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
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
I
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 35oC) and high temperatures (50°C and 65°C),
which resulted in different designs for reactor volume for a given duty based on different
temperatures.
II
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.
III
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
IV
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
V
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
VI
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 3.4: Reaction rate constants at 2 wt.% H2SO4..................................................... 57
Table 3.5: Reaction rate constants at 1 wt.% H2SO4..................................................... 58
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
VII
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
VIII
Figure 3.9: Graph of ln
CFFA0
as a function of time (H2SO4 concentration=3
CFFA
wt.%) (a) T=35°C (b) T=50°C (c) T=62°C ................................................... 53
Figure 3.10: Graph of ln
CFFA0
as a function of time (H2SO4 concentration=2
CFFA
wt.%) (a)T=35°C (b) T=50°C (c) T = 62°C ................................................... 54
Figure 3.11: Graph of ln
CFFA0
as a function of time (H2SO4 concentration=1
CFFA
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
IX
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 4.5: Effect of temperature on the conversion profile at 0.6 wt.% NaOH
(Methanol/Canola oil=9:1(molar ratio); mixing speed=200 rpm) ................. 85
Figure 4.6: Effect of temperature on the conversion profile at 1.0 wt.% NaOH
(Methanol/Canola oil=9:1(molar ratio); mixing speed=200 rpm) ................. 86
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 4.8: Effect of catalyst concentration on the conversion profile at 35°C
(Methanol/Canola oil=9:1 (molar ratio); mixing speed=200 rpm) ................ 90
Figure 4.9: Effect of catalyst concentration on the conversion profile at 50°C
(Methanol/Canola oil=9:1 (molar ratio); mixing speed=200 rpm) ................ 91
Figure 4.10: Effect of catalyst concentration on the conversion profile at 65°C
(Methanol/Canola oil=9:1 (molar ratio); mixing speed=200 rpm) ................ 92
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
X
Figure 4.12: Effect of free fatty acid content on the biodiesel conversion at 0.6
wt.% NaOH (Methanol/Canola oil=9:1 (molar ratio); mixing
speed=200 rpm; T=65°C) ............................................................................ 96
Figure 4.13: Effect of free fatty acid content on the biodiesel conversion at 1.0
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
1
Figure 4.15: Plots of ln(
) vs. reaction time, t, (a) NaOH concentration =0.2
1 x
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
XI
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
XII
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)
R1 - COOH
FFA
R2  OH
alcohol
R1  COOR2
biodiesel
S
interfacial surface (m2)
SnCl2
tin chloride
SOx
sulfur oxides
t
time (s)
T
temperature ( K)
VRX
reaction volume (m3)
Greek letters
XIII
α
the reaction order with respect to FFA
β
reaction order with respect to alcohol
XIV
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).
1
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;
2
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)
3
(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
Catalyst
CH1-OOC-R2 +
3R'OH
R2-COO-R'
R3-COO-R'
CH2-OOC-R3
Glyeride
Esters
Alcohol
CH2-OH
R1-COO-R'
+
CH1-OH
CH2-OH
[1.1]
Glycerol
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,
4
Table 1.2: FFAs content in various biodiesel feedstocks
Feedstock
FFAs
Refined vegetable oils
< 0.05 %
0.3 – 0.7%
Crude vegetable oil
Restaurant waste grease
2 – 7%
Animal fat
5 – 30%
Trap grease
40 – 100%
Source: (Van Gerpen, Shanks et al. 2004)
5
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
6
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.
R1 - COOH R 2  OH  R1  COOR2  H2O
FFAs
alcohol
biodiesel
[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
7
Acid Catalyst
Method І:
Alkaline Catalyst
Esterification
Transesterification
Purification
Acid Catalyst
Method П:
Esterification
+
Transesterification
Purification
Figure 1.1: Simplified scheme of two biodiesel production methods from low quality
feedstocks. Source: (Van Gerpen, Shanks et al. 2004)
8
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
9
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.
10
70
Sendzikiene, Makareviciene et al.
(2004)
60
Activation Energy kJ/mol
Berrios, Siles et al. (2007)
50
Aranda, Santos et al. (2008)
40
Supardan (2008)
30
Thiruvengadaravi, Nandagopal et al.
(2009)
20
Praveen K.S Yadav.et al (2010)
10
0
0
2
4
6
8
10
12
H2 SO4 wt.%
Figure 1.2: Acitavation energy obtained in different studies at different catalyst
concentrations
11
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.
12
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.
13
Direction of mass transfer
Interface
Concentration of solute A
Reaction zone
yA,G
CA,i
Gas film
Liquid film
yA,i
CA,L
Distance from interface
Figure 2.1: Mass transfer process based on the two-film theory (Redrawn from Astaria,
Savage et al. (1983))
14
(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:
Catalyst
Alcohol + FFA
Ester + H2O
[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:
*
 rFFA  kFFAOa(CFFAO  CFFA
i )
[2.2]
where rFFA= the mass transfer rate of FFA to the interface; kFFAO = the mass transfer
coefficient of FFA in the oil film; CFFAO = the concentration of FFA in the oil bulk;
*
CFFA
i = the concentration of FFA at the interface in the oil film; a = the interfacial area
in per unit volume.
15
Concentration of FFA
Interface
Oil Film
Methanol Film
CFFA-O
C*FFA-i
Oil bulk
P*FFA-i
Methanol
bulk
PFFA-AL
Direction of massbtransfer
Figure 2.2: Mass transfer process
of FFA from the oil phase to the methanol phase
b
16
In the methanol film:
*
 rFFA  kFFA ALa(PFFA
i  PFFA AL )
where
[2.3]
*
kFFA AL = the mass transfer coefficient of FFA in the methanol film; PFFA
i = the
concentration of FFA at the interface in the methanol film;
PFFA AL = 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:
*
FFFA  rFFAVRX  kFFAOa(CFFAO  CFFA
i )VRX
[2.4]
and the equation for the overall reaction rate at the interface is:

 rRX  kRXC*
FFAi
C  ALi
[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. CALi = 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. C  ALi becomes a constant and can combine with kRX. Therefore,
the overall reaction rate can be expressed with respect to the concentration of FFA:
*
 rRX  k 'RX CFFA
i
where k'RX = the pseudo-rate constant.
17

[2.6]
The flow rate of FFA at the interface is:

*
FFFA  rFFAVRX  k 'RX CFFA
i VRX
[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:
*
*
CFFAO  (CFFAO  CFFA
i ) Eq2.4  (CFFAi ) Eq2.7 =
FFFA
FFFA
F
1
1

 FFA (

)
k FFAO aVRX k ' RX VRX VRX k FFAO a k ' RX
[2.8]
The overall reaction rate of the esterification reaction is obtained by rearranging Equation
2.8:


FFFA
1
 r  
 1
1
VRX

k
 FFAO a k 'RX


C
 FFAO


[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
18
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]
19
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.
20
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
21
[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.
22
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
23
Table 2.1: Literatures on kinetic study of the acid-catalyzed esterification reaction using homogenous catalysts
References
Yadav, Singh et al.
(2010)
Thiruvengadaravi,
Nandagopal et al.
(2009)
Cardoso, Neves et al.
(2008)
Research Objectives



Optimized the pretreatment process.
Undertook kinetic and
thermodynamic studies of
esterification

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


Aranda, Santos et al.
(2008)

Yalçinyuva, Deligöz
et al. (2008)

Supardan (2008)
Optimized the conditions for
production of palm fatty acid methyl
esters

Studied the esterification of palm
fatty acids, by-products of edible
palm Oil production, to produce
biodiesel, using homogeneous acid
catalysts.
Studied the esterification kinetics of
myristic acid with isopropyl alcohol
with both homogeneously and
heterogeneously catalyzed systems
Studied the effect of operational
variables on the esterification of
FFA in low grade CPO;
Studied the influence of operational
variables on the kinetics.
Test Conditions
Important Findings




Reactants: Palm fatty acid(FFA=93 wt.%)
Alcohol: methanol
Catalyst: H2SO4
500 mL three neck flask, stirrer





FFA: FFA in Pongamia
Alcohol: methanol
Catalyst: H2SO4
Bath reactor, mechanical stirrer, speed
(N/A)




FFA: oleic acid in soybean oil
Alcohol: ethanol
Catalyst: H2SO4, tin chloride (SnCl2)
50 mL three-necked glass flask, magnetic
stirrer speed N/A
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














FFA: myristic acid
Alcohol: isopropyl alcohol
Catalyst: ρ-toluene sulfonic acid, amberlyst15 and Degussa (acidic cation exchange
resin)
250 mL round bottomed reactor, magnetic
stirrer, mixing speed=450 rpm
FFA: low grade CPO with FFA content of
5.6% and 33.3%.
Alcohol: methanol
Catalyst: H2SO4
Mechanical agitation =464 rpm
24












Pseudo first-order kinetics for
esterification.
Et=15.31 kJ mol-1
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.%,
SnCl2 is a potential catalyst for the low
quality raw materials.
A first order dependence for both
esterification reaction catalyzed by
H2SO4 and SnCl2.
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.%.
Second-order kinetics for the
homogeneous catalyst.
No pore diffusion when using
heterogeneously catalyst limitation.
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
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






Kocsisová, Cvengroš et
al. (2005)

Sendzikiene,
Makareviciene et al.
(2004)

Studied the reaction of ester
preparation in short reaction time
with small excess of alcohol, low
catalyst, but higher conversion to
esters



FFA: Fatty acids in sunflower oil
Alcohol: methanol
Catalyst: H2SO4
Bath reactor, magnetic agitation speed 600
rpm
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)





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)
25


A first-order kinetic law for the
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.%
A first-order kinetic law for the
reactions.
The reaction rate is two to three
times higher than at the
temperatures close to the boiling
point of MeOH
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
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.
26
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%
27
-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
28
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
29
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 byproducts 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.
30
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 alkalicatalyzed transesterification process converted the products of the first step to monoesters 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 acidcatalyzed 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
31
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.
32
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;
33
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%
34
Mechanical Stirrer
Condenser
Temperature
Indicator
Thermostatic
Water Bath
Impeller
Water Jacket
Figure 3.1: Schematic diagram of experimental setup for acid-catalyzed esterification
reaction
35
Reactor
Water bath
Figure 3.2: Photographs of the esterification experimental setup (Original in color)
36
(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
37
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 acidcatalyzed 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:
 mg KOH  volume KOH (mL)  N KOH (mmol/mL) 56.1(mg/mmol)
  
Acid value

sample weight (g)
 g sample  

FFAconversion(%) 
Ai  At
100
Ai
[3.1]
[3.2]
where Ai= the initial acid value; and At= the acid value at a certain reaction time.
38
Methanol/H2SO4
Reaction
mixture
Simulated low quality oil
Esterification
reaction
Reaction
mixture
Quench
Dried
biodiesel
Centrifuge
Reaction
Figure 3.3: Experimental procedure for the acid-catalyzed esterification
39
Titration
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
40
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
41
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
42
100
FFA Conversion (%)
80
no agitation at two phase
agitate oil phase 50 rpm
60
agitate methanol phase at 50 rpm
40
20
0
0
10
20
30
40
50
60
Reaction Time (min)
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
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.
44
Low
Agitation Speed
High
Figure 3.5: Change of the interface state with increasing agitation speeds
45
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;
46
100
III
FFA Covnersion (%)
80
II
I
60
Reaction Time
20 min
40
40 min
60 min
20
0
0
20
40
60
80
100
120
140
160
180
200
220
Agitation Speed (rpm)
Figure 3.6: Effect of agitation speed on the FFA conversion rate (T=50°C, H2SO4
concentration=3wt.%, FFA content=37 mgKOH/g)
47
100
II
FFA Conversion (%)
80
60
Reaction Time
20 min
40
40 min
60 min
20
0
0
20
40
60
80
100
120
Agitation Speed (rpm)
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
100
FFA Conversion (%)
80
III
Reaction Time
60
0 min
20 min
40
40 min
60 min
20
0
0
20
40
60
80
100
120
140
Agitation Speed (rpm)
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
(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:
FFFA
VRX


1
 r  
 1
1


 k FFAO a k ' RX


C
 FFAO


[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:
 r  k' RX CFFA
50
[3.3]
(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 ln
ln
CFFA0
against time.
CFFA
CFFA0
 k ' RX t
CFFA
[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 k'RX value is determined at different temperatures, the activation
energy for the esterification can be estimated by using the Arrhenius equation:
Ea
 RT
k 'RX  Ae
[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.
51
According to Equation 3.4, the relationship between ln
CFFA0
and reaction time (t)
CFFA
is linear. The value of the rate constant (k´RX) equals the slope of the linear regression
trendline. Thus, ln
CFFA0
is plotted against t in different experimental conditions. As
CFFA
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:
lnk 'RX  lnA 
Ea
RT
[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
52
3
2.5
ln(CFFA0 /CFFA)
2
1.5
Initial FFA Content
35.47 mgKOH/g
1
13.32 mgKOH/g
0.5
5.90 mgKOH/g
0
0
60
120
180
240
300
360
Reaction Time (min)
(a)
3
ln(C FFA0 /C FFA)
2.5
2
Initial FFA Content
1.5
36.39 mgKOH/g
1
14.14 mgKOH/g
0.5
4.86 mgKOH/g
0
0
60
120
180
240
Reaction Time (min)
(b)
3
ln(C FFA0 /C FFA)
2.5
2
Initial FFA Content
1.5
35.62 mgKOH/g
1
15.30 mgKOH/g
4.74 mgKOH/g
0.5
0
0
60
120
180
240
Reaction Time (min)
(c)
Figure 3.9: Graph of ln
CFFA0
as a function of time (H2SO4 concentration =3 wt.%)
CFFA
(a)T=35°C (b) T=50°C (c) T=62°C
53
3
ln(CFFA0 /CFFA)
2.5
2
Intial FFA Content
1.5
37.33 mgKOH/g
1
5.20 mgKOH/g
0.5
0
0
60
120
180
240
300
360
Reaction Time (min)
(a)
3
ln(CFFA0 /CFFA)
2.5
2
1.5
Intial FFA Content
1
37.33 mgKOH/g
0.5
4.72 mgKOH/g
0
0
60
120
180
240
300
Reaction Time (min)
(b)
3
2.5
2
ln CFFAO/CFFA
Intial FFA Content
1.5
37.33 mgKOH/g
1
4.72 mgKOH/g
0.5
0
0
60
120
180
240
300
Reaction Time (min)
(c)
Figure 3.10: Graph of ln
CFFA0
as a function of time (H2SO4 concentration =2 wt.%)
CFFA
(a)T=35°C (b) T=50°C (c) T=62°C
54
3
ln (C FFA0 /C FFA)
2.5
2
Initial FFA Content
1.5
35.87 mgKOH/g
1
14.60 mgKOH/g
0.5
5.27 mgKOH/g
0
0
60
120
180
240
300
360
420
Reaction Time (min)
(a)
3
2.5
ln(CFFA0/CFFA)
2
Initial FFA Content
1.5
36.09 mgKOH/g
1
4.74 mgKOH/g
0.5
0
0
60
120
180
240
300
360
Reaction Time (min)
(b)
3
ln(C FFA0 /C FFA)
2.5
2
Initial FFA Content
1.5
36.09 mgKOH/g
1
4.74 mgKOH/g
0.5
0
0
60
120
180
240
300
Reaction Time (min)
(c)
CFFA0
Figure 3.11: Graph of ln
as a function of time (H2SO4 concentration =1 wt.%)
CFFA
(a)T=35°C (b) T=50°C (c) T=62°C
55
Table 3.3: Reaction rate constants at 3 wt.% H2SO4
Initial FFA content
Temperature
Reaction rate constant
(mgKOH/g)
(°C)
k´RX (min-1) ×102
Correlation
coefficients
(R2)
0.53
0.9376
0.53
0.9910
5.90
0.47
0.9559
36.39
1.07
0.9583
1.12
0.9970
4.86
0.95
0.9920
35.62
1.63
0.9801
1.83
0.9910
1.77
0.9862
35.47
13.32
14.14
15.30
35
50
62
4.74
56
Table 3.4: Reaction rate constants at 2 wt.% H2SO4
Initial FFA content
Temperature
Reaction rate constant
(mgKOH/g)
(°C)
k´RX (min-1) ×102
Correlation
coefficients
(R2)
0.59
0.9855
0.54
0.9968
0.97
0.9889
0.93
0.9941
1.48
0.9885
1.61
0.9876
37.33
35
5.20
37.33
50
4.72
37.33
62
4.72
57
Table 3.5: Reaction rate constants at 1 wt.% H2SO4
Initial FFA content
Temperature
Reaction rate constant
(mgKOH/g)
(°C)
k´RX (min-1) ×102
Correlation
coefficients
(R2)
0.49
0.9537
14.6
0.50
0.9384
5.90
0.37
0.9884
36.09
0.76
0.9819
4.86
0.80
0.9914
36.09
1.25
0.9836
1.29
0.9987
35.47
35
50
62
4.74
58
0
-1
H 2 SO 4 Concentration
-2
lnk'RX
1wt.%
-3
2wt.%
3wt.%
-4
-5
-6
0.00295
0.003
0.00305 0.0031
0.00315 0.0032
0.00325 0.0033
1/T
Figure 3.12: Arrhenius plot of lnk'RX against 1/T (Esterification of oleic acid)
59
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
60
Table 3.6: Activation energy in the esterification reaction of oleic acid
H2SO4 concentration
Ea
Correlation coefficients
A
(R2)
(wt.%)
(kJ/mol)
1
32.48
1436.55
0.9961
2
31.96
1450.99
0.9956
3
39.14
22136.87
0.9999
61
100
FFA Conversion (%)
80
Temperature
60
35℃
40
50℃
62℃
20
0
0
200
400
600
800
Reaction Time (min)
(a)
FFA Conversion (%)
100
80
Temperature
60
35℃
40
50℃
62℃
20
0
0
100
200
300
400
500
Reaction Time (min)
(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
62
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
63
2.00
1.80
k'RX( min-1 )×10 2
1.60
1.40
1.20
1.00
Initial FFA Content
0.80
35-38 mg KOH/g
0.60
13-16 mg KOH/g
0.40
0.20
0.00
0
10
20
30
40
50
60
70
Reaction Temperature (℃)
Figure 3.14: Effect of temperature on the reaction rate constant (H2SO4 concentration=3
wt.%)
64
100
FFA Conversion (%)
80
60
H2SO 4 Concentration
40
1 wt.%
2 wt.%
20
3 wt.%
0
0
60
120
180
240
300
360
Reaction Time (min)
(a)
FFA Conversion (%)
100
80
60
H2SO 4 Concentration
40
1 wt.%
2 wt.%
20
3 wt.%
0
0
60
120
180
240
300
360
Reaction Time (min)
(b)
100
FFA Conversion (%)
80
60
H2SO4 Concentration
1 wt.%
40
2 wt.%
20
3 wt.%
0
0
60
120
180
240
300
Reaction Time (min)
(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
65
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
66
1.80
1.60
k' RX ( min-1 )×10 2
1.40
1.20
Temperautre
1.00
35℃
0.80
50℃
0.60
62℃
0.40
0.20
0.00
0
1
2
3
4
H2 SO4 Concentration (wt.%)
Figure 3.16: Effect of catalyst concentration on reaction rate constant (Initial FFA
content =35-38 mgKOH/g)
67
FFA Content (mgKOH/g)
40
Initial FFA Content
4.86 mgKOH/g
14.14 mgKOH/g
20
36.39 mgKOH/g
0
0
60
120
Reaction Time (min)
(a)
FFA Content (mgKOH/g)
40
Initial FFA Content
4.74 mgKOH/g
20
15.30 mgKOH/g
35.62 mgKOH/g
0
0
60
120
Reaction Time (min)
(b)
Figure 3.17: Change of FFA content as a function of reaction time
(a) T=50°C (b) T=62°C
68
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
69
2.00
1.80
k'RX ( min-1 )×10 2
1.60
1.40
1.20
1.00
Temaperture
0.80
50℃
0.60
62℃
0.40
0.20
0.00
0
5
10
15
20
25
30
35
40
Initial FFA Content (mgKOH/g)
Figure 3.18: Effect of the initial FFA content on the reaction rate constant
70
.
0
-1
H 2 SO 4 Concentration
lnk' RX
-2
1wt.%
-3
3wt.%
-4
-5
-6
0.00295
0.003
0.00305
0.0031
0.00315
0.0032
0.00325
0.0033
1/T
Figure 3.19: Arrhenius plot of lnk΄RX against 1/T (Esterification of linoleic acid)
71
Table 3.7: Activation energy of the esterification reaction using linoleic acid
H2SO4 concentration
Ea
Correlation coefficients
A
(R2)
(wt.%)
(kJ/mol)
1
31.58
1344.261
0.9584
3
34.23
4072.857
0.9770
72
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.
73
2.5
k' RX (min-1 ) × 10 2
2
1.5
1
0.5
0
0:1
1/4:3/4
1/2:1/2
3/4:1/4
1:0
Oleic acid: Linoleic acid
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
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,
75
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%
76
Cold water out
Condenser
Cold water in
Thermometer
Sampling point
Three necked flask
Stirring hot plates
Figure 4.1: Schematic diagram of the alkali-catalyzed transesterification experimental
setup
77
Reactors connected with
condensers
Water bath for preheating
methanol
Separating funnels
Figure 4.2: Photographs of the alkali-catalyzed transesterification experimental setup
(Original in color)
78
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 (ColeParmer 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
79
crude biodiesel in the top layer was collected and gently washed with 10 mL of deionized 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 (splitlessmode). 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:
80
Water
Methanol/NaOH
Canola oil
Transesterification
Reaction
Reaction
mixture
GC/MS
Phase
Phase
Separation
Separation
Crude
biodiesel
Water
wash
Washed
biodiesel
Drying
Na2SO4
Dried
biodiesel
Titration
Glycerol (Disposal)
Waste Water
Figure 4.3: Experimental procedure for the alkali-catalyzed transesterification
81
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
82
 ( A  AR )  CRVR
wt.%    i

AR

 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
83
Methyl Esters Content (wt.%)
100
90
80
70
60
Reaction Temperature
50
25℃
40
35℃
30
50℃
20
60℃
10
0
0
1
2
3
4
5
6
7
Reaction Time (h)
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
Methyl Esters Content (wt.%)
100
90
80
70
60
Reaction Temperature
50
25℃
40
35℃
30
50℃
20
65℃
10
0
0
1
2
3
4
5
6
7
Reaction Time (h)
Figure 4.5: Effect of temperature on the conversion profile at 0.6 wt.% NaOH
(Methanol/Canola oil=9:1(molar ratio); mixing speed=200 rpm)
85
Methyl Esters Content (wt.%)
100
90
80
70
Reaction Temperature
60
50
25℃
40
35℃
30
50℃
65℃
20
10
0
0
1
2
3
4
5
Reaction Time (h)
Figure 4.6: Effect of temperature on the conversion profile at 1.0 wt.% NaOH
(Methanol/Canola oil=9:1(molar ratio); mixing speed=200 rpm)
86
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 mm2/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
87
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.
88
Methyl Esters Content (wt.%)
100
90
80
70
60
50
NaOH Concentration
40
30
0.2 wt.%
20
0.6 wt.%
10
1.0 wt.%
0
0
1
2
3
4
5
6
Reaction Time (h)
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
100
Methyl Esters Content (wt.%)
90
80
70
60
50
NaOH Concentration
40
0.2 wt.%
30
0.6 wt.%
20
1.0 wt.%
10
0
0
1
2
3
4
Reaction Time (h)
Figure 4.8: Effect of catalyst concentration on the conversion profile at 35°C
(Methanol/Canola oil=9:1 (molar ratio); mixing speed=200 rpm)
90
100
Methyl Esters Content (wt.%)
90
80
70
60
50
NaOH Concentration
40
0.2 wt.%
30
0.6 wt.%
20
1.0 wt.%
10
0
0
0.5
1
1.5
2
Reaction Time (h)
Figure 4.9: Effect of catalyst concentration on the conversion profile at 50°C
(Methanol/Canola oil=9:1 (molar ratio); mixing speed=200 rpm)
91
100
Methyl Esters Content (wt.%)
90
80
70
60
50
Catalyst Concentration
40
0.2 wt.%
30
0.6 wt.%
20
1.0 wt.%
10
0
0
0.5
1
1.5
2
Reaction Time (h)
Figure 4.10: Effect of catalyst concentration on the conversion profile at 65°C
(Methanol/Canola oil=9:1 (molar ratio); mixing speed=200 rpm)
92
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
93
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.
94
Methyl Esters Content (wt.%)
100
90
80
70
60
50
Acid Value
40
Canola Oil
30
0.94mgKOH/g Canola Oil
20
4.84mgKOH/g Canola Oil
10
0
0
0.5
1
1.5
2
2.5
Reaction Time (h)
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
100
Methyl Esters Content (wt.%)
90
80
70
60
Acid Value
50
Canola Oil
40
4.84mgKOH/g Canola Oil
30
6.171mgKOH/g Canola Oil
20
8.35mgKOH/g Canola Oil
10
9.53mgKOH/g Canola Oil
0
0
0.5
1
1.5
2
2.5
Reaction Time (h)
Figure 4.12: Effect of free fatty acid content on the biodiesel conversion at 0.6 wt.%
NaOH (Methanol/Canola oil=9:1 (molar ratio); mixing speed=200 rpm;
T = 65°C)
96
Methyl Esters Content (wt.%)
100
90
80
70
60
Acid Vaule
50
40
Canola Oil
30
0.94mgKOH/g Canola Oil
20
4.84mgKOH/g Canola Oil
10
0
0
0.5
1
1.5
2
2.5
Reaction Time (h)
Figure 4.13: Effect of free fatty acid content on the biodiesel conversion at 1.0 wt.%
NaOH (Methanol/Canola oil=9:1 (molar ratio); mixing speed=200 rpm;
T =65°C)
97
(a)
(b)
(c)
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
98
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:
r 
dCTG
 kobsC TGt C  ALt
dt
[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, C  ALt
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:
r 
where
dCTG
 k 'obs C TGt
dt
[4.3]
'
kobs
 kobsC β ALt
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)
99
[4.4]
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:
r 
dCTGt d[CTG0 (1 x)]
dx

 CTG0
dt
dt
dt
[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 :
 1 
ln
  k 'obs t
1  x 
[4.6]
1 
As shown in Equation 4.6, k'obs is the slope of the graph ln 
 as a function of
1 x 
reaction time t.
We have observed that the reaction conversion profiles at lower temperatures
(i.e., 25oC and 35oC) 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.
100
Table 4.3: Duration time and conversion rate for slow reaction region (200 rpm)
Conversion rate (%)
NaOH
(wt.%)
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
101
1 
By fitting ln 
 vs. reaction time, t, at different temperatures, a good liner
1 x 
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:
r 
dCTG
 k 'obs CTGt
dt
[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.
102
2.50
ln1/(1-x)
2.00
Reaction Temperature
1.50
25℃
35℃
1.00
50℃
0.50
65℃
0.00
0
50
100
150
200
250
300
Reaction Time (min)
(a)
2.50
ln1/(1-x)
2.00
1.50
Reaction Temperature
25℃
1.00
35℃
50℃
0.50
65℃
0.00
0
50
100
Reaction Time (min)
150
(b)
2.5
2
ln1/(1-x)
Reaction Temperature
1.5
25℃
35℃
1
50℃
0.5
65℃
0
0
10
20
30
40
Reaction time (min)
50
60
(c)
1 
Figure 4.15: Plots of ln 
 vs. reaction time, t, (a) NaOH concentration =0.2 wt.%
1 x 
(b) NaOH concentration =0.6 wt.% (c) NaOH concentration =1 wt.%
103
Table 4.4: Observed reaction rate constant for alkali-catalyzed transesterification (200 rpm)
k'obs (min-1)
Temperature
NaOH
NaOH
NaOH
(°C)
0.2 wt.%
R2
0.6 wt.%
R2
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)
104
Water
Water
Alcohol
Biodiesel
Ester
Dryer
TG
Alcohol
Alcohol
Catalyst
Water
Wash
Water
Batch Reactor
Crude Glycerol
Figure 4.16: Batch reaction process (Source:Van Gerpen, Shanks et al. 2004)
105
Alcohol
Alcohol
Ester
Alcohol
TG
Heater
Alcohol
PFR1
TG
PFR 2
Separator
Catalyst
Ester
Glycerol
Glycerol
Figure 4.17: Plug flow reaction process (Source:Van Gerpen, Shanks et al. 2004)
106
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:
 rate of reactant 


 rate of reactant   rate of reactant   loss due to
  rate of accumulation 

 
 

 
 flowintoelement   flow out of element   chemcialreaction   of reactantin element
 of volume
  of volume
  withinthe element  of volume


 



 of volume



[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:
 rate of reactant 


 loss due to
  rate of accumulation 



0  0  chemcialreaction    of reactantin element



 within theelement  of volume

 of volume



0  0  rAV 
dnA
dt
Here: dnA  nA0 dxA
dt
dt
107
[4.9]
[4.10]
[4.11]
FA
CA0
V0
CA
rA
xA
Figure 4.18: Schematic of a batch reactor
108
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.
t  ts  t f
[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]
109
Table 4.5: Experimental data for slow reaction region (200 rpm)
T=25°C
NaOH
T=35°C
ts
xs
ts
xs
min
%
min
%
0.2
60
1.92
60
4.05
0.6
60
19.65
30
13.96
1
30
11.03
30
16.68
(wt.%)
Cs= CA0(1-xs)
110
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:
VR 
FA0t
CA0
[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:
 rate of reactant 


 rate of reactant   rate of reactant   loss due to


 
 
0
flow
into
element

flow
out
of
element

chemcial
reaction

 
 

 of volume
  of volume
  withinthe element

 

 of volume



where
[4.17]
FA  (FA  dFA )  (rA )dVR
[4.18]
dFA  d FA0 (1 - xA )  FA0dxA
[4.19]
After rearranging, we obtain:
FA0dxA  (rA )dVR
VR  FA0 
Thus:
x Af
0
dxA
 rA
[4.20]
[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 FA0 and required conversion
of xAf is
111
dV
CA0
FA0
CA0
xA0=0
FA
FA+dFA
xA
xA+dxA
Figure 4.19: Schematic of a plug flow reactor
112
CAf
FAf
xAf
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.
VR  Vs  Vf
Vs 
where
[4.22]
FA0t
C A0
[4.23]
V f  FA0 (1  xs )
x Af
xs
dxA
 rA
[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
 rate of reactant 


 rate of reactant   rate of reactant   loss due to


 
 

 flow into element   flow out of element   chemcialreaction  0
 of volume
  of volume
  withinthe element

 

 of volume



[4.25]
FA0  FA0(1 - xAf )  (rAf )dV
[4.26]
where rAf is the reaction rate at exit of the reactor
113
FA0
CA0
V0
FA
CA
rAf
xAf
CA
rAf
xAf
Figure 4.20: Schematic of a continuous stirred tank reactor
114
After rearranging Equation 4.26, the volume of the reactor for CSTRs can be
determined by using Equation 4.27.
VR 
FA0 xAf
(rA ) f
[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.
VR  Vs  Vf
where
Vf 
FA0 (1  xs ) xAf
(rA ) f
115
[4.28]
[4.29]
Table 4.6: Summary of reactor design at different temperatures (200 rmp)
Reactor Volume at Different Temperatures
Reactor Type
25°C and 35°C
VR 
Batch Reactor
PFR
CSTR
VR 
50°C and 65°C
FA0ts
CA0
VR 
FA0 (ts  t f )
CA0
x Af dx
FA0t
A
 FA0 (1  xs )
x
s
CA0
 rA
VR  FA0 
FA0t FA0 (1  xs ) xAf

CA0
(rA ) f
VR 
VR 
Note: 1) For values of ts, xs, and Cs, refer to Table 4.5, 2)
0
dxA
 rA
FA0 xAf
(rA ) f
rA  k'obs CA0 (1  x) ,
(rA ) f  k'obs CA0 (1  x f ) , 3) k'obs values refer to Table 4.4.
116
x Af
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 twophase 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:
117

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
 r  k 'RX CFFAs:
(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.
118
 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.
119

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)
120
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.
121
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