

H2SO4 was 3 wt.%, the FFA conversion reached as high as 70%. Then, after the reaction

reached equilibrium, an increase of H2SO4 concentration could not lead to a further

increase in the FFA conversion. The maximum FFA conversions were almost the same

when comparing three different concentrations of H2SO4. This indicates that an increased

64

Figure 3.14: Effect of temperature on the reaction rate constant (H2SO4 concentration=3

wt.%)

0.00

0.20

0.40

0.60

0.80

1.00

1.20

1.40

1.60

1.80

2.00

0 10 20 30 40 50 60 70

k' R

X( m

in-1

)×1

02

Reaction Temperature (℃)

Initial FFA Content

35-38 mg KOH/g

13-16 mg KOH/g

65

(a)

(b)

(c)

Figure 3.15: Effect of catalyst concentration on the FFA conversion (Initial FFA

content=35-38 mg KOH/g) (a) T=35°C (b) T=50°C (c) T=62°C

0

20

40

60

80

100

0 60 120 180 240 300 360

FF

A C

on

ver

sion

(%

)

Reaction Time (min)

H2SO4 Concentration

1 wt.%

2 wt.%

3 wt.%

0

20

40

60

80

100

0 60 120 180 240 300 360

FF

A C

on

ver

sion

(%

)

Reaction Time (min)

H2SO4 Concentration

1 wt.%

2 wt.%

3 wt.%

0

20

40

60

80

100

0 60 120 180 240 300

FF

A C

on

ver

sion

(%

)

Reaction Time (min)

H2SO4 Concentration

1 wt.%

2 wt.%

3 wt.%

66

concentration of catalyst could effectively reduce the reaction time but could not change

the maximum conversion efficiency. However, at a low temperature, such as 35°C, the

FFA conversion remained almost constant over the same reaction time when the

concentration of H2SO4 increased from 1 wt.% to 3 wt.%. The different effects of the

catalyst concentration on the FFA conversion resulted in different reaction rate constants.

Figure 3.16 shows the reaction rate constant as a function of the catalyst concentration at

35°C, 50°C, and 62°C. It is clear that at 50°C and 62°C, an increase of catalyst

concentration caused the reaction rate constant to increase, resulting in an increased

reaction rate. However, at 35°C, the change of reaction rate constant with the catalyst

concentration is very small, leading to an unchanged reaction rate in different catalyst

concentrations.

3.2.3.3 Effect of FFA Content

The effect of initial FFA content on the reaction rate was investigated at two

different temperatures: 50°C and 62°C. Figure 3.17 shows the profiles of acid value as a

function of reaction time (t) in three different initial FFA contents: 35 mgKOH/g, 14

mgKOH/g, and 4 mgKOH/g. The results illustrate that the initial FFA content had a

significant impact on the reaction rate, which is the slope of the CFFA-t plot. An increase

in the initial FFA content led to an increased reaction rate until the reaction reached

equilibrium. This result can be simply explained by Equation 3.3. Since the reaction is an

equilibrium reaction, increasing the initial FFA content drives the equilibrium forward

67

Figure 3.16: Effect of catalyst concentration on reaction rate constant (Initial FFA

content =35-38 mgKOH/g)

0.00

0.20

0.40

0.60

0.80

1.00

1.20

1.40

1.60

1.80

0 1 2 3 4

k' R

X ( m

in-1

)×1

02

H2SO4 Concentration (wt.%)

Temperautre

35℃

50℃

62℃

68

(a)

(b)

Figure 3.17: Change of FFA content as a function of reaction time

(a) T=50°C (b) T=62°C

0

20

40

0 60 120

FF

A C

on

ten

t (m

gK

OH

/g)

Reaction Time (min)

Initial FFA Content

4.86 mgKOH/g

14.14 mgKOH/g

36.39 mgKOH/g

0

20

40

0 60 120

FF

A C

on

ten

t (m

gK

OH

/g)

Reaction Time (min)

Initial FFA Content

4.74 mgKOH/g

15.30 mgKOH/g

35.62 mgKOH/g

69

and increases the reaction rate. Figure 3.18 shows the reaction rate constant as a function

of the initial FFA content. The values of the reaction rate constants were very close and

have no clear trend as the FFA content increased. This indicates the initial FFA content

had no impact on the reaction rate constant.

3.2.3.4 Effect of the Type of FFA

Previous studies on the esterification reaction were conducted by using an oleic

acid as FFA, which was added into canola oil for simulating a low quality feedstock.

Since linoleic acid is another major component of FFA in low quality feedstocks,

esterification reactions using a mixture of linoleic acid and canola oil as feedstock were

also studied and the activation energies were also determined using the same method as

the previous studies with the oleic acid.

Figure 3.19 shows the Arrhenius plot of lnk'RX against 1/T at two different H2SO4

concentrations: 1 wt.% and 3 wt.%. By performing a linear regression of the lnk'RX -1/T

plot, the activation energy and the frequency factor were determined. As seen in Table

3.7, the esterification reaction of linoleic acid had a very similar activation energy at the

two different catalyst concentrations. Similarly to the esterification reaction of oleic acid,

for the esterification reaction of linoleic acid, the catalyst concentration had no impact on

the activation energy. By comparing the activation energy of the esterification reaction of

oleic acid in Table 3.6 with that of linoleic acid in Table 3.7, it shows that the

esterification reactions of oleic acid and linoleic acid have very similar activation

energies

70

Figure 3.18: Effect of the initial FFA content on the reaction rate constant

0.00

0.20

0.40

0.60

0.80

1.00

1.20

1.40

1.60

1.80

2.00

0 5 10 15 20 25 30 35 40

k' R

X ( m

in-1

)×1

02

Initial FFA Content (mgKOH/g)

Temaperture

50℃

62℃

71

.

Figure 3.19: Arrhenius plot of lnk΄RX against 1/T (Esterification of linoleic acid)

-6

-5

-4

-3

-2

-1

0

0.00295 0.003 0.00305 0.0031 0.00315 0.0032 0.00325 0.0033

lnk

' RX

1/T

H2SO4Concentration

1wt.%

3wt.%

72

Table 3.7: Activation energy of the esterification reaction using linoleic acid

H2SO4 concentration

(wt.%)

Ea

(kJ/mol)

A

Correlation coefficients

(R2)

1 31.58 1344.261 0.9584

3 34.23 4072.857 0.9770

73

Mixtures of oleic acid and linoleic acid in different ratios were also investigated

as FFA in the canola oil in the esterification reaction. The experiments were conducted

using 3 wt.% H2SO4 (as catalyst at 65℃. Figure 3.20 shows the reaction rate constants in

different ratios of oleic acid to linoleic acid. It indicates that the reaction rate constant did

not change with the ratio of oleic acid to linoleic acid. Because the esterification reactions

of oleic acid and linoleic acid have very similar activation energies, the unchanged

reaction rate constant in different ratios of oleic acid to linoleic acid reveals that the oleic

acid and linoleic acid have the same reaction behavior in the esterification reaction.

74

Figure 3.20: Comparison of the reaction rate constants by using mixed FFA with

different ratios of oleic acid versus linoleic acid (T=62°C, H2SO4

concentration =3 wt.%)

0

0.5

1

1.5

2

2.5

0:1 1/4:3/4 1/2:1/2 3/4:1/4 1:0

k'

RX

(m

in-1

) ×

10

2

Oleic acid: Linoleic acid

75

Chapter 4 Alkali-catalyzed Transesterification Reaction

4.1 Alkali-catalyzed Transesterification Experiments

A series of bench-scale experiments were carried out to evaluate the reaction

kinetics on alkali-catalyzed transesterification. Virgin canola oil was used as a source of

high quality feedstocks, and it reacted with methanol in the presence of NaOH to yield

biodiesel. Details of the experimental apparatus, experimental procedure, sample analysis,

and data analysis are provided below.

4.1.1 Materials

Methanol (purity: 99.98%) and sodium hydroxide-NaOH (purity: 99.1%) were

purchased from Fisher Scientific (Ottawa, Ontario). Hexane (purity: 99.99%) and

anhydrous sodium sulfate-Na2SO4 (purity: 99.9%) used for sample preparation and

analysis were also obtained from Fisher Scientific. Oleic acid (purity ≥ 99%) and methyl

heptadecanoate (internal standard for gas chromatography with purity of 99.5%) were

obtained from Sigma-Aldrich (Oakville, Ontario). The purities and suppliers of chemicals

used in the experiment are listed in Table 4.1.

4.1.2 Experimental Setups

Figures 4.1 and 4.2 show a schematic diagram and photographs of the

experimental setup designed for the alkali-transesterification study. The setup was

designed to facilitate three consecutive steps of operation: preheating methanol,

76

Table 4.1: Purities and suppliers of chemicals

Chemical name Supplier Purity

Anhydrous Na2SO4 Fisher Scientific 99.9%

Hexane Fisher Scientific 99.99%

Linoleic acid Sigma-Aldrich 90%

Methanol Fisher Scientific 99.98%

Methyl heptadecanoate Sigma-Aldrich 99.5%

Oleic acid Sigma-Aldrich 90%

NaOH Fisher Scientific 99.1%

77

Figure 4.1: Schematic diagram of the alkali-catalyzed transesterification experimental

setup

Condenser

Sampling point

Three necked flask Stirring hot plates

Thermometer

Cold water out

Cold water in

78

Figure 4.2: Photographs of the alkali-catalyzed transesterification experimental setup

(Original in color)

Reactors connected with

condensers Separating funnels Water bath for preheating

methanol

79

transesterification reaction, and product separation. The apparatus used for

transesterification reaction includes

(1) one 125 mL glass reactor;

(2) one reflux condenser, which was connected with the reactor in order to

prevent material loss from vaporization;

(3) one Thermo Scientific Super-Nuova

multi-position stirring hot plate (Cole-

Parmer Canada Inc.) to control the reaction temperature and agitation speed;

(4) one magnetic stir to supply a desired mixing intensity; and

(5) one thermometer to measure the reaction temperature.

4.1.3 Experimental Procedure and Conditions

The transesterification reaction took place in a 125 mL glass reactor. Prior to the

reaction, 50 mL of virgin canola oil was added into the reactor, which was placed on a

stirring hot plate. The reaction temperature and agitation speed of the stir were adjusted

by the stirring hot plate to meet the desired experimental conditions. At the same time, a

known amount of NaOH (catalyst) was mixed with a pre-measured amount of methanol.

The mixture of catalyst and methanol was then heated to the reaction temperature in a

water bath.

The transesterification reaction took place by introducing the mixture of catalyst

and methanol to the canola oil in the reactor. The reaction temperature and agitation

speed were controlled for a particular period of time until the reaction reached its

equilibrium. The final reaction mixture was transferred to a separating funnel and kept

undisturbed for 12 hours to separate the glycerol phase and crude biodiesel phase. The

separated glycerol phase in the bottom layer was disposed from the funnel, and a 10 mL

80

crude biodiesel in the top layer was collected and gently washed with 10 mL of de-

ionized water three times to remove the unreacted catalyst, methanol residual, glycerol,

and trace amount of soaps. The washed biodiesel was then dried over sodium sulfate

(Na2SO4) and injected into a gas chromatograph with a mass spectrometry detector

(GC/MS) for analysis of methyl ester concentration. Figure 4.3 illustrates the

transesterification experimental procedure in steps. All the experiments were conducted

in the fume hood for safety purposes. Experimental conditions for the alkali-catalyzed

transesterification are listed in and Table 4.2.

4.1.4 Analytical Methods

A dried and washed sample from the transesterification reaction was analyzed for

the content of FAME, i.e., biodiesel, by using a GC/MS.

The GC/MS was equipped with an Econo-Cap EC-WAX Capillary Column (30.0

m in length × 250 m in diameter × 0.25 m in film thickness). The GC oven was

maintained at 50°C for 3 min, and then heated to 210°C at a rate of 10°C per minute and

held at 210°C for 9 min. The front inlet temperature of the oven was 255°C (splitless-

mode). The carrier gas was helium with a flow rate of 12 mL/min. The analysis of FAME

was carried out by injecting 1.0 L of a sample solution that was prepared by blending

the biodiesel sample with a prepared internal standard of GC, i.e., methyl heptadecanoate.

The FAME content by weight was determined from Equation 4.1:

81

Figure 4.3: Experimental procedure for the alkali-catalyzed transesterification

Crude

biodiesel

Washed

biodiesel

Water

GC/MS

Titration

Dried

biodiesel

Waste Water

Reaction

mixture Canola oil

Methanol/NaOH

Transesterification

Reaction

Phase

Separation

Phase

Separation

Glycerol (Disposal)

Water

wash

Drying

Na2SO4

82

Table 4.2: Experiment conditions for the alkali-catalyzed transesterification reaction

Experimental Parameter Condition

Methanol to Oil (molar ratio) 9:1

NaOH concentration (wt.%) 0.2, 0.6, 1.0

Temperature (°C)* 25, 35, 50, 65

Agitation speed (rpm) 200

* Accuracy = 2.0°C

83

( )

.%i R R R

R

A A C Vwt

A W

[4.1]

where Ai= the peak area from chromatogram of FAME; AR= the peak area from

chromatogram of internal standard; CR= the concentration of the internal standard; VR=

the volume of the internal standard; and W = the total weight of the biodiesel sample.

4.2 Results and Discussion

The main task of this part was to investigate the parametric effects, including

reaction temperature, catalyst concentration, and initial FFA content, on the biodiesel

conversion profile and reaction rate when using NaOH as a catalyst. The biodiesel

conversion performance was evaluated in terms of FAME content (wt.%) of the

reaction product as a function of reaction time. The change in FAME content with

time provided an insight into the effects of these reaction parameters on the biodiesel

conversion rate. It should be noted that an agitation speed of 200 rpm was chosen in

this study because the speed is adequate for facilitating the reaction between oil and

methanol but gentle enough to clearly reveal crucial information on the advance of

biodiesel conversion with the reaction time.

4.2.1 Effect of Reaction Temperature

Figures 4.4-4.6 show the effect of reaction temperature on the FAME content

as a function of reaction time. Three catalyst concentrations, including 0.2, 0.6, and

1.0 wt.%, were studied at four different temperatures: 25°C, 35°C, 50°C, and 65°C. It

is clear that, regardless of the catalyst concentration, raising the reaction temperature

caused the biodiesel conversion to proceed at a greater rate as indicated by a faster

increase in FAME content, especially during the first part of the reaction period. For

84


(Methanol/Canola oil=9:1(molar ratio); mixing speed=200 rpm)

0

10

20

30

40

50

60

70

80

90

100

0 1 2 3 4 5 6 7

Meth

yl

Est

ers

Co

nte

nt

(wt.

%)

Reaction Time (h)

Reaction Temperature

25℃

35℃

50℃

60℃

85



0

10

20

30

40

50

60

70

80

90

100

0 1 2 3 4 5 6 7

Meth

yl

Est

ers

Co

nte

nt

(wt.

%)

Reaction Time (h)


25℃

35℃

50℃

65℃

86



0

10

20

30

40

50

60

70

80

90

100

0 1 2 3 4 5

Meth

yl

Est

ers

Co

nte

nt

(wt.

%)

Reaction Time (h)


25℃

35℃

50℃

65℃

87

instance, in Figure 4.5, at 0.6 wt.% catalyst concentration, a FAME content of 85 wt.%

was achieved within 10 min at 65°C while it took as long as 110 min to reach the

same conversion level at 25°C. The increased rate of conversion with temperature is

probably caused by two main factors: (1) an increase in kinetic reaction rate of

transesterification reaction with temperature and (2) a reduction in viscosity of

feedstock oil with temperature that helps promote ultimate mixing between oil and

methanol. From Figures 4.4-4.6, conversion profiles at lower temperatures (i.e., 25°C,

35°C) appear to have two distinct conversion regions: slow conversion for FAME

content below 20 wt.% and rapid conversion for content above 20 wt.%. The slow

conversion region is an indication of poor mixing between feedstock oil (with low

FAME content) and methanol, which was probably caused by the large difference in

viscosity between these two reactant phases. For instance, kinematic viscosity of

canola oil at 40°C is 36 mm2/s whereas methanol viscosity is only 0.59 mm

2/s (Perry

and Green 1997). As the conversion progressed, the increasing content of FAME

(with kinematic viscosity of 5 mm2/s at 40°C) caused the viscosity of oil phase to

decrease significantly, resulting in improved mixing between oil and methanol, which

in turn promoting rapid conversion as seen in the second part of the conversion

profiles at the low temperatures in Figures 4.4-4.6. In contrast, at a high temperature,

there was no dramatic change in the conversion rate. It appears from the profiles that

the conversion at 50°C and 65°C proceeded rapidly as soon as the reaction started.

This simply demonstrates that the negative impact of viscosity observed at the low

reaction temperature was eliminated at the high temperature, offering an improvement

in the rate of conversion.

In addition to the rate of conversion, Figures 4.4-4.6 also provide important

information on the maximum conversion level where the FAME content reaches the

88

highest value and remains relatively constant as time progresses. They show that the

maximum FAME content ranges from 80 to 90 wt.% regardless of the catalyst

concentration and reaction temperature.

4.2.2 Effect of Catalyst Concentration

Figures 4.7-4.10 show the effect of NaOH concentration on the conversion

profile of the transesterification reaction at four reaction temperatures: 25°C, 35°C,

50°C, and 65°C. In general, an increase in the catalyst concentration offers a higher

rate of conversion, as indicated by the shorter reaction time, for achieving the

maximum level of FAME content. As shown in Figure 4.7, at 25°C, the reaction

required about 5 hours to yield the maximum conversion of 79 wt.% when the catalyst

concentration was 0.2 wt.%. As the catalyst concentration increased to 0.6 wt.% and

even 1.0 wt.%, the reaction time required was reduced to 2 and 1 hours, respectively.

At 35°C, in Figure 4.8, the maximum conversion for 0.2 wt.% catalyst was achieved

within 180 minutes while the catalyst concentrations of 0.6 wt.% and 1.0 wt.% offered

a shorter reaction times of 60 and 45 minutes, respectively. At 65°C, in Figure 4.10,

increasing the concentration of catalyst from 0.2 wt.% to 1.0% caused the reaction

time to decrease from 30 minutes to less than 5 minutes. Therefore, the conversion

rate increased as the concentration of catalyst increased.

89


(Methanol/Canola oil=9:1 (molar ratio); mixing speed=200 rpm)

0

10

20

30

40

50

60

70

80

90

100

0 1 2 3 4 5 6

Meth

yl

Est

ers

Co

nte

nt

(wt.

%)

Reaction Time (h)

NaOH Concentration

0.2 wt.%

0.6 wt.%

1.0 wt.%

90



0

10

20

30

40

50

60

70

80

90

100

0 1 2 3 4

Meth

yl

Est

ers

Co

nte

nt

(wt.

%)

Reaction Time (h)

NaOH Concentration

0.2 wt.%

0.6 wt.%

1.0 wt.%

91



0

10

20

30

40

50

60

70

80

90

100

0 0.5 1 1.5 2

Meth

yl

Est

ers

Co

nte

nt

(wt.

%)

Reaction Time (h)

NaOH Concentration

0.2 wt.%

0.6 wt.%

1.0 wt.%

92



0

10

20

30

40

50

60

70

80

90

100

0 0.5 1 1.5 2

Meth

yl

Est

ers

Co

nte

nt

(wt.

%)

Reaction Time (h)

Catalyst Concentration

0.2 wt.%

0.6 wt.%

1.0 wt.%

93

4.2.3 Effect of FFA Content

As mentioned in the previous chapter, using low quality feedstock such as

waste cooking oil or animal fat for biodiesel production presents a number of

advantages over the conventional virgin vegetable oil since the cost of waste cooking

oil or fat is much lower and their availability is not directly affected by crop growing

variables. However, low quality feedstock contains a large amount of FFA. In the

alkali-catalyzed transesterification process, the FFA in the feedstock can react with

the alkaline catalyst to form undesirable soap products, resulting in a loss of catalyst

as well as a reduction in biodiesel production efficiency. Therefore, in this part of the

study, the effect of FFA content on the reaction conversion was quantified through a

number of experiments. Oleic acid was added into the base canola oil to form

simulated low quality oil containing different levels of acid number.

Figures 4.11-4.13 show the experimental results revealing the effect of FFA

content on the reaction conversion when the catalyst concentrations are 0.2 wt.%, 0.6

wt.%, and 1.0 wt.%. In general, the presence of FFA caused the rate of reaction

conversion to drop as the amount of alkaline catalyst available for transesterification

reaction was reduced or depleted. At a catalyst concentration of 0.2 wt.%, in Figure

4.11, the conversion required as long as 50 minutes to produce 87% FAME product

from a feedstock containing 0.5% FFA (acid value of 0.94 mg KOH/g) whereas it

took only 20 minutes to yield a similar product from the same feedstock containing no

FFA. With higher FFA content (acid value of 4.84 mg KOH/g), there was no FAME

produced in the system even within 120 minutes. This indicates that the catalyst was

completely consumed by FFA and no catalyst was left for the transesterification

reaction. A similar result was obtained as shown in Figure 4.12 when a catalyst

concentration of 0.6% was used. From Figure 4.12, increasing the FFA content from

94

nil to about 2.4% (acid value of 4.84 mgKOH/g) resulted in an increase in conversion

time from less than 8 minutes to about 25 minutes in order to produce 80% FAME.

The conversion rate was further reduced when the FFA content was increased to

about 4.2 % (8.35 mgKOH/g), and there was no conversion as soon as the FFA

content reached 4.8% (9.53 mgKOH/g). The reduction in conversion rate due to FFA

content can also be seen at 1.0% catalyst concentration as shown in Figure 4.13.

In addition to the rate of conversion, the presence of FFA in the feedstock can

also have a negative impact on the FAME content in the reaction product produced

from the alkali-catalyzed process. The results in Figure 4.12 show that the FAME

content in the product decreased with the increasing percentage of FFA. An 88%

FAME was obtained from the use of original canola oil while an 81% product was

produced from the same oil with 2.4% FFA. The FAME content was reduced further

to 55-65% when the canola oil with 4.2% FFA was used. In Figure 4.14, the

appearances of the separation of the reaction mixtures are directly illustrated that the

presence of FFA in the feedstock make the separation process difficult due to the soap

formation.

95

Figure 4.11: Effect of free fatty acid content on the biodiesel conversion at 0.2 wt.%

NaOH (Methanol/Canola oil=9:1 (molar ratio); mixing speed=200 rpm;

T= 65°C)

0

10

20

30

40

50

60

70

80

90

100

0 0.5 1 1.5 2 2.5

Meth

yl

Est

ers

Co

nte

nt

(wt.

%)

Reaction Time (h)

Acid Value

Canola Oil

0.94mgKOH/g Canola Oil


96



T = 65°C)

0

10

20

30

40

50

60

70

80

90

100

0 0.5 1 1.5 2 2.5

Meth

yl

Est

ers

Co

nte

nt

(wt.

%)

Reaction Time (h)

Acid Value

Canola Oil





97



T =65°C)

0

10

20

30

40

50

60

70

80

90

100

0 0.5 1 1.5 2 2.5

Meth

yl

Est

ers

Co

nte

nt

(wt.

%)

Reaction Time (h)

Acid Vaule

Canola Oil



98

Figure 4.14: Photographs showing appearances of separation of reaction mixtures in

the separating funnel (Sample collected at reaction time=1 hour; NaOH

(wt.%)=0.6%; methanol/Canola oil=9:1 (molar ratio); mixing

speed=200 rpm; reaction temperature=65°C) (a) Acid value=4.84 mg

KOH/g (b) Acid value=8.35 mgKOH/g (c) Acid value=9.53 mgKOH/g

(a) (b) (c)

99

4.2.4 Determination of Reaction Rate Constant

Since the reaction system we applied in the alkali-catalyzed transesterification

is neither a pseudo-homogenous system nor a two-phase reaction system (i.e., the

acid-catalyzed reaction system applied in the previous study), the reaction rate

constant calculated is the observed value. The generalized transesterification reaction

is shown in the following Equation 2.26

Triglyceride (TG) + 3R′OH Glycerol (GL) + 3R′COOR3 [2.26]

Because the excess methanol is used to drive the reaction forward, the reverse

reaction is ignored. Then, the reaction rate can be given by Equation 4.2:

tALtTGobs

TG CCkdt

dCr

[4.2]

where kobs is the observed reaction rate constant, CTG-t is the concentration of

triglyceride at time t, and CAL-t is the concentration of alcohol at time t. Here, tALC

could be considered as a constant since the methanol concentration is larger than the

concentration of triglyceride. Then, Equation 4.2 can be derivated to Equation 4.3:

tTGobs

TG Ckdt

dCr ' [4.3]

where tALβ

obs'obs Ckk

In addition, the concentration of triglyceride at reaction time t can be expressed using

the conversion rate at reaction time t and initial concentration of triglyceride, as

shown in Equation 4.4:

CTG-t = CTG-0(1-x) [4.4]

100

where CTG-0 is the initial concentration of triglyceride and x is the triglyceride

conversion rate at time t. Thus, the reaction rate can be expressed using the initial

concentration of triglyceride and conversion rate as shown in Equation 4.5:

dt

dxC

dt

xCd

dt

dCr TG

TGtTG0

0 )]1([

[4.5]

If we assume reaction order α=1 and combine Equation 4.3 and 4.5, we get Equation

4.6 to calculate k'obs :

tkx

obs'1

1ln

[4.6]

As shown in Equation 4.6, k'obs is the slope of the graph

x1

1ln as a function of

reaction time t.

We have observed that the reaction conversion profiles at lower temperatures

(i.e., 25oC and 35

oC) appear to have two distinct conversion regions before the

reaction reaches the equilibrium: a slow conversion region and rapid conversion

region. At a low conversion region, the reaction kinetics is controlled by mass transfer

between the reactants. The duration time and the conversion rate in the slow

conversion region varied with temperature and catalyst concentration as shown in

Table 4.3. At a lower temperature (i.e., 25°C and 35°C), the k'obs is calculated by

using the experimental data from the second rapid conversion region since the

reaction kinetics is controlled by the chemical reaction in this region.

http://dict.cn/equilibrium

101

Table 4.3: Duration time and conversion rate for slow reaction region (200 rpm)

NaOH

(wt.%)

Conversion rate (%)

T=25°C T=35°C

30 min 60 min 30 min 60 min

0.2 1.26 1.92 2.72 4.05

0.6 8.13 19.65 13.96

1.0 11.03 16.68

102

By fitting

x1

1ln vs. reaction time, t, at different temperatures, a good liner

relationship between plots was satisfied and supports the hypothesis that the reaction

could be considered as first order (Figure 4.15). Table 4.4 gives the observed reaction

rate constants with respect to different temperatures and different catalyst

concentrations. The rate equation is expressed as follows:

tTGobsTG Ck

dt

dCr ' [4.7]

4.2.5 Demonstration of Reactor Design

Currently, the most common reactors used in biodiesel production in industry

are batch reactors and continuous reactors. In batch reactors, the oil is first charged to

the reactor, followed by the catalyst and methanol in the determined amount. The

reactor is then closed and controlled to operate under the desired reaction conditions.

After reaction is complete, the reacted mixture is removed from the reactor and sent

for purification processing. Batch reactors are better suited to smaller plants that do

not need 24/7 operation. The batch reaction process is showed in Figure 4.16.

Continuous stirred tank reactors (CSTRs) and plug flow reactors (PFRs) are

two types of continuous reactors applied in biodiesel production in industry, and they

are more efficient than batch reactors when large quantities of feedstocks are to be

processed. In CSTRs the reactants with a steady flow are continuously fed into the

reactors and the products are continuously withdrawn. Adequate mixing is required to

ensure that the concentration of any chemical involved should be approximately

constant anywhere in the reactor at all times. For PFRs, the reactants are fed into one

side of the reactor and travel in the axial direction of the reactor. Figure 4.17 is an

example of biodiesel production using a PFR process.

http://en.wikipedia.org/wiki/Continuous_stirred-tank_reactor



103

(a)

(b)

(c)

Figure 4.15: Plots of

x1

1ln vs. reaction time, t, (a) NaOH concentration =0.2 wt.%

(b) NaOH concentration =0.6 wt.% (c) NaOH concentration =1 wt.%

0.00

0.50

1.00

1.50

2.00

2.50

0 50 100 150 200 250 300

ln1

/(1

-x)

Reaction Time (min)


25℃

35℃

50℃

65℃

0.00

0.50

1.00

1.50

2.00

2.50

0 50 100 150

ln1

/(1

-x)

Reaction Time (min)


25℃

35℃

50℃

65℃

0

0.5

1

1.5

2

2.5

0 10 20 30 40 50 60

ln1

/(1

-x)

Reaction time (min)


25℃

35℃

50℃

65℃

104

Table 4.4: Observed reaction rate constant for alkali-catalyzed transesterification (200 rpm)

Temperature

(°C)

k'obs (min-1

)

NaOH

0.2 wt.% R2

NaOH

0.6 wt.% R2

NaOH

1 wt.% R2

25 0.0050 0.9982 0.0373 0.9409 0.0761 0.9913

35 0.0101 0.9688 0.0496 0.9978 0.0704 0.9008

50 0.0262 0.9743 0.0780 0.9718 0.3891 0.9839

65 0.0631 0.9989 0.1842 0.9777 0.4454 0.9998

(Note: k'obs for 25°C and 35°C

are for the second, fast conversion region; the information for the first, slow

conversion region is given in Table 4.3)

105

Alcohol

Catalyst

Crude Glycerol

Wash

Water

Water

Alcohol

Ester

Alcohol Water

Water

Dryer

Biodiesel

Batch Reactor

TG

Figure 4.16: Batch reaction process (Source:Van Gerpen, Shanks et al. 2004)

106

Figure 4.17: Plug flow reaction process (Source:Van Gerpen, Shanks et al. 2004)

Heater PFR1

Separator

PFR 2

Glycerol

Ester Alcohol

Alcohol

Glycerol

Ester

TG

Alcohol

Alcohol

TG

Catalyst

107

For a given duty, once the reactor has been selected, the proper volume of the

reactor is one of the main parameters for reactor design, and the starting point for reactor

design is based on the material balance equation, as showed in Equation 4.8:

volume of

element in reactant of

onaccumulati of rate

volume of

element the within

reaction chemcial

to due loss

reactant of rate

volume of

element of out flow

reactant of rate

volume of

element into flow

reactant of rate

[4.8]

Since the reaction kinetics for biodiesel reaction at lower temperatures (i.e., 25°C

and 35°C) are different from the reaction kinetics for biodiesel production at higher

temperatures (i.e., 50°C and 65°C), in the following sections, we discuss the volume of

the selected reactor for high-temperature design and low-temperature design separately.

For a batch reactor, since the composition is uniform throughout at any instant of

time, the material balance accounts for the whole reactor. Figure 4.18 shows simple

schematic of a batch reactor. There are no reactants entering or products leaving the

reaction mixture during the reaction so the input of reactant and output of product are

equal to zero, and evaluating the terms of Equation 4.8, we get:

volumeof

elementin reactant of

onaccumulati of

volumeof

element within the

reaction chemcial

todue loss

reactant of

00

rate

rate

[4.9]

dt

dnVr A

A 00 [4.10]

Here: dt

dxn

dt

dn AA

A0 [4.11]

108

Figure 4.18: Schematic of a batch reactor

FA

CA0

V0

CA

rA

xA

109

where t is the reaction time, rA is the reaction rate of material A, V is the volume of

reaction mixture, nA is the mole of A at time t, and nA0 is the mole of A at reaction time

zero.

Rearranging and integrating then gives:

[4.12]

where xA is the conversion rate of A at time t.

It is assumed that the density of the reaction mixture during the reaction remains

constant during the reaction and thereby obtain:

[4.13]

where CA0 is the initial concentration of A, CA is the concentration of A at reaction time t,

and xA is the conversion rate of A at time t.

For transesterification reactions at higher temperatures (i.e., 50°C and 65°C), the

reaction time can be determined using Equation 4.13. However, for reactions at lower

temperatures, the reaction time for transesterification is the reaction time for the slow

reaction region, ts, and the reaction time for fast reaction region, tf.

fs ttt [4.14]

Since the rate equation at the slow reaction regime is complicated and difficult to

determine, we can use empirical data from experiments to determine the required reaction

time, ts and the conversion rate, xs, as summarized in Table 4.5. The reaction time for the

fast region is determined using following Equation 4.15, and the conversion rate is xs

instead of x0:

[4.15]

110

Table 4.5: Experimental data for slow reaction region (200 rpm)

NaOH

(wt.%)

T=25°C T=35°C

ts

min

xs

%

ts

min

xs

%

0.2 60 1.92 60 4.05

0.6 60 19.65 30 13.96

1 30 11.03 30 16.68

Cs= CA0(1-xs)

111

where Cs is the initial concentration of A in the fast reaction region, CA is the

concentration of A at reaction time t, and xs is the initial conversion rate of A in the fast

reaction region. The rate equation obeys Equation 4.7.

The volume of the batch reactor can be calculated using Equation 4.16:

0

0

A

AR

C

tFV [4.16]

where FA0 is the processing flow of A per unit time and VR is the reactor volume.

In a plug flow reactor, the composition of the reaction mixture changes from place

to place, and the material balance for component A must be considered for a different

element of volume dV, as shown in Figure 4.19. The material balance for A becomes:

0

volume of

element the within

reaction chemcial

to due loss

reactant of rate

volume of

element of out flow

reactant of rate

volume of

element into flow

reactant of rate [4.17]

RAAAA dVrdF FF )()( [4.18]

where AAAAA dxFx-(1 FddF 00 ) [4.19]

After rearranging, we obtain:

RAAA dVrdxF )(0 [4.20]

Thus:

Afx

A

AAR

r

dxFV

00 [4.21]

When biodiesel production is conducted using a plug flow reactor process at

higher temperatures, the reactor size for a given flow rate of 0AF and required conversion

of xAf is

112

CA0

FA0

CA0

xA0=0

FA

xA

FA+dFA

xA+dxA

CAf

FAf

xAf

dV

Figure 4.19: Schematic of a plug flow reactor

113

determined by Equation 4.21. At lower temperatures, the size of the reactor required for

the slow reaction region and the size for the high reaction region must each be

determined and, then, added together to get the total size of the reactor required, as

described by Equation 4.22. Since the reaction progress is extremely slow in the slow

reaction region, the concentration and conversion of A are the same at any position of the

plug flow reactor and in CSTR. Then, the volume is determined by Equation 4.23, and

for fast reaction region, the volume is calculated using Equation 4.24.

fsR VVV [4.22]

where 0

0

A

As

C

tFV [4.23]

Af

s

x

xA

AsAf

r

dxxFV )1(0 [4.24]

In CSTRs, the concentration of any chemical and reaction rate are constant

anywhere in the reactor at all times, as shown in Figure 4.20. The material balance is as

shown in Equations 4.25 and 4.26

0

volume of

element the within

reaction chemcial

to due loss

reactant of rate

volume of

element of out flow

reactant of rate

volume of

element into flow

reactant of rate [4.25]

dVr x-(1FF AfAfAA )()00 [4.26]

where rAf is the reaction rate at exit of the reactor

114

FA0

CA0

V0

FA

CA

rAf

xAf

CA

rAf

xAf

Figure 4.20: Schematic of a continuous stirred tank reactor

115

After rearranging Equation 4.26, the volume of the reactor for CSTRs can be

determined by using Equation 4.27.

fA

AfA

Rr

xFV

)(

0 [4.27]

Equation 4.27 is used for determining the volume of the CSTRs for biodiesel production

at higher temperatures, and the volume of the CSTRs for biodiesel reaction at lower

temperatures is determined with Equation 4.28 where the volumes for the slow reaction

region and fast reaction region are determined with Equation 4.23 and Equation 4.29,

respectively.

fsR VVV [4.28]

where fA

AfsA

fr

xxFV

)(

)1(0 [4.29]

116

Table 4.6: Summary of reactor design at different temperatures (200 rmp)

Reactor Type

Reactor Volume at Different Temperatures

25°C and 35°C 50°C and 65°C

Batch Reactor 0

0

A

sAR

C

tFV

0

0 )(

A

fsA

RC

ttFV

PFR

Af

s

x

xA

AsA

A

AR

r

dxxF

C

tFV )1(0

0

0

Afx

A

AAR

r

dxFV

00

CSTR fA

AfsA

A

AR

r

xxF

C

tFV

)(

)1(0

0

0

fA

AfA

Rr

xFV

)(

0

Note: 1) For values of ts, xs, and Cs, refer to Table 4.5, 2) )1(' 0 xCkr AobsA ,

)1(')( 0 fAobsfA xCkr , 3) obsk' values refer to Table 4.4.

117

Chapter 5 Conclusions and Recommendations

5.1 Conclusions

This thesis studied biodiesel production by using simulated low quality feedstocks

(i.e., mixtures of the canola oil and oleic acid/linoleic acid). Since the high content of

FFA in the low quality feedstock will greatly reduce the biodiesel production rate in an

alkali-catalyzed transesterification process, esterification was used to effectively decrease

the FFA content prior to the alkali-catalyzed transesterification. The previous studies on

esterification were conducted in an heterogeneous reaction system, in which the

hydrodynamic effect and the chemical reaction control the production efficiency and

reaction rate. A new kinetic model of esterification was developed in an immiscible two-

phase reaction system, in which the hydrodynamic effect was completely eliminated

under an appropriate agitation speed in the oil phase. Thus a real, kinetically controlled

reaction system was achieved. Based on the new reaction system, a number of

experiments were carried out to determine the reaction rate constant and activation

energy. The parametric effects on the reaction rate were examined and discussed. In

addition, study of the parametric effects on alkali-catalyzed transesterification was also

successfully carried out through a series of experiments.

This research covered three aspects of biodiesel production using a simulated low

quality feedstock. The following are the conclusions drawn from this study:

1) A new kinetic reaction system was developed for esterification in an immiscible

reaction system:

118

The mass transfer resistance was found to be negligible in the methanol

phase.

The mass transfer resistance of FFA in the oil phase had a great impact on

the overall reaction rate and was negligible in the particular range of

agitation speeds.

Based on the above findings and the following conditions and assumptions,

the esterfication reaction was only controlled over by the pure chemical

reaction, and the hydrodynamic effect was completely removed, the

reaction rate can be simply determined by the equation FFAsRX Ckr ':

(1) The oil phase is gentle agitated to keep the interface undisturbed.

(agitation speed controlled at 80 rpm in this study)

(2) Assume that all reacted FFA are converted to biodiesel.

(3) Assume the concentration of methanol remains constant during the

esterification reaction because it is pure and in an excessive amount.

2) The rate constant and activation energy of esterification were determined and the

parametric effects on the reaction rate were discussed:

The reaction was found to proceed in the first order reaction as a function

of the FFA content.

An increase of temperature leads to an increased reaction rate constant in a

proportional manner, resulting in an increase in reaction rate.

119

An increase of catalyst concentration caused the reaction rate to increase

at temperatures of 50°C and 62°C. At a temperature of 35°C, the change

of reaction rate with the catalyst concentration is negligible.

An increase in the initial FFA content leads to an increased reaction rate

until the reaction reaches equilibrium. The initial FFA content has no

impact on the reaction rate constant.

The esterification reactions of oleic acid and linoleic acid have very

similar activation energies. The activation energies of esterification of

oleic acid were 32.48, 31.96, and 39.14 kJ/mol at concentrations of H2SO4

1 wt.%, 2 wt %, and 3 wt.%, respectively. The activation energies of

esterification of linoleic acid were 31.58 and 34.23 kJ/mol at

concentrations of H2SO4 1 wt.% and 3 wt.%, respectively.

3) The parametric effects on the alkali-catalyzed transesterification reaction were

studied:

Raising the reaction temperature increased the biodiesel conversion rate.

The increasing conversion rate with temperature is caused by an increase

in the transesterification reaction rate and a reduction in viscosity of

feedstock oil, which promotes ultimate mixing between oil and methanol

phases.

The maximum FAME content obtained ranged from 80 to 90 wt.%

regardless of the catalyst concentration and temperature.

120

The reaction conversion rate increased as the concentration of catalyst

increased.

The presence of FFA caused the conversion rate to drop and made the

separation process difficult because of the soap formation.

4) The observed reaction rate constants at different temperatures were determined

and the reactor design for a given duty was summarized. Since the reaction

kinetics are different at low temperatures (25°C and 35°C) and high temperatures

(50°C and 65°C), the required reactor volume for a given duty must be

determined based the temperature.

5.2 Recommendations for Future Work

This work proposed a new reaction system for esterification of biodiesel

production. The kinetic data obtained were intrinsic, and they can be used in industrial

design for resizing or optimizing the reactor. The following are our recommendations for

future work:

1) The reaction system is based on the condition that the interfacial surface area is

fixed and equal to the area of undisturbed two-phase interface. Visual observation

was used to make the judgment on the interface change. Future work may use

other advanced technologies in order to precisely indentify the interface change

and control it undisturbed.

2) Hydrodynamics is another important factor which affects the reaction rate of

heterogeneous reaction of biodiesel. The study of hydrodynamic effect had been

conducted in another research project separately. (Nath, D. 2012)

121

3) This study was conducted using a simulated low quality feedstock for studying

the reaction kinetics. Therefore, real low quality feedstocks such as waste cooking

oil containing a high amount of impurities and different types of FFA should be

investigated.

4) The alkali-catalyst transesterification system is also an immiscible two-phase

system, so the new kinetic model can also be applied in the system under

particular conditions for precisely evaluating its reaction kinetics.

5) Since the esterification is a pre-treatment process of FFA prior to the

transesterification process, a continuous process study including the esterification

and transesterification is necessary for the commercial production of biodiesel

using low quality feedstock in the future.

122

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