Kinetics of Canola Oil Transesterification in a Membrane Reactor

Ind. Eng. Chem. Res. 2009, 48, 2533–2541

2533

Kinetics of Canola Oil Transesterification in a Membrane Reactor Peigang Cao, Andre´ Y. Tremblay,* and Marc A. Dube´ Department of Chemical and Biological Engineering, UniVersity of Ottawa, 161 Louis Pasteur Street, Ottawa, ON, K1N 6N5 Canada

Biodiesel is quickly gaining attention as a renewable and environmentally friendly replacement for petroleumbased diesel. The dominant process involved in its production is transesterification consisting of three consecutive reversible reactions. For this work, the transesterification of canola oil was conducted using a continuous membrane reactor in the presence of NaOH as a catalyst. The forward and reverse rate constants of all three steps involved in the transesterification in the membrane reactor are reported. The proposed mathematical model fitted the experimental results well. It was found that increasing the catalyst concentration increased the reaction rates and the residence time did not have a significant influence on the reaction rates. Runs were performed at 0.05, 0.1, and 0.5 wt % NaOH based on the oil. A mole ratio of 24:1 methanol/oil was used in this work. The forward rate constants were greater than previously reported for a batch process. This was attributed to the excellent mixing in the membrane reactor loop, the higher methanol/oil mole ratio used here, and the continuous removal of product from the reaction medium. The advantages of using a membrane rector to enhance the reaction rate in the transesterification of canola oil in a membrane reactor were clearly shown. 1. Introduction The necessity for alternative energy sources across the globe is increasing as nonrenewable resources are being depleted. It is for this reason that biodiesel has been the focus of many studies as an environmentally friendly and renewable replacement for petroleum-based diesel. Biodiesel is obtained by transesterifying lipids of vegetable or animal origin with alcohols. Triglycerides (TGs) react with alcohols in a reaction that is catalyzed by acid or base. This produces fatty acid alkyl esters (FAAEs) and glycerine as a coproduct. The characteristics and composition of the final biodiesel product vary with the type of lipid reacting with the alcohol.1 The transesterification reaction has a low rate because of the immiscibility of the two reactants, as oil and alcohol require vigorous stirring to promote good interphase contact.2 Boocock et al. proposed the use of tetrahydrofuran (THF) to overcome this difficulty,3 which reduced the reaction time to less than 15 min depending on the catalyst concentration. However, singlephase homogeneous solutions for biodiesel production will solubilize hydrophobic unreactable or unsaponifiable matter in the feedstock and retain it in the FAAE product. This will occur not only when converting low-quality lipids, but even when converting virgin oil, which can contain detectable amounts of unsaponifiable material.4 Other efforts such as using microwave or ultrasonic irradiation have also been made to accelerate the transesterification.5,6 The reversibility of transesterification also interferes with the completion of this process, which implies the presence of glyceride in the final biodiesel product. A membrane reactor technology has been proposed to produce high-quality biodiesel.7-9 The membrane reactor integrates reaction and separation in a single unit, provides continuous mixing of raw materials, and maintains high mass transfer between the immiscible phases. It was found that this novel reactor enabled the continuous separation of reaction products (FAAE/glycerol in alcohol) from the original lipid feed, which shifts the reversible transesterification to the product side. * To whom correspondence should be addressed. Tel.: +1 (613) 5625800ext. 6108. E-mail: [email protected].

Meanwhile, the straightforward separation of the FAAE phase in the permeate stream also enabled the recycling of the polar phase to the membrane reactor, thus lowering the overall alcohol-to-lipid molar ratio to 10:1.10 Some researchers have reported reaction kinetics for both acid- and base-catalyzed transesterification processes. At 32 °C, transesterification was 99% completed in 4 h when using base catalyst (NaOH), whereas at 60 °C, the reaction was completed in 1 h.11 The transesterification of soybean oil with methanol was reported to proceed with pseudo-first-order or second-order kinetics, whereas the reverse reaction was second-order.12 The methanolysis of sunflower oil at a methanol/oil molar ratio of 3:1 was reported to begin with second-order kinetics, but the rate decreased because of the formation of glycerol.13 Noureddini and Zhu studied the kinetics of transesterification of soybean oil.14 They used the same reaction model proposed by Freedman et al.12 and performed experiments at various mixing intensities, as measured by the Reynolds number. The reaction rates for the reverse of the first two reactions were found to be larger than the rates for the forward reactions. In addition, secondorder kinetics of palm oil transesterification were studied in batch mode.15 Vicente et al. conducted a comprehensive study on the kinetics of sunflower oil transesterification with a 6:1 methanol/oil molar ratio.16 The kinetic parameters of the homogeneous base-catalyzed methanolysis of soybean oil were determined by Doell et al.,17 who reported much larger rate constants than did Karmee et al. for the methanolysis of Pongamia oil.18 In Doell’s work, however, THF was added to the reaction mixture with a methanol/oil molar ratio of 27:1 to achieve and maintain a monophasic system throughout the reactions. In the work reported herein, we investigated the transesterification of canola oil with methanol using a membrane reactor. Methanol was the alcohol used in the reaction because of its low cost and high reactivity compared to longer-chain alcohols,19 and the FAAE produced was fatty acid methyl ester (FAME). The membrane reactor runs employed a continuous feed of oil and methanol, and the product was continuously removed from the reaction zone simultaneously. The present study was focused

10.1021/ie8009796 CCC: $40.75  2009 American Chemical Society Published on Web 01/28/2009

2534 Ind. Eng. Chem. Res., Vol. 48, No. 5, 2009

Figure 1. Schematic of the membrane reactor.

on the effects of residence time and catalyst concentration on the reaction rates. A modeling approach similar to the works of Freedman et al. and Vicente et al. was used, with the modification that the mass of permeate phase leaving the membrane reactor was identical to the combined mass of oil and methanol fed to the reactor. The six reaction rate constants in the model were estimated from measurements performed on the contents of the reactor (retentate) and the product stream (permeate). 2. Experimental Section 2.1. Materials. Food-grade canola oil (No-Name, Toronto, ON, Canada) was purchased at a local food store. Methanol (99.85% purity) was purchased from Commercial Alcohols Inc. (Brampton, ON, Canada). Tetrahydrofuran (THF, 99.98% purity, EMD Chemicals Inc., Gibbstown, NJ) and hydrochloric acid (HCl, 36.5-38%, Fisher Scientific Co., Nepean, ON, Canada) were both reagent grade. The THF was used only for HPLC analysis. Sodium hydroxide pellets (NaOH, reagent grade, ACP Chemicals Inc., Montreal, QC, Canada) were used as obtained. 2.2. Continuous Membrane Reactor. A schematic drawing of the membrane reactor is shown in Figure 1. The figure shows the feed system, the back-pressure valve and cooling system for permeate, the circulation pump, the membrane module, the heat-exchange system, and the safety and relief valve system. The circulating pump, membrane module, and heat exchanger make up a circulating loop within which the transesterification reaction occurs at a controlled temperature and pressure. A feed system and a back-pressure valve regulate the pressure in the loop. The feedstock oil and methanol/catalyst are pumped proportionately into this pressurized loop and circulated using the circulating pump. The multicomponent mixture flows on the retentate side of the membrane as a suspended mixture of lipids dispersed in a methanol-rich phase. The FAME/methanol/ glycerol phase permeates through the membrane and is collected for purification and/or recycling. The permeate stream is brought to atmospheric conditions by the back-pressure valve and cooler. The cooler reduces the temperature of the permeate, thus preventing the volatilization of methanol on release to atmospheric conditions. The retentate stream is heated by a heat exchanger, and the fluid is looped back to the feed side of the circulation pump. A thermal fluid is used in the heat exchanger.

The fluid is heated and circulated by a thermal bath (Neslab Instruments, Inc., Portsmouth, NH). In this study, a filtanium ceramic membrane (TAMI, Nyons, France) constructed of a titanium oxide support and active layer was employed. The overall internal volume of the membrane reactor is 6.0 L. 2.3. Experimental Procedures. For all runs, the methanol/ oil volume ratio fed to the reactor was 1:1. Four runs were carried out to compare different residence times: 15, 30, 45, and 60 min. The catalyst concentration for these four runs was held constant at 0.5 wt % NaOH by weight of oil. Another two runs were carried out to compare the lower catalyst concentrations of 0.1 and 0.05 wt % at a 60-min residence time. Initially, for all runs, 6 L of methanol and catalyst was charged to the reactor through the feed system and heated to 65 °C. At the same time, canola oil was preheated to 65 °C in the oil feed tank. With the circulation pump off, 3 L of canola oil was then quickly pumped into the reactor, displacing 3 L of methanol. The permeate back-pressure valve was shut initially, and the heat exchanger and circulating pump were turned on. The circulation pump was then switched on, and the feed system activated. Both the methanol/catalyst and oil feed rates were set according to the residence time. The canola oil and methanol/ catalyst solution were fed continuously to the pressurized loop. Almost immediately after the feed streams were fed to the reactor, permeate began to flow. Reactor pressure, temperature, and transmembrane pressure (TMP) were recorded continuously. Following each run, the membrane was backwashed with pure methanol, and the system was flushed for 30 min with 6 L of pure methanol and drained. Several repeat runs were performed using this equipment10,20 having a volume of 6 L and, in earlier studies, a smaller membrane reactor having an internal volume of 0.3 L.8,9 All quantities fed to the reactor were determined gravimetrically to (0.1 g. In Tremblay et al.,20 the error involved in feeding both methanol and oil to the membrane reactor in a 50:50 volume ratio, over the course of 3 h, was (2%. The variation in determining the composition of retentate and permeate for two to three repeat runs performed under identical conditions in our previous studies8-10 was (2%, and that for conversion was (3%. We do not further report on errors in this study as they were estimated to fall within the (2-3% range found in earlier studies.10,20

Ind. Eng. Chem. Res., Vol. 48, No. 5, 2009 2535

equations characterizing the stepwise reactions involved in the transesterification where the catalyzed reactions were first-order with respect to the catalyst concentration.16 In eqs 1-6, r denotes the reaction rates of the six components in units of (mol/L)/min rTG ) -(k1C + k10)[TG][Methanol] + (k2C + k20)[DG][FAME] )

d[TG] (1) dt

rDG ) (k1C + k10)[TG][Methanol] (k2C + k20)[DG][FAME] (k3C + k30)[DG][Methanol] + (k4C + k40)[MG][FAME] )

Figure 2. Reaction scheme of the transesterification.

2.4. Sampling and Analysis. Samples from the membrane reactor were taken at various time intervals. The samples were neutralized by the addition of 12 N hydrochloric acid to a pH of 7.0. Then, 0.4 g of sample was diluted with 20 g of THF in a 40-mL vial. The sample was then passed through a 0.2-µm polytetrafluoroethylene (PTFE) syringe filter and analyzed by gel-permeation chromatography (GPC), using THF as the mobile phase, according to the methods reported by Dube´ et al.21 The GPC system (Waters Corp.) consisted of a pump, a flow and temperature controller, a differential refractive index detector, and two 300 × 7.5 mm Phenogel columns of 3-µm and 100-Å pore size (Phenomenex, Torrance, CA) connected in series. The system was operated by Waters Millennium 32 software. HPLCgrade THF was used as the mobile phase at a flow rate of 0.05 mL/min at 25 °C. The sample injection loop was 200 µL, and the injected sample volume was 20 µL. The running time required for product characterization was approximately 60 min. Calibration curves were generated for the following standards (Sigma-Aldrich): triolein (TG), diolein (DG), monoolein (MG), methyl oleate (FAME), glycerol, and methanol. The areas under the peaks in the chromatograms for the product samples were used, together with the calibration curves, to determine the masses of the constituents (TG, DG, MG, FAME, glycerol, and methanol) present in the samples. 3. Kinetic Model for the Transesterification Process in a Membrane Reactor The transesterification of TG follows a three-step reaction as shown in Figure 2. These steps all occur at different rates of reaction and have their own rate constants (kn). In the first step, TG is converted to DG and a methyl ester, which, in turn, is converted to MG and a methyl ester in the second step. In the third and final step, MG is converted to glycerol and a methyl ester. In each step, one molecule of methyl ester is formed for every molecule of glyceride reacted. In this study, the shunt reaction involving the direct reaction of TG and methanol proposed by Freedman et al. to yield methyl ester and glycerol was not considered, because the proposed shunt reaction mechanism was not deemed necessary.12,14 For the above three-step reversible reaction, the forward and reverse reactions follow second-order overall kinetics.12,14 The reaction rate can be described as the sum of the rates from both uncatalyzed reactions and catalyzed reactions. Vicente et al. proposed a general form for the governing set of differential

d[DG] (2) dt

rMG ) (k3C + k30)[DG][Methanol] (k4C + k40)[MG][FAME] (k5C + k50)[MG][Methanol] + (k6C + k60)[G][FAME] )

d[MG] (3) dt

rFAME ) (k1C + k10)[TG][Methanol] (k2C + k20)[DG][FAME] + (k3C + k30)[DG][Methanol] (k4C + k40)[MG][FAME] + (k5C + k50)[MG][Methanol] (k6C + k60)[G][FAME] )

d[FAME] (4) dt

rG ) (k5C + k50)[MG][Methanol] (k6C + k60)[G][FAME] )

d[G] (5) dt

rMethanol ) -(k1C + k10)[TG][Methanol] + (k2C + k20)[DG][FAME] (k3C + k30)[DG][Methanol] + (k4C + k40)[MG][FAME] (k5C + k50)[MG][Methanol] + (k6C + k60)[G][FAME] ) d[Methanol] ) -rFAME (6) dt In eqs 1-6 above, k1, k3, and k5 are the forward rate constants for the catalyzed reactions; k2, k4, and k6 are the reverse rate constants for the catalyzed reactions; k10, k30, and k50 are the forward rate constants for the uncatalyzed reactions; and k20, k40, and k60 are the reverse rate constants for the uncatalyzed reactions. C is the catalyst concentration. The catalyst concentration in the reaction mixture remained constant because a catalyst/methanol solution with a constant catalyst concentration was fed steadily and the side reaction that consumed the catalyst could be neglected. Thus, we can use the effective rate constants as shown in eq 7 to simplify eqs 1-6. k′1 k′2 k′3 k′4 k′5 k′6

) ) ) ) ) )

k1C k2C k3C k4C k5C k6C

+ + + + + +

k10 k20 k30 k40 k50 k60

(7)


FTG-in d[TG] ) - k′1[TG][Methanol] + k′2[DG][FAME] dt V0 (11) Similarly, for DG in the membrane reactor

Figure 3. Schematic of the membrane reactor used for the biodiesel production kinetic model. *Based on the total volume of the emulsion in the reactor (6 L). **Concentration in the mobile phase ) concentration of the permeate.

A schematic representation of the transesterification process in a membrane reactor is illustrated in Figure 3. The TG and methanol were fed constantly to the membrane reactor at a known rate, whereas the mobile phase (methanol-rich phase) in the reactor could leave the reactor through the membrane as it was displaced by the incoming reactants. The circulation rate of the pump in the system allowed for the complete circulation of the reactor contents in 15 s, which permitted excellent mixing in the membrane reactor loop. According to the reaction mechanism shown in Figure 2, six components (namely, TG, DG, MG, FAME, glycerol, and methanol) are found in the reactor retentate side. Permeate crosses the membrane and is continuously removed from the reactor. TG molecules agglomerate as droplets and do not pass through the membrane pores unless the reaction mixture becomes a single phase.9 The circulating mass flow rate in the reactor was 20 kg/min. For the kinetic study, we assumed that the components on the retentate side of the membrane were perfectly mixed, so that the properties (e.g., concentration, temperature) of the reaction mixture were uniform in the circulating loop of the reactor, and the mass flow rate of the permeate (m ˙ Total-out) was also equal to the mass flow rate of the feedstock (m ˙ Total-in). Thus V˙out )

m ˙ Total-in FMethanol-inMWMethanol + FTG-inMWTG ) (8) Fout Fout

(9)

where ni is the number of moles of component i in the reactor. To provide for a less cumbersome nomenclature, the values of ni are abbreviated by the component name in the following equations. TG is assumed to be totally retained in the reactor; thus, its molar accumulation rate is d(TG) ) FTG-in + rTGV0 dt

(12)

FDG-out ) [DG]outV˙out

(13)

(FDG-in - [DG]outV˙out) d[DG] ) + k′1[TG][Methanol] dt V0 k′2[DG][FAME] - k′3[DG][Methanol] + k′4[MG][FAME] (14) For MG in the membrane reactor d(MG) ) rMGV0 - FMG-out dt

(15)

FMG-out ) [MG]outV˙out

(16)

(-[MG]outV˙out) d[MG] ) + k′3[DG][Methanol] dt V0 k′4[MG][FAME] - k′5[MG][Methanol] + k′6[G][FAME] (17) For FAME in the membrane reactor d(FAME) ) rFAMEV0 - FFAME-out dt

(18)

FFAME-out ) [FAME]outV˙out

(19)

(-[FAME]outV˙out) d[FAME] ) + k′1[TG][Methanol] dt V0 k′2[DG][FAME] + k′3[DG][Methanol] - k′4[MG][FAME] + k′5[MG][Methanol] - k′6[G][FAME] (20) For glycerol in the membrane reactor

A molar balance on component i in the reactor yields dni ) Fi-in + riV0 - Fi-out dt

d(DG) ) FDG-in + rDGV0 - FDG-out dt

(10)

When eq 10 is combined with eq 1, we obtain the TG concentration in the reactor as a function of time as

d(G) ) rGV0 - FG-out dt

(21)

FG-out ) [G]outV˙out

(22)

(-[G]outV˙out) d[G] ) + k′5[MG][Methanol] - k′6[G][FAME] dt V0 (23) For methanol in the membrane reactor d(Methanol) ) FMethanol-in + rMethanolV0 - FMethanol-out dt (24)

FMethanol-out ) [Methanol]outV˙out

(25)

(FMethanol-in - [Methanol]outV˙out) d[Methanol] ) dt V0 k′1[TG][Methanol] + k′2[DG][FAME] k′3[DG][Methanol] + k′4[MG][FAME] k′5[MG][Methanol] + k′6[G][FAME] (26) The model also assumes that the DG, MG, FAME, glycerol, and methanol form a single mobile or continuous phase in the reactor, whereas TG forms a separate phase in the reactor. The compounds in the mobile phase can pass through the membrane pores as a single phase and form the permeate. This implies that the concentrations of DG, MG, FAME, glycerol, and methanol in the permeate stream noted by the subscript “out” are equal to their corresponding mobile-phase concentrations in the reactor. The volume of the mobile phase (Vmobile) is Vmobile

[TG]V0MWTG ) V0 FTG

(27)

The component concentrations in the permeate can be simply expressed as [DG]out )

[MG]out )

[FAME]out )

[G]out )

[Methanol]out )

V0[DG] ) Vmobile

V0[MG] ) Vmobile

(

[DG] [TG]MWTG 1FTG

(

[MG] [TG]MWTG 1FTG

V0[FAME] ) Vmobile

V0[G] ) Vmobile

(

)

(

)

[FAME] [TG]MWTG 1FTG

[G] [TG]MWTG 1FTG

V0[Methanol] ) Vmobile

(

(28)

(29)

)

(30)

)

[Methanol] [TG]MWTG 1FTG

(31)

(


)

[DG]V˙out d[DG] )+ k′1[TG][Methanol] dt Vmobile k′2[DG][FAME] - k′3[DG][Methanol] + k′4[MG][FAME] (34)

(

)

[MG]V˙out d[MG] )+ k′3[DG][Methanol] dt Vmobile k′4[MG][FAME] - k′5[MG][Methanol] + k′6[G][FAME] (35)

(

( )

[G]V˙out d[G] )+ k′5[MG][Methanol] - k′6[G][FAME] dt Vmobile (37)

(

(32)

If we incorporate eqs 28-32 into eqs 11, 14, 17, 20, 23, and 26, a set of differential equations of component concentrations in the reactor as a function of time can be derived as shown in eqs 33-38. The six component concentrations in the reactor at a given time can be determined from the experimental data. The unknown parameters are the six effective rate constants. FTG-in d[TG] ) - k′1[TG][Methanol] + k′2[DG][FAME] dt V0 (33)

)

FMethanol-in [Methanol]V˙out d[Methanol] ) dt V0 Vmobile k′1[TG][Methanol] + k′2[DG][FAME] k′3[DG][Methanol] + k′4[MG][FAME] k′5[MG][Methanol] + k′6[G][FAME] (38) To calculate the effective rate constants, we need to integrate the eqs 33-38. A computer code for finite-difference methods was employed in MATLAB v6.5 (The MathWorks Inc., Natick, MA) to solve the governing set of differential equations for the entire experimental reaction time. The size of the time step (∆t) used in the simulation affected the results; however, a time step of less than 1 s gave similar results compared to a time step of 1 s.15 In this work, a time step of 0.01 min (0.6 s) was used in the simulation; the initial rate constant ranges were selected based on values from the literature.14 The relative average difference [S (%)] can be used as the objective criterion of the correctness of the model for base-catalyzed transesterification.18 We normalized the residuals by dividing them by the experimental values as follows15 S (%) )

)

)

[FAME]V˙out d[FAME] )+ dt Vmobile k′1[TG][Methanol] - k′2[DG][FAME] + k′3[DG][Methanol] - k′4[MG][FAME] + k′5[MG][Methanol] - k′6[G][FAME] (36)

[

m

t

i

0

∑∑

]

|yi(t)exp - yi(t)cal | × 100/N yi(t)exp

(39)

where m is the number of reactants, yi(t) is the concentration of component i at time t, subscript “exp” denotes the experimental value, subscript “cal” denotes the calculated value, and N is total number of measurements in the interval from time 0 to time t. Hence, S (%) indicates the average relative difference between all experimental and calculated concentrations of the components. 4. Results and Discussion 4.1. Calculation of the Effective Rate Constants. Six reactions were carried out at varying catalyst concentrations and residence times. The temperature was kept at 65 °C and the methanol/oil molar ratio was 24:1. The TMP across the membrane was 46, 42, and 32 kPa for the runs at 0.5, 0.1, and 0.05 wt % catalyst, respectively. These values were less than 7% of the maximum TMP (691 kPa) for this membrane.

2538 Ind. Eng. Chem. Res., Vol. 48, No. 5, 2009 Table 1. Calculated Rate Constants in the Membrane Reactor effective rate constants [L/(mol · min)]

catalyst conc (wt %)

residence time (min)

S (%)

k′1

k′2

k′3

k′4

k′5

k′6

0.5 0.5 0.5 0.5 0.1 0.05

60 45 30 15 60 60

14.83 12.47 11.73 14.48 17.71 13.97

0.1193 0.1475 0.1398 0.1096 0.0158 0.0026

0.7414 0.9227 0.8742 0.6857 0.6620 0.6450

1.4003 1.7748 1.6814 1.3190 0.4968 0.1492

3.6814 4.6744 4.4284 3.4739 2.0977 1.7979

0.3914 0.4815 0.4562 0.3586 0.2344 0.1365

0.0292 0.0349 0.0331 0.0260 0.0006 0.00034

Based on the experimental results, the effective rate constants (k′) were calculated by the procedure described in section 3. Table 1 lists all of the effective rate constants and the corresponding S (%) values. The parameters of the kinetic model were determined using an Excel spreadsheet; the model for the membrane reactor process was validated by making a mass balance with time for the reactor. The mass balance on the reactor was verified by using the calculated rate constants found in Table 1 to determine the mass of permeate leaving the reactor. The differences between the mass fed into the reactor with the mass leaving the reactor were always lower than 2%. This validates the model and the assumption that the volume of the continuous phase (methanol-rich phase) in the reactor can be determined by subtracting the oil phase from the total volume of the reactor. For the catalyst concentration of 0.5 wt % and a 1-h residence time reaction, the experimental and simulation results for the concentrations of the reaction mixture during the first 60 min of reaction time are shown in Figure 4. The transesterification mechanism consists of an initial masstransfer-controlled region followed by a kinetically controlled region,14 and the lag time at the initial stage occurs because of poor diffusion between methanol and oil. In this work, the concentration of FAME in the reactor increased rapidly and then leveled off. The TG conversion did not show a sigmoidal (S-shaped) pattern consisting of a low rate at the beginning followed by a sudden surge and finally a low rate again,12,22 even though the TG was continuous fed into the reaction zone. In Table 1, the residence time represents the feed flow rate to the reactor. The values for the effective rate constants, k′1-k′6, with the 0.5 wt % catalyst concentration above 15 min did not change significantly at different residence times. Under these conditions, the residence time does not have a significant impact on the kinetics of transesterification in the membrane reactor.

For the lower catalyst concentrations of 0.1 and 0.05 wt %, Figures 5 and 6 show both the experimental and simulated results for the concentrations of the reaction mixture during the first 20 min of the experiment. The concentrations of FAME and other components changed slowly with time. The concentrations of the components at 0.1 wt % took a longer time to reach steady state than did those at the 0.5 wt % catalyst concentration; the concentration of the components at 0.05 wt % still did not approach steady state when the run was carried out for 20 min. As shown in Table 1, the calculated values of the effective rate constants were also much smaller than those for the 0.5 wt % catalyst concentration, which implies that the reaction rates for catalyst concentrations of 0.1 and 0.05 wt % are lower. The lower reaction rates do not affect the purity of

Figure 4. Experimental points and simulation curves for the concentration of the reaction mixture during the transesterification in the membrane reactor with a catalyst concentration of 0.5 wt % for 1 h at residence time.




Figure 7. TG conversion as a function of time.

the product (FAME), as TG-free biodiesel can be obtained using a membrane reactor at low catalyst concentrations.22 The average value of S obtained in this work, as indicated in Table 1 above, was 14.2%. This is well under the value obtained by Komers et al.23 as 22.5% for the transesterification of rapeseed oil using KOH. The value of S indicates the average percent relative difference between all experimental and calculated concentrations of all components in this study. One main constraint in transesterifications using batch reactors is that the concentration of catalyst decreases over time. Catalyst concentration has a proportional effect on reaction rates. The rate constants in the case of a batch process cannot be assumed to remain constant for the duration of the run. The better results obtained by applying conventional kinetic models to the membrane reactor are most likely due to the constant feed of the methanol/catalyst solution to the reactor. This maintains a reasonably constant concentration of catalyst in the reactor. Thus, a better fit is obtained for the model because the rate constants do not vary significantly during the course of the run. The results of this study indicate that a membrane reactor is an excellent tool for studying transesterification reactions. 4.2. Effect of Catalyst Concentration. Figure 7 shows the effect of catalyst on the TG conversion. At the 0.05 wt % catalyst concentration, the TG conversion progressed slowly; the overall TG conversion was only 32.7% after 20 min. Masstransfer limitations resulted from the formation of an emulsion in the reaction mixture, but the membrane reactor can take advantage of this biphasic condition and produce biodiesel without TG. In Figure 7, the overall TG conversion for the 0.1 wt % catalyst run reached 95.5% after 20 min, and the TG concentration reached more than 98% after only 5 min for the 0.5 wt % catalyst concentration. According to the results of Vicente et al.,16 the effective rate constants increased linearly with the catalyst concentration. Figure 8 illustrates this linear relationship between the effective rate constants and the catalyst concentration, where the slopes represent the rate constants of the catalyzed reaction at 65 °C. These rate constants, irrespective of catalyst concentration, are listed in Table 2. As stated previously, the methanolysis process involves three steps, i.e., from TG to DG, from DG to MG, and from MG to glycerol. The ratio of kforward to kreverse for each step in the membrane reactor can be calculated, giving 1.48, 0.629, and 7.46, respectively. The ratio of the last step is greater than that of the second. This indicates that the reaction step from MG to glycerol is faster than the step from DG to MG,

Figure 8. Effective rate constants as a function of catalyst concentration: 0, k′1; ×, k′2; ∆, k′3; *, k′4; O, k′5; +, k′6. Table 2. Catalyzed Reaction Rate Constants [L/(mol · min)] at 65 °C in the Membrane Reactor catalyzed reaction rate constant

canola oila

sunflower oilb

Brassica carinata oilc

k1 k2 k3 k4 k5 k6

2.1991 1.4862 22.466 35.6950 4.1078 0.5503

0.6116 3.1346 2.0642 1.6055 2.0642 -

0.0510 0.3823 0.7785 1.0958 0.0510 -

a

This work. b Vicente et al.16 c Vicente et al.24

which is also in agreement with the work of Vicente et al.16 and the work of Komers et al.23 In a membrane reactor, the slowest step of DG to MG is the control step for the canola oil methanolysis; MG reacts with methanol at the highest rate and is the least stable component in the membrane reactor, which agrees satisfactorily with the MG concentration levels in this work and previous reported values.1,9,10


Compared to the work of Vicente et al.,16,24 the forward rate constants in the continuous membrane reactor were over 450% higher. This was attributed to the greater methanol/oil mole ratio used in this work (24:1 vs 6:1 used in Vicente et al.’s work), the excellent mixing in the membrane reactor loop, and the continuous removal of product from the reaction medium. This work shows the advantages of product removal in enhancing the reaction rate in the transesterification of canola oil in a membrane reactor. 5. Conclusions A kinetic model was proposed to simulate the methanolysis of canola oil in a membrane reactor, which gives a good fit to the experimental data for all of the conditions tested. It was found that mass-transfer effects were not significant in the membrane reactor at the methanol-to-oil molar ratio of 24:1. When the catalyst concentration was 0.5 wt %, the TG conversion reached above 98% after only 5 min. The results obtained show that a decrease in the catalyst concentration lowers the reaction rate; the calculated effective reaction rate constants increase linearly with the catalyst concentration. Furthermore, the second step of the methanolysis was deemed as the slowest process that controls the overall reaction rates, and MG was found to be the least stable component in the membrane reactor, in agreement with previously reported results. It was concluded that the residence time does not play an important role in the kinetics of canola oil in the membrane reactor, as the rate constants are at same level, but the continuous product removal from the reaction medium enhances the reaction rate in the transesterification of canola oil in the membrane reactor. A better fit to conventional kinetics models was obtained in this study. These models assume that the rate constants do not vary throughout the course of the run, which is not the case in batch transesterifications. The results of this study indicate that the membrane reactor is an excellent tool for studying transesterification reactions. Acknowledgment The authors thank the Natural Sciences and Engineering Research Council of Canada (NSERC) and the Ford Foundation for financial support. List of Symbols [DG] ) concentration of DG in the reactor based on the reactor volume, mol/L [DG]out ) concentration of DG in the permeate, mol/L [FAME] ) concentration of FAME in the reactor based on the reactor volume, mol/L [FAME]out ) concentration of FAME in the permeate, mol/L [G] ) concentration of glycerol in the reactor based on the reactor volume, mol/L [G]out ) concentration of glycerol in the permeate, mol/L [Methanol] ) concentration of the in membrane reactor based on the reactor volume, mol/L [Methanol]out ) concentration of methanol in the permeate, mol/L [MG] ) concentration of MG in the reactor based on the reactor volume, mol/L [MG]out ) concentration of MG in the permeate, mol/L [TG] ) concentration of TG in the reactor based on the reactor volume, mol/L FDG-in ) DG fed to reactor, mol/min; neglected for canola oil feed FDG-out ) DG outflow rate, mol/min

FFAME-out ) FAME outflow rate, mol/min FG-out ) glycerol outflow rate, mol/min FMethanol-in ) methanol feeding rate, mol/min FMethanol-out ) methanol outflow rate, mol/min FMG-out ) MG outflow rate, mol/min FTG-in ) TG feeding rate, mol/min ki ) rate constants of reaction i, L/(mol · min) ki′ ) effective rate constant of reaction i, L/(mol · min) m ˙ Total-in ) total mass flow rate of feedstocks entering the reactor, g/min m ˙ Total-out ) total mass flow rate of permeate, g/min MW ) molecular weight, g/mol r ) reaction rate, (mol/min)/L t ) time, min V0 ) volume of the membrane reactor, L V˙out ) volumetric flow rate of the permeate, L/min Fout ) density of permeate from the reactor, g/mL FTG ) density of TG in the reactor, g/mL

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ReceiVed for reView June 23, 2008 ReVised manuscript receiVed November 18, 2008 Accepted December 1, 2008 IE8009796

Kinetics of Canola Oil Transesterification in a Membrane Reactor

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