Sim ASPEN Biodiesel Cat MgO

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Simulation of heterogeneously MgO-catalyzed transesterificationfor fine-chemical and biodiesel industrial production

Tanguy F. Dossin, Marie-Françoise Reyniers *, Rob J. Berger, Guy B. Marin

Laboratorium Voor Petrochemische Techniek, UGent Krijgslaan 281 (S5), B-9000 Gent, Belgium

Received 19 January 2006; received in revised form 5 April 2006; accepted 8 April 2006

Available online 5 June 2006

Abstract

A heterogeneous magnesium oxide catalyst is a good alternative for homogeneous catalysts for the transesterification of alkyl esters for theproduction of fine-chemicals as well as for the production of biodiesel. The transesterification of ethyl acetate with methanol was used as a model

reaction to simulate fine-chemical production in a batch slurry reactor at industrial conditions. The transesterification of triolein with methanol to

methyl oleate was chosen to simulate continuous production of biodiesel from rapeseed oil. A kinetic model based on a three-step ‘Eley–Rideal’

type of mechanism in the liquid phase was used in both process simulations. The transesterification reaction occurs between methanol adsorbed on

a magnesium oxide free basic site and ethyl acetate or the glyceride from the liquid phase. Methanol adsorption is assumed to be rate-determining

in both processes. Activity coefficients were required to account for the significant non-ideality of the reaction mixture in the simulations of both

processes. The simulations indicate that a production of 500 tonnes methyl acetate per year can be reached at ambient temperature in a batch

reactor of 10 m3 containing 5 kg of MgO catalyst, and that a continuous production of 100,000 tonnes of biodiesel per year can be achieved at

323 K in a continuous stirred reactor of 25 m3 containing 5700 kg of MgO catalyst. Although various assumptions and simplifications were made

in these explorative simulations the assumptions concerning the reaction kinetics used, the results indicate that for both processes a heterogeneous

magnesium oxide catalyst shows promising potential as a viable industrial scale alternative.

# 2006 Elsevier B.V. All rights reserved.

Keywords: Transesterification; Solid base catalyst; Magnesium oxide; Reactor modeling; Industrial simulation; Biodiesel; Rapeseed oil; Kinetics

1. Introduction

Transesterification of alkyl esters plays an important

industrial role with numerous applications for large and small

volume productions. Large-scale applications are, e.g. the

production of biodiesel [1], polyester or PET in the polymer

industry [2,3] while small-scale fine-chemical production

includes synthesis of intermediates for the pharmaceutical

industry, the production of food additives or surfactants and thecuring of resins in the paint industry [2,4,5]. Nowadays, most

industrial applications are performed in batch or continuous

reactors depending on the production volume. Temperatures

range from 333 to 523 K while pressure varies from 0.1 to

10 MPa and alcohol to ester molar ratios from 5 to 15.

Transesterification reactors are usually coupled with distillation

columns to separate the alcohols from the reaction mixture

[6–11]. Reactive distillation is also currently investigated as

transesterification process [12–15]. Transesterification reac-

tions are mainly performed using acid or base homogeneous

catalysts: sulfuric, sulfonic, phosphoric and hydrochloric acidsas acid catalysts [2,14] and alkaline hydroxides, metal

alkoxides or acetates as base catalysts [2,6]. Base catalysts

are preferred over acid catalysts because of the higher reaction

rates and the lower process temperatures [16]. However, the use

of homogeneous catalysts requires neutralization and separa-

tion from the reaction mixture, leading to a series of

environmental problems related to the use of large amounts

of solvents and energy. Heterogeneous solid base catalysts, able

to catalyze the transesterification of alkyl esters could solve

these issues: they can be easily separated from the reaction

www.elsevier.com/locate/apcatbApplied Catalysis B: Environmental 67 (2006) 136–148

Abbreviations: BuOAc, butyl acetate; BuOH, n-butanol; BuOK, potassium

n-butanolate; D, diolein; EtOAc, ethyl acetate; EtOH, ethanol; FAME, fatty

acid methyl ester; G, glycerine; M, monoolein; M/E, methanol to ethyl acetate

molar ratio; M/T, methanol to triolein molar ratio; MeOAc, methyl acetate;

MeOH, methanol; MgO, magnesium oxide; MeOl, methyl oleate; T, triolein

* Corresponding author. Tel.: +32 9 264 45 17; fax: +32 9 264 49 99.

E-mail address: [email protected] (M.-F. Reyniers).

0926-3373/$ – see front matter # 2006 Elsevier B.V. All rights reserved.

doi:10.1016/j.apcatb.2006.04.008

mailto:[email protected]://dx.doi.org/10.1016/j.apcatb.2006.04.008http://dx.doi.org/10.1016/j.apcatb.2006.04.008mailto:[email protected]


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mixture without requiring the use of a solvent, show easy

regeneration, and have a less corrosive character, leading to

safer, cheaper and more environment-friendly operation.

Recently, numerous studies have been performed to investigate

the use of heterogeneous base compounds as catalysts for

transesterification reactions. This includes the use of catalysts

T.F. Dossin et al. / Applied Catalysis B: Environmental 67 (2006) 136–148 137

Nomenclature

ai activity of component i (mol m3)

a LS liquid–solid interfacial area (m2 m3)

amn group interaction parameter (in calculation

activity coefficient)

A pre-exponential factor (m3 kgcat

1 s

1)C i concentration of component i (mol m

3)

d 1,2 Flory–Huggins combinatorial term

d I impeller diameter (m)

d p solid particle diameter (m)

Di liquid diffusion coefficient of component i

(m2 s1)

Di,eff effective diffusion coefficient of component i in

the pellet pores (m2 s1)

Dim bulk diffusivity of component i (m2 s1)

D0 AB binary diffusivity of component B in A (m2 s1)

E a activation energy (J mol1)

F molar flow (mol s

1

)g acceleration of the gravity (m2 s1)

k BuOK reaction rate coefficient for BuOK catalyzed

reaction (m6 /mol2 s)

k MeOH reaction rate coefficient of methanol adsorption

step (m3 kgcat1 s1)

k LS ,i liquid–catalyst mass transfer coefficient for com-

ponent I (m3 m2 s1)

K eq equilibrium constant of the overall reaction

K A equilibrium constant of alcohol adsorption

(m3 mol1)

li volume parameter (in calculation activity coeffi-

cient)

MWi molecular weight of component i (kg/mol)ni number of moles of component i (mol)

N total number of components

N I impeller revolution speed (s1)

N P impeller power number (=5) (s1)

p pressure (atm)

P yearly production (tonnes year1)

q induction parameter

qi molecular surface area parameter (in calculation


r BuOK BuOK-catalyzed reaction rate (mol m3 s1)

r i van der Waals volume parameter (in calculation

activity coefficient)r MgO MgO-catalyzed reaction rate (mol kgcat1 s1)

R universal gas constant (J mol1 K 1)

Re Kolmogoroff Reynolds number = r3 L d 4 p N pd

5 I N

3i =

ðm3 L V L ÞSci Schmidt number component i = k LS ,i d p / Dim (s)

t time (s)

T temperature (K)

V L liquid volume (m3)

W mass of catalyst (kgcat)

xi molar fraction of component i

X i conversion of component i

Y i product yield of component i

Greek symbols

a acidity parameter

b basicity parameter

g i activity coefficient of component ig 1i activity coefficient of component i at infinite

dilution

g C i combinatorial contribution to activity coefficientof component i

g Ri residual contribution to activity coefficient of

component i

Fi segment fraction (in calculation activity coeffi-

cient)

u i area fraction (in calculation activity coefficient)

G k group residual activity coefficient

G ðiÞk group residual activity coefficient in a reference

solution containing only molecules of type i

e p catalyst porosity

es solid fraction in the slurry reactor

l dispersion parameterLi, j Wilson binary parameter of component i in

solvent j

m L liquid viscosity (kg m1 s1)

n liquid molar volume at 293 K (cm3 mol1)

j asymmetry parameter due to hydrogen bonding

r L liquid density (kg m3)

r p catalyst pellet density (kg m3)

r p,wet density of the catalyst particle filled with liquid

(kg m3)

t polar parameter or catalyst tortuosity

c asymmetry parameter due to polarity difference

cmn group interaction parameter (in calculation


Subscripts and superscripts

0 initial condition* basic site

cat catalyst

eq equilibrium

i component i

in inlet

I impeller

j parameter j

k k th experimental point

out outlet p catalyst pellet

R reactor

s solvent

tot total


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such as alkaline-earth oxides, zeolites, and hydrotalcites

[3,4,17–20]. Nowadays only one industrial application,

developed by the IFP, uses heterogeneous base catalysts for

the transesterification of vegetable oils to produce biodiesel

[21].

Although several authors investigated the kinetics of

transesterification catalyzed by homogeneous base catalysts

[6,14,16,22–24], there is very little information available

concerning the kinetics of the transesterification catalyzed by

solid bases. Hattori et al. [18] have proposed a mechanism for

the transesterification of ethyl acetate with different alcohols

catalyzed by a variety of solid base catalysts, particularly

alkaline-earth-metal oxides. However, this mechanistic study

did neither provide values of activation energies nor rate

constants. Recently, a kinetic model was proposed to describe

the transesterification of ethyl acetate with methanol catalyzed

by a heterogeneous magnesium oxide catalyst [25]. This model

is based on a three-step ‘Eley–Rideal’ type of mechanism

applied in liquid phase where methanol adsorption is assumed

to be rate determining. The proposed value of the activationenergy is 20 kJ/mol and the value of the methanol adsorption

equilibrium coefficient is 3.13 103 m3 /mol. From severalkinetic studies on homogeneous base catalysts followed a

pseudo-second-order rate law with activation energies ranging

from 22 to 83 kJ/mol, depending on the type of alcohol and

ester used [6,14,16,22,23].

The objective of this study was to evaluate the use of the

heterogeneously MgO-catalyzed transesterification reaction in

batch and continuous stirred tank reactors at industrially

relevant conditions using the kinetic model based on the three-

step ‘Eley–Rideal’ type mechanism assuming methanol

adsorption as rate-determining step. Two processes have beensimulated: the transesterification of ethyl acetate with methanol

in a batch slurry reactor for fine-chemical production and the

transesterification of triolein with methanol to form methyl

oleate to simulate biodiesel production from rapeseed oil in a

continuous slurry reactor. The simulation results from both

have been compared with those of homogeneously catalyzed

transesterification processes.

Biodiesel consists of alkyl esters of long chain fatty acids

originating from a renewable resource such as vegetable oils.

Biodiesel can be used in usual diesel engines and presents some

advantages compared to traditional petroleum-based diesel.

Biodiesel is an environment-friendly compound, non-toxic and

emits less carbon monoxide, sulphur compounds, particulatematter and unburned hydrocarbons than traditional diesel.

However, NO x emissions are 10% higher compared to

petroleum-based diesel. Moreover, since biodiesel is produced

from a renewable source, the net production of carbon dioxide

is lowered with reduced contribution to the greenhouse effect.

Biodiesel is usually produced by transesterification of

vegetable oils, largely consisting of triglycerides, with

methanol to produce fatty acid methyl esters (FAME), i.e.

biodiesel, and glycerine.

The catalysts used are typically alkali hydroxides such as

KOH or NaOH and alkali alkoxides such as sodium methoxide

or sodium butoxide. Non-ionic base catalysts such as TBD

(1,5,7-triazabicyclo[4.4.0]dec-5-ene, C7H13N3) are currently

studied [20,26]. The main purpose of transesterification is to

reduce the viscosity of the oil. Vegetable oils cannot be used as

such in common diesel engines because their viscosity is 10–20

times that of conventional diesel, causing injector coking and

engine deposits [1,26]. Biodiesel does not show these problems

and offers the same performance and engine durability as

petroleum-based diesel. Therefore, the importance of biodiesel

increases due to environmental concerns and uncertainties

concerning petroleum price and availabilities.

The transesterification of triolein with methanol to methyl

oleate (CH3(CH2)7CH=CH(CH2)7COOCH3) has been chosen

as a model process for the transesterification of rapeseed oil to

produce biodiesel. Rapeseed oil is the most abundant oil

produced in Europe and triolein (C57H104O6) has been chosen

to represent rapeseed oil as oleic acid is the major fatty acid in

rapeseed oil [26].

2. Modeling procedures

2.1. Transesterification of ethyl acetate with methanol

2.1.1. Thermodynamics

Thermodynamic properties of the components and the

transesterification reaction between 283 and 323 K were

calculated using the ASPEN Engineering SuiteTM 11.1 [27].

The reaction enthalpy varies from 2.8 at 273 K to 3.2 kJ/ mol at 328 K, showing a slightly exothermic reaction. The

value of the equilibrium coefficient K eq ranges from 1.875 at

293 K to 1.716 at 323 K, which is similar to values reported in

previous studies of transesterification reactions [14,22]. Theequilibrium conversion of ethyl acetate X eq varies from 9 to

95% at molar ratios of methanol to ethyl acetate (M/E) ranging

from 0.1 to 10.

The non-ideal character of the reaction mixture was assessed

and activity coefficients g i for each component i were

calculated at temperatures ranging from 283 to 323 K based

on the Wilson equation as follows [28]:

ln g i ¼ ln

X N j¼1

x jLi; j

þ 1

X N k ¼1

xk Lk ;iP N j¼1 x jLk ; j

(1)

where, N , xi and L i, j are the number of components, the molefraction of each component and the Wilson binary parameter of

component i diluted in solvent j. The Wilson binary parameters

are calculated based on Eq. (1) assuming infinite dilution of

component i in solvent j. The activity coefficient at infinite

dilution is then calculated based on Eq. (2) by means of the

MOSCED method [28]:

ln g 1i ¼ ni

RT

ðls liÞ

2 þq2sq

2i ðt s t iÞ

2

cs

þðas aiÞðbs biÞ

j s

þ d s;i (2)

T.F. Dossin et al. / Applied Catalysis B: Environmental 67 (2006) 136–148138


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where g 1i is the activity coefficient of component i in solvent s

at infinite dilution, and a, b, l, n, j , t , c, q and d s,i are tabulated

parameters, specific for each component.

Table 1 shows the influence of the M/E ratio at 303 K on the

activity coefficients of each component at conditions corre-

sponding to the start of the reaction and to the equilibrium

conversion. Since the calculated activity coefficients were

fairly different from 1.0, the non-ideality of the mixture was

taken into account in the simulations.

2.1.2. Kinetic model

The kinetic model used to describe the transesterification of

ethyl acetate with methanol to form methyl acetate and ethanol

is based on a three-step ‘Eley–Rideal’ type mechanism [25].

The elementary steps of the mechanism are given in Table 2.

The first step describes the adsorption of methanol on a

catalytically active site. Then, adsorbed methanol reacts with

ethyl acetate from the liquid phase to form methyl acetate and

adsorbed ethanol. Ethanol finally desorbs in the third step. The

proposed model assumes methanol adsorption as rate-deter-

mining step, leading to the following rate equation:

r MgO ¼k MeOHðaMeOH ð1=K eqÞðaMeOAcaEtOH=aEtOAcÞÞ

1 þ ðK A=K eqÞðaMeOAcaEtOH=aEtOAcÞ þ K AaEtOH(3)

where k MeOH is the methanol adsorption rate constant, ai, the

activities of reaction component i, K eq the equilibrium coeffi-

cient of the overall transesterification reaction and K A is the

equilibrium constant of the methanol adsorption step. The value

of the adsorption equilibrium coefficient is assumed to be the

same for both alcohols. The values of the activation energy E aand the pre-exponential factor A of the rate constant k MeOH as

well as the adsorption equilibrium coefficient K A have been

estimated and their values are presented in Table 3. Theparameter estimation procedure and its results are described

elsewhere [25]. The value of K eq has been calculated based on

thermodynamic data for all compounds.

According to Madon and Iglesia [29], it is appropriate in

non-ideal systems to use the activities instead of the

concentrations in the rate expression if the rate-determining

step is an adsorption or desorption step, because the

corresponding activated complexes may be solvated by the

surrounding fluid phase. However, if the solvation effect of the

activated complex is similar to that of the reactants (products),

the concentrations instead of the activities should be applied

because a cancellation of activity coefficients leads to the

apparent dependence of rates on concentration rather than

activities. For the current reaction, it is unclear which situation

applies and therefore the effect of this was investigated, thus

using the following rate equation:

r MgO ¼k MeOHðC MeOH ð1=K eqÞðaMeOAcaEtOH=aEtOAcÞÞ

1þ ðK A=K eqÞðaMeOAcaEtOH=aEtOAcÞ þ K AC EtOH(3a)

Since the location of the thermodynamic equilibrium in non-

ideal mixtures is determined by the equilibrium constants in

combination with the component activities rather than the

component concentrations, the activities are maintained in

the terms correcting for the equilibrium. This thus results in

a rather unusual rate expression containing activities as well as

concentration terms. A similar form of rate equation, however,

has been proposed by Madon and Iglesia [29] for the cyclo-

hexene hydrogenation over Pd.

Simulations results using rate Eq. (3a) are compared to

simulation results using Eq. (3) to study the possible influence

of solvation in the non-ideal liquid phase.

2.1.3. Reactor model

The slurry reactor, operated batch wise or in a continuous

way, was chosen as industrial reactor because of its easy

operability and frequent use in fine-chemical industry. The


Table 1

Influence of the M/E ratio on the activity coefficients of MeOH, EtOH, MeOAc,

EtOAc at 303 K, at the start of the reaction and at equilibrium

M/E ratio 0.1 1 5 10

Conversion (mol%) 0 9 0 58 0 89 0 95

MeOH 3.08 3.11 1.29 1.30 1.03 1.03 1.01 1.01

EtOH – 2.79 – 1.26 – 1.09 – 1.09MeOAc – 1.00 – 1.30 – 1.90 – 2.10

EtOAc 1.02 1.02 1.42 1.42 2.22 2.24 2.51 2.53

Table 2

Elementary reactions of Eley–Rideal type reaction mechanism

CH3OH + * CH3OH*

CH3OH* + CH3COOCH2CH3 CH3CH2OH

* + CH3COOCH3CH3CH2OH

* CH3CH2OH + *

Table 3

Parameter values of the kinetic model used

A (m3 /kgcat s) 0.148

E a (103 J/mol) 20.1

K A (103 m3 /mol) 5.29

Table 4

Range of process parameters

Conditions

T (K) 293–323

p (atm) 1

M/E () 0.1–30W (kg) 5–10000

F inEtOAc(mol s1) 1–100

Reactor

V L (m3) 10–40

d R (m) 2.3–3.7

h R (m) 2.3–3.7

N I (s1) 1

d I (m) 1.2–1.9

Catalyst

d p (106 m) 25

e p 0.45

t p 5.0

r p (kg/m3) 2460


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reactor was considered to be cylindrical with equal reactor

diameter and height and with a volume between 10 and 40 m 3.

The range of process parameters considered in the simulations

is given in Table 4. Since transesterification reactions are

equilibrium-controlled reactions, an excess of methanol is used

to shift the equilibrium towards the product side. The methanol

to ethyl acetate molar ratio (M/E) varies from 5 to 30 [16]. The

reactor geometry follows the standard tank geometry for

turbulent mixing with a radial flow impeller proposed by

Tatterson [30]. The impeller is a traditional Rushton turbine

with a diameter from 1.2 to 1.9 m. The impeller speed is

calculated to fulfill perfect mixing and complete suspension of

solid particles, following Eq. (4) [31]:

N I >

12

g

d I

0:5 r p;wet r L r L

0:5d p

d I

0:165e

0:25S (4)

where N I is the impeller speed, d I , the impeller diameter, g, the

gravitational constant, r p;wet, the density of the catalyst pellet

filled with liquid,r L , the liquid density, d p, the particle diameterand es is the catalyst fraction in the slurry. Since the criterion

was only validated in the range 0.01 < eS


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values are close to 1.0 since the reactions are almost

thermoneutral and the entropy change is likely to be very

small as well. For simplicity, all the equilibrium constants were

taken equal to 1.0 in the simulations, independent of

temperature.

Since it was observed that the reaction mixture consisting of

ethyl acetate, methanol, methyl acetate and ethanol was not

ideal (see Section 2.1.1), the need for taking non-ideality into

account was also checked for the biodiesel simulations.

However, unlike the transesterification of ethyl acetate with

methanol, the MOSCED method could not be applied to solve

the Wilson equation since data for glyceride molecules are nottabulated. Therefore, the UNIFAC contribution method was

used to calculate activity coefficients [28]. The activity

coefficient is separated into two parts: one part provides the

contribution due to molecular size and shape and is defined as

the combinatorial contribution (ln g C i ), and the other provides

the contribution due to molecular interactions, known as the

residual contribution (ln g Ri ):

ln g i ¼ ln g C i þ ln g

Ri (10)

where

ln g C i ¼ lnFi

xiþ 5qiln u i

Fiþ li Fi

xi

Xallcomp: j

x jl j and

ln g R ¼Xallgroupsk

nðiÞðlnG k lnG ðiÞk Þ;

u i ¼ xiqiP

j x jq j; Fi ¼

xir iP j x jr j

;

li ¼ 5ðr i qiÞ ðr i 1Þ; r i ¼Xk

nðiÞk Rk ;

qi ¼Xk

nðiÞk Qk

and

lnG k ¼ Qk

1 ln

Xallcomp:m

u mC mk

Xallcomp:m

u mC kmPallcomp:n u nC nm

;

C mn ¼ exp

amn

T

:

Rk and Qk are group volumes and area parameters and aretabulated as a function of the type of group. T is the temperature

expressed in Kelvin, nðiÞk is the number of groups k present in the

molecule i and xi is eitherthe molar fraction of component i in the

mixture or the mole fraction of group i in the component. Table 5

shows the influence of the M/T ratio at 323 K on the activity

coefficientsof each component andat thestart of thereactionand

at equilibrium in terms of mole fraction. Since the calculated

activity coefficients were fairly different from 1.0, the non-

ideality of the mixture was taken into account in the simulations.

2.2.2. Kinetic model

Several kinetic studies of methyl ester production from

different vegetable oils have been performed using homo-geneous base catalysts [6,16,20,23]; however, to the best of our

knowledge, no kinetic data of vegetable oil transesterification

involving heterogeneous catalysts have been reported. All

kinetic studies performed with homogenous catalysts show that

all three consecutive transesterification reaction rates follow a

pseudo-second-order law with respect to the glyceride and the

alcohol. As can be seen in Freedman’s work [16], the three

consecutive transesterification reactions have significantly

different activation energies.

Since the reaction mechanism of the transesterification

reactions over the MgO catalyst is assumed to be of the

Eley–Rideal type with the adsorption of methanol being the


Table 5

Influence of the M/T ratio on the activity coefficients (g i) of all components at 323 K, at the start of the reaction and at equilibrium ( xi = mole fraction)

Triolein Diolein Monoolein Methyl oleate Glycerine MeOH

M/T = 5

xi 0.16648 0.00004 0.00006 0.00029 0.00003 0.83311

g I 1.34479 0.65299 0.83860 2.15770 28.27353 1.23042

xi (equilibrium) 0.01987 0.06435 0.07885 0.23293 0.00368 0.60033

g I 1.23714 0.60129 0.77292 2.06543 26.08368 1.26260

M/T = 10

xi 0.09086 0.00000 0.00002 0.00012 0.00003 0.90897

g I 2.58711 0.71287 0.51951 3.67922 9.93927 1.12314

xi (equilibrium) 0.00432 0.02538 0.05645 0.15259 0.00478 0.75648

g I 2.35993 0.64583 0.46744 3.50896 8.88203 1.13984

M/T = 20

xi 0.04759 0.00000 0.00001 0.00007 0.00002 0.95231

g I 6.56754 0.97234 0.38074 6.66597 3.91396 1.05075

xi (equilibrium) 0.00077 0.00826 0.03344 0.09059 0.00516 0.86177

g I 6.22731 0.90549 0.34822 6.45176 3.51569 1.05628


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rate-determining step, the kinetic model used for simulation of

the three heterogeneous transesterification reactions over the

MgO catalyst is assumed to be the same as that for the

transesterification of ethylacetate, as discussed in Section 2.1.2.

As this mechanism is of the Eley–Rideal type with the

adsorption of methanol being the rate-determining step, it can

be assumed that the values of the activation energy E a and the

pre-exponential factor A for each of the three transesterifica-

tions are equal to the values for the transesterification of ethyl

acetate with methanol, as presented in Table 3. The rate

equations for the three glyceride transesterifications are thus

assumed to be, analogous to Eq. (3):

r T ¼

k MeOH

aMeOH

1K eq

aMeOla DaT

1þ K AK eq

aMeOla DaT

þ K Aa D þ K Aa M þ K AaG(11a)

r D ¼

k MeOH

aMeOH

1K eq

aMeOla M a D

1 þ K AK eq

aMeOla M a D

þ K Aa D þ K Aa M þ K AaG (11b)

r M ¼

k MeOH

aMeOH

1K eq

aMeOlaGa M

1 þ K AK eq

aMeOlaGa M

þ K Aa D þ K Aa M þ K AaG(11c)

where k MeOH is the methanol adsorption rate constant equal for

all three reactions, ai, the activities of each reaction component,

K eq the equilibrium coefficient of the overall transesterification

reaction and K A is the adsorption equilibrium constant of the

alcohols present, assumed to be the same for all alcohols

present. The elementary steps and the corresponding reactionmechanism for each transesterification reaction are shown in

Table 6.

It should be mentioned that the assumption of exactly the

same reaction kinetics for the transesterification of large

glyceride molecules as that for the transesterification of ethyl

acetate is rather speculative. The long organic tails of the

glyceride molecules will probably have some influence on the

kinetics. Additionally, the assumption of equal rate coefficients

for the three consecutive reactions seems rather speculative in

view of the significant differences between the activation

energies of the three consecutive transesterification reactions

observed with homogeneous catalysts, as mentioned above.

Similarly as for batch slurry reactor simulations (see Section

2.1.2), the results obtained using the rate equations based on

the component activities (11a)–(11c) were compared to the

results using the rates based on the component concentrations

(Eqs. (12a)–(12c)):

r T ¼k MeOH

C MeOH

1

K eq;1

aMeOla D

aT

1 þ K AK eq;1

aMeOla DaT

þ K AC D þ K AC M þ K AC G(12a)

r D ¼

k MeOH

C MeOH

1K eq;2

aMeOla M a D

1 þ K AK eq;2

aMeOla M a D

þ K AC D þ K AC M þ K AC G(12b)

r M ¼

k MeOH

C MeOH

1K eq;3

aMeOlaGa M

1 þ K AK eq;3

aMeOlaGa M

þ K AC D þ K AC M þ K AC G(12c)

2.2.3. Reactor model

As for the simulation of fine-chemical production, a slurry

reactor has been chosen to simulate continuous biodiesel

production. All assumptions mentioned for the simulations in

the EtOAc/MeOH case are also valid for the biodiesel

production process (see Section 2.1.3). Additionally, the

continuous slurry reactor was assumed to be at steady state,

start-up time was not taken into account and the catalyst was

retained in the reactor by a microporous filter.

For simplicity reasons, the possibility of phase separation

between alcohols and glycerides was not included in the

explorative simulations discussed in this paper. According toNoureddini and Zhou [23], this phase separation disappears

rapidly if a high impeller speed of at least 5 s1 is applied since

methyl oleate functions as mutual solvent to form a single-

phase system as reaction advances. However, it seems more

likely that the good mixing leads to a micro-emulsion of a polar

and a non-polar liquid phase [1]. Particularly the high activity

coefficients of methyl oleate and glycerol (Table 5), which were

calculated assuming one homogenous liquid phase, indicate the

possibility of such a phase separation.

The existence of a phase separation may strongly influence

the reactor performance. On the one hand it may negatively

influence the performance if mass transport limitations occur

between both phases. On the other hand it may influence theoverall equilibrium composition, since a phase separation

generally leads to a significant decrease of the values of the

activity coefficients. Table 5 shows the composition at

equilibrium as a function of the M/T if no phase separation

is taken into account. For instance, at a M/T ratio of 10, the

maximum achievable mole fraction of methyl oleate amounts to

0.1526, corresponding to a yield of only 0.56. According to

these calculations, the reaction product will thus always contain

significant amounts of diolein and monoolein besides methyl

oleate, as shown in Table 5.

Experimental studies using homogeneous alkali catalysts,

however, have shown that methyl ester yields of over 0.9 can be


Table 6

Elementary steps and reaction mechanisms for the biodiesel transesterification

reactions

Elementary reactions Reaction mechanism

T ! D D ! M M ! G

CH3OH + * CH3OH* 1 1 1

CH3OH* + T D* + MeOl 1

CH3OH* + D M* + MeOl 1

CH3OH* + M G* + MeOl 1

D* D + * 1

M* M + * 1

G* G + * 1


8/13

reachedatanM/Tabove5 [23,24,46,39]. Besides the existence of

a phase separation, this discrepancy may also originate from the

estimation of the equilibrium constants or from the UNIFAC

contribution method used to estimate the activity coefficients.

Nevertheless, to the best of our knowledge, the UNIFAC

contribution method is currently the best method to estimate

activity coefficients of complex and non-tabulated molecules.

The occurrence of phase separation therefore seems to be the

most likely explanation of the discrepancy since the activity

coefficients of the final twoproducts, methyl oleate and glycerol,

are the highest of all components, thus limiting the equilibrium

conversion to a relatively low value, which can be understood

based on Eq. (11c); a decrease of the activity coefficients as a

result of a phase separation will thus increase the conversion to

the desired product methyl oleate at equilibrium. Due to the

significant deviation of the equilibrium conversion, the simula-

tion results are to be interpreted in a qualitative way only.

Another important aspect in this process is the presence of

mass transport limitations around and inside the catalyst

particles as a result of the relatively low diffusivity, which iscaused by the high viscosity of the reaction mixture. It may give

rise to significant concentrations gradients. In contrast, heat

transport limitations do not have to be taken into account as a

result of the thermoneutrality of the reaction. In order to take

the effect of mass transfer limitation into account, the reactor

model equation in the bulk phase for a continuous slurry reactor

considering steady-state operation and mass transport limita-

tions is as follows:

F ini F outi ¼ V L k LS ;ia LS ðC L ;i C S ;iÞ (13)

where F ini and F outi are inlet and outlet molar flows of compound

i, k LS ,i the liquid–solid mass transfer coefficient of the compo-nent i in methanol, assumed to be in excess in the reactor, and

a LS the liquid–catalyst interface area, and V L is the total liquid

volume. C L ,i and C S ,i are the concentrations in the liquid phase

and at the external pellet surface, respectively.

The mass transfer coefficient is calculated from Sano et al.

[40]:

k LS ;i ¼ 2þ 0:4 Re0:25Sc0:333i (14)

The required liquid density and viscosity were calculated as the

weight-weighted average value of the values of the individual

components, which were takenfrom various sources as indicated

in Table 7.

Since uniform spherical catalyst particles are assumed, the

mass balance inside the pellet is, applying spherical coordi-

nates:

4

e pd 2 p

1

j 2@

@j

j 2 Deff ;i

@C i

@j

¼ r

r p

e p

(15)

Deff,i and C i are the effective diffusion coefficient and theconcentration of component i in methanol inside the pellet;

e p and r p are the porosity and the density of the catalyst pellet,

respectively. j is the spatial coordinate inside de pellet, equal to

1 at the surface and 0 at the centre of the pellet. r is the net rate

of formation of this specific component. The effective diffu-

sivity Deff,i depends on the concentrations through the depen-

dence of the activity coefficients on the mole fractions. The

effective diffusion coefficient was estimated from [41]:

Deff ;i ¼ e p

t p Dim (16)

The bulk diffusivity Dim was calculated from the binary diffu-

sion coefficient in methanol ( A = i, B = methanol) [42] and

corrected for the effect of non-ideality [43] using:

D AB ¼ ð x A D0 BA þ x B D

0 ABÞ

1 þ

@ ln g A@ ln x A

(17)

Since x A D0 BA x B D

0 AB, this equation was simplified to:

D AB ¼ x B D0 AB

1 þ

@ ln g A@ ln x A

(18)

resulting in:

Dim ¼ xMeOH D0

im

1 þ

@ ln g i@ ln xi

(19)

At the pellet surface, the effective diffusion of the component i

inside the pellet is equal to the diffusion of this component from

the bulk phase to the pellet surface. This is expressed by the

following equation:

2 Deff ;i

k LS ;id p

@C i

@j

j ¼1

¼ C L ;i C S ;i (20)

Eq. (13) was applied to all reaction compounds and the reaction

rate r in this equation is the sum of the three transesterification

reaction rates (Eqs. (11a)–(11c)) taking the reaction stoichio-

metry into account. Simulations were performed with the


Table 7

Pure component densities and viscosities used in the biodiesel production process simulations

Component Densitya (kg/m3) Source Viscosityb (mPa s) T 0 (K) Source

Triolein 1113–0.684T [27,47] 45.1 313 [47,48]

Diolein 1042–0.42T [27] 68 313 [49]

Monoolein 1108–0.543T [27] 91.5 313 [49]

Methyl oleate 1082–0.712T [27,47] 5.2 313 [48,47,50]

Glycerine 1453–0.656T [27,51] 1500 293 [51]

Methanol 1079–0.971T [27,51] 0.55 298 [51]

a T : temperature (K).b The Lewis and Squires liquid viscosity–temperature correlation [28] wasusedto obtain theviscosity at therequiredtemperature:m0:2661T ¼ m

0:2661T 0

þ T T 0233

, with

mT = liquid viscosity at the required temperature T (K), and mT 0 = known viscosity at temperature T 0 (with viscosities expressed in mPa s).


9/13

Athena Visual Workbench software developed by Stewart &

Associates Engineering Inc., using the PDAPlus solver. This

software allows the solution of boundary value problems

involving partial differential equations. The integration of

the catalyst particles was performed using the method of finite

differences with 150 discretization points.

The biodiesel production in a continuous slurry reactor

typically amounts to 100,000 tonnes year1 [44,45]. The

applicable temperature range is narrow due to the high

viscosity of the glycerides and the low boiling point of

methanol. Since the melting point of the monoglyceride is

about 310 K and the methanol boiling point is about 340 K, the

operating temperature window only ranges from 320 to 330 K.

Nowadays, the methanol to oil molar ratio is typically set to 6,

although a recent patent showed that higher molar ratios can

improve the yield sufficiently to make it economically

attractive [46]. Therefore, the operating molar methanol/

triolein ratio will be in the range between 5 and 30.

The conversions and the yields were calculated according to:

xT ¼ 1 F T ;out

F T ;feed; xMeOH ¼ 1

F MeOH;out

F MeOH;feed(21)

Y MeOl ¼F MeOl;out

3F T ;feed; Y D ¼

F D;out

F T ;feed; Y M ¼

F M ;out

F T ;feed;

Y G ¼ F G;out

F T ;feed

(22)

The yearly production of methyl oleate and glycerine were

calculated as follows:

PMeOl ¼ 3Y

MeOl W

MWMeOl

60

60

24

365

1000ðW =F T ;feedÞ

ðtonnes=yearÞ (23)

PG ¼Y G W MWG 60 60 24 365

1000ðW =F T ;feedÞ ðtonnes=yearÞ

(24)

3. Results and discussion

3.1. Batch slurry reactor for fine-chemical production

Simulations of a batch slurry reactor for the heterogeneouslyMgO-catalyzed transesterification of ethyl acetate with

methanol at industrial process conditions have been performed

using Eq. (5). Influences of temperature, amount of catalyst and

molar ratio have been investigated to achieve a typical

production of 500 tonnes MeOAc year1. Traditionally, batch

times do not exceed 2 days.

Table 8 shows the evolution of the annual production as a

function of temperature ranging from 293 to 323 K, at M/E = 10,

W = 2 kg, V L = 10 m3. Fig. 1 shows the corresponding ethyl

acetate conversion as a function of the batch time. The minimum

required batch time is defined as the time to reach 95% of the

ethyl acetate equilibrium conversion. At the conditions stated

above, this corresponds to an EtOAc equilibrium conversion of

90%. As expected, the minimum required batch time decreases

with increasing temperature. For the considered process

parameters, a yearly production of almost 500 tonnes year1

is already achieved at 323 K (see Table 8). Nevertheless, it is

useful to study the influence of the other process parameters such

as the amount of catalyst and the M/E molar ratio on the yearly

production. The influence of the catalyst mass on the yearly

production is presented in Table9. From these results follows that

a similar yearly production (524 tonnes year1) can be achieved

at 293 K if the amount of MgO catalyst is increased from 2 to4.5 kg (M/E = 10, V L = 10 m

3). Running the process at 293 K

could eliminate theneed forreactor heating, whichmay make the

process easier and less expensive, assuming the cost of the

catalyst is much smaller than the cost of the process equipment

combined with the energycost. Theinfluence of the M/E ratio on

the minimum required batch time at 293 K, W = 5 kg,

V L = 10 m3 is given in Table 10 while Fig. 2 shows the evolution

of ethyl acetate conversion as a function of time for differentM/E

ratios at the same reaction conditions. The minimum required

batch time needed to reach 95% of the equilibrium conversion is

reduced when M/E is increased. The yearly production is also

increased. However, the absolute amount of product obtained perbatch is also reduced since the initial amount of ethyl acetate is

lowerat higherM/E molarratios. This results in an increase of the

number of batches and thus an increase of the time needed for

emptying and reloading the reactor, thus resulting in an overall

productivity loss. A more in-depth productivity analysis is

thereforeneeded to determinethe optimal M/Eratio to maximize

the profit.

The effect of using concentrations instead of activities in the

rate expressions, as discussed in Section 2.1.1, was found to be

small around the optimal process conditions, i.e. 293 K and

M/E = 10, as shown in Table 10 (compare the first entry with

the second entry). At M/E > 1, the activity coefficients of

methanol and ethanol are close to 1.0 and only a small influencecan thus be expected.

A comparison has been made with homogeneously

catalyzed transesterification. Schmidt et al. have proposed a


Table 8

Influence of the temperature on the annual production (simulation performed using Eqs. (3) and (5) at M/E = 10, W = 2 kg, V L = 10 m3, X EtOAc = 90%)

Temperature (K) Minimum required batch time (h) MeOAc production (tonnes year1) batches year1

293 51.1 233.5 171.3

303 39.6 301.3 221

313 31.8 375.1 275.2

323 25.6 465.3 341.3


10/13

kinetic model for the transesterification of butyl acetate withethanol catalyzed by potassium butanolate to form ethyl acetate

and butanol [22]. The activation energy and the pre-exponential

factor of the rate coefficient were derived from experimental

data, based on a second-order rate law (Eq. (25)) and

accounting for the polarity of the reaction mixture:

r BuOK ¼ k BuOK

C BuOAcC EtOH

C BuOHC EtOAc

K eq

(25)

This model has been applied to the homogeneous catalyzed

transesterification of ethyl acetate with methanol using the

kinetic parameters reported by Schmidt et al.: activation energyand pre-exponential factor amount to 51 kJ/mol and 24.4 m6 /

mol2 s, respectively [22]. Note that this activation energy is

twice that for the heterogeneously catalyzed reaction used in

this paper. A simulation has been performed at the same process

conditions as the heterogeneous catalyzed industrial transes-

terification, i.e. 293 K, M/E = 10, V L = 10 m3 integrating

Eq. (5) using Eq. (25) as rate equation. In order to achieve

the same yearly production of 500 tonnes year1, 0.150 mol/m3

catalyst concentration is needed, i.e. 0.170 kg of potassium

butanolate catalyst for a comparable batch time and a 10 m 3

reactor volume. The required mass of homogeneous catalyst is

thus almost 30 times smaller than the mass of solid magnesium

oxide catalyst to achieve the same production. However,

the lower cost of the solid magnesium oxide catalyst, the

elimination of the reaction mixture neutralization step and

the easy catalyst separation from the reaction mixture make

the heterogeneous process a promising alternative to the

homogeneous process.

3.2. Continuous slurry reactor for biodiesel production

Simulations were performed of the industrial scale biodiesel

production in a continuously perfectly mixed slurry reactor by

transesterification of triolein with methanol using an hetero-geneous MgO catalyst. The influence of the temperature, the

space time and the MeOH/triolein feed molar ratio (M/T ratio)

on the production has been investigated to achieve about

100,000 tonnes of methyl oleate per year. As discussed in

Section 2.2.3, the results are to be interpreted in a qualitative way

only, since the equilibrium conversion to the desired product

methyl oleate is much lower than the ones found experimentally.

In order to obtain a high productivity, the amount of catalyst

was chosen quite high: 5700 kg. In a total liquid volume of

25 m3, the volume fraction of magnesium oxide catalyst

amounts to about 0.091. According to Eq. (4), with this catalyst

fraction, the impeller revolution speed should be at least0.76 s1 at the standard conditions: 323 K, a M/T molar ratio of

10, a space time of 1000 kg-cat s/mol-triolein, and a catalyst

particle size of 25 mm. Therefore, an impellor speed of 1.0 s1

seems sufficient. However, impeller revolutions of 5 and 10 s1


Table 9

Influence of the catalyst mass on the annual production (simulation performed using Eqs. (3) and (5) at T = 293 K, M/E = 10, V L = 10 m3, X EtOAc = 90%)

W (kg) Minimum required batch time (h) MeOAc production (tonnes year1) Batches year1

2 51.1 233.5 171.3

4.5 22.7 525.5 385.4

5 20.4 583.9 428.2

10 10.2 1167.8 856.515 6.8 1751.6 1284.7

Table 10

Influence of the M/E molar ratio on the annual production (simulation performed using Eqs. (3) and (5) at T = 293 K, W = 5 kg, V L = 10 m3, X EtOAc = 95% of X eq)

M/E Minimum required batch time (h) MeOAc production (tonnes year1) Batches year1

10 20.4 583.9 428.2

10a 19.9a 601.4a 441.1a

15 9.9 873.8 883.3

20 5.7 1163.3 1545.8

30 2.8 1637.9 3122.1

a

Simulations using rate Eq. (3a) with concentrations instead of activities.

Fig. 1. Conversion of ethyl acetate vs. batch time at different temperatures, at

M/E = 10, W = 2 kg, V L = 10 m3: (—) 293 K; ( ) 303 K; (– – –) 313 K;

(– - – -) 323 K. The simulated values were obtained from Eqs. (3) and (5)

using the parameter values of Table 3.


11/13

were chosen in order to achieve a homogeneous mixture (or amicro-emulsion) of the methanol and glycerides.

The reaction rates were calculated using Eqs. (12a)–(12c).

The conversion of the triolein, together with the yearly

production of methyl oleate, is plotted in Fig. 3 as a function of

space time at the standard set of reactor conditions: M/T feed

ratio = 10, T = 323 K, d p = 25 mm, V L = 25 m3, triolein feed

molar flow = 5.0 mol s1, W cat = 5700 kg, N I = 5 s1. As

expected, the conversion increases with space time until

equilibrium conversion is reached. The yearly production of

methyl oleate decreases strongly with increasing space time

due to the decrease of the reaction rate upon approaching the

equilibrium composition. Working at short space times willthus result in a high production; however, the quality of the

biodiesel will be worse due to the larger amounts of unreacted

triolein and diolein. This can be seen in Fig. 4 in which the

yields of the products are plotted as a function of space time.

For the explorative simulations discussed in this paper, it was

assumed that a production of 100,000 tonnes methyl oleate per

reactor per year is typical. Additionally, it was assumed that theminimum required triolein conversion and methyl oleate yields

are close to the values shown in Table 11. Of course, a thorough

economic evaluation is needed to determine the real optimum

operation conditions, which includes the maximum allowed

contents of remaining glycerides in the biodiesel according to

the ASTM or DIN standards, and also the value of the

byproduct glycerine. Currently glycerine is still a valuable

product that can be used in several industrial applications;

however, its value will decrease once the biodiesel production

takes off on a large scale.

The influence of temperature, M/T ratio, particle diameter,

and impeller revolution speed on conversions and yields wereinvestigated and some of these results are presented in Table 11.

According to the simulations, a production of 100,000 tonnes

methyl oleate is achieved at a space time of 871 kg s/mol-

triolein at 323 K, M/T feed ratio = 10, d p = 25 mm and

N I = 5 s1. A triolein conversion of 93% is achieved while

methyl oleate and glycerine yields amount to 0.54 and 0.05,

respectively. At the same conditions, according to Eq. (24),

944 tonnes of glycerine per year are produced. Separation from

methyl oleate is easy since glycerine is practically immiscible

with methyl oleate and therefore can be recovered by settling

(decantation) [7] or centrifugation [46].

Increasing theM/T ratio from 10 to 20 results in a significantly

higher triolein conversion and also higher yields of methyl oleateand glycerine (compare first entry with the second entry in

Table 11) as was also pointed out by Boocock [46]. These yields

rise from 0.54 and 0.05 to 0.62 and 0.10, respectively. However,

this increase of M/Tratio also increasesmethanol flow, leading to

a larger effort to recycle the methanol, as pointed out by

Connemann et al. [7] and Zhang et al. [8]. The extra cost for

separating the methanol from the reaction mixture by means of

distillation will thus compensate the advantage of the higher

yields of methyl oleate and glycerine. An economic evaluation

will thus be needed to estimate the optimal M/T ratio.

It appears that the influence of the temperature, as well as

that of the catalyst particle size on the reactor performance is


Fig. 2. Conversion of ethyl acetate vs. batch time at different M/E ratios, at

T = 293 K, W = 5 kg, V L = 10 m3: (—) M/E = 10; ( ) 15; (– – –) 20; (– - – -)

30. The simulated values have been obtained using Eqs. (3) and (5). The

simulated values were obtained from Eqs. (3)and (5) using the parameter values

of Table 3.

Fig. 3. Conversion of triolein (—; left axis) and the yearly production (kilo-

tonnes year1) of methyl oleate (– – –; right axis) and glycerine ( ; right axis)as a function of space time (kgcat s/mol-triolein) at 323 K, M/T = 10,

W = 5700 kg, V L = 25 m3, N I = 5 s

1, d p = 25 mm. The simulated values were

obtained using Eqs. (11a)–(11c), (13), (15), and (20), and using the kinetic

parameter values from Table 3.

Fig. 4. Yields of diolein (– – –), monoolein (– - – -), glycerine ( ), andmethyl oleate (—) as a function of space time (kgcat s/mol-triolein) at

323 K, M/T = 10, W = 5700 kg, V L = 25 m3, N I = 5 s

1, d p = 25 mm. The simu-

lated values were obtained using Eqs. (11a)–(11c), (13), (15), and (20), and

using the kinetic parameter values from Table 3.


12/13

very small. This is due to the rather close approach to

equilibrium, which is neither influenced by the temperature,

nor by the particle size. However, since a rather strong diffusion

limitation of triolein occurs at low conversion, the use of smaller

catalyst particles will significantly increase the initial conversion

rate whereas the temperature has only a minor effect. The strong

diffusion limitation of triolein at low conversion is due to the

combination of the relatively fast reaction rate and the fact that

the reaction rate is zeroth order with respect to triolein. This

zeroth order causes the rate to remain high if the triolein

concentration already decreases due to the diffusion limitation.Obviously, both external and internal transport limitations

disappear when equilibrium is approached.

Also the impeller revolution speed has very little influence

on the reactor performance since the conversion rate is not

significantly influenced by external mass transfer limitation

(i.e. transfer from the liquid bulk towards the catalyst particles).

Nevertheless, it remains crucial to check whether the stirring

rate is really sufficient to achieve a single-phase system or

micro-emulsion system and a good suspension of the catalyst

particles in the reactor.

The effect of using concentrations instead of activities in the

rate expressions, as discussed in Section 2.2.2 was found to besmall at all conditions considered. At the standard conditions the

calculated effect is shown in Table 11 (compare the first entry

with the third entry). Apparently, the net effect of the activity

coefficients on thevalue of thedenominators of therate equations

is small. The numerator of the rate equations hardly changes

since the activity coefficient of methanol is always close to 1.0.

As for the EtOAc/MeOH-case, the biodiesel production

catalyzed by homogeneous catalysts, based on data from Zhang

et al. [8] and Connemann et al. [45], has been compared to the

heterogeneously catalyzed transesterification. According to

Zhang et al., a stream of 10 kg/h of sodium hydroxide is needed

to reach a production of 8000 tonnes biodiesel year1 at 333 K

with M/E = 6. This leads to a yearly catalyst consumption of 88tonnes. Based on Connemann’s data on existing biodiesel

plants, which typically operate at 333–348 K, 3–4 kg catalyst

per tonne of biodiesel is needed to reach a yearly production of

80,000–125,000 tonnes of biodiesel. This thus requires more

than 300 tonnes of homogeneous base catalyst. These amounts

are much larger than the estimated amount of 5.7 tonnes of

heterogeneous MgO catalyst that follows from the simulations

in this paper. In contrast to the batch reactor simulation for the

EtOAc/MeOH-case, the heterogeneous catalyst thus seems to

be much more efficient for biodiesel production than the

homogeneous catalysts. Although this is an encouragement to

intensify the research on the heterogeneously catalyzed system,

it must be interpreted with care in view of the assumptions and

simplifications made in the explorative simulations presented

here. Particularly the assumption of the reaction kinetics to be

exactly the same as for the transesterification of ethyl acetate

may have a major impact. On the other hand it might be that the

catalytic activity of the KOH catalyst in the homogeneous

process is rather low. Despite the uncertainties in the

explorative simulation results of the biodiesel production

process, they offer a useful starting point towards further

investigation of the industrial biodiesel production using MgO

catalysts. Moreover, these simulation results indicate thatheterogeneous catalysts may allow good production levels and

present an interesting alternative to homogeneous catalysts

towards a more environment-friendly process.

4. Conclusions

Magnesium oxide offers a viable heterogeneous solid base

catalyst for the standard transesterification of ethyl acetate or

triolein with methanol at industrial conditions for the industrial

production of fine-chemicals or biodiesel. Simulation of batch

and continuous slurry reactors, based on an Eley–Rideal kinetic

model, show that industrial production of fine-chemicals can bereached at ambient temperature and atmospheric pressure using

finely dispersed MgO catalyst. This allows the development of

a more environmentally friendly chemical process for the

transesterification reaction of low and high methyl acetate

production volumes since reaction neutralization is avoided and

product separation is made much easier. Also, the use of MgO

makes the batch process cost-effective since reactor heating is

not required.

Explorative process simulations for the continuous produc-

tion of biodiesel indicate that commercially attractive produc-

tion rates can be achieved. However, further experimental

investigation is needed to quantitatively determine the potential

of MgO as an alternative towards a more sustainabledevelopment of the industrial biodiesel production process.

Acknowledgement

This research has been funded in the frame of the STWW

project ‘‘HYPCAT’’ (STWW 141) of the IWT of the Flemish

Government.

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Table 11

Space time,yields, and conversions obtained at a production rate of 100,000 tonnes methyl oleate per year as a function of the methanol/triolein feed ratio and the use

of concentrations instead of activities in the rate equations (Eqs. (12a)–(12c)) instead of Eqs. (11a)–(11c))

T (K) M/T d p (mm) N I (s1) W / F tot (kg s/mol) Y MeOl Y G Y D Y M X T X MeOH

323 10 25 5 870.6 0.54270 0.05002 0.28953 0.59090 0.92938 0.16215

323 20 25 5 1013.5 0.62279 0.10319 0.18789 0.68123 0.97137 0.09300

323 10 25 5 851.6 0.52718a 0.04366a 0.30708a 0.56878a 0.91843a 0.15757a

a Simulations using Eqs. (12a)–(12c) with concentrations instead of activities.


13/13

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