Transesterification of Canola Oil to Fatty Acid Methyl Ester (FAME) in a

TG conversions and DG, MG, and FAME yields in the packed bed reactor were calculated by subtracting the amount of reaction that occurred during the ti...
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Energy & Fuels 2008, 22, 3551–3556

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Transesterification of Canola Oil to Fatty Acid Methyl Ester (FAME) in a Continuous Flow Liquid-Liquid Packed Bed Reactor Fadi Ataya,† Marc A. Dube´,*,† and Marten Ternan†,‡ Department of Chemical and Biological Engineering, Centre for Catalysis Research and InnoVation, UniVersity of Ottawa, 161 Louis Pasteur Street, Ottawa, Ontario K1N 6N5, Canada, and EnPross, Inc., 147 Banning Road, Ottawa, Ontario K2L 1C5, Canada ReceiVed June 14, 2008. ReVised Manuscript ReceiVed July 28, 2008

Experiments were performed to study the mass-transfer limitations during the acid-catalyzed transesterification reaction of triglyceride (TG) with methanol (MeOH) to fatty acid methyl ester (FAME or biodiesel). The experiments were carried out as both agitated two- and single-phase reactions in an empty pipe and a packed bed reactor. The TG conversion rate increased with an increase in total superficial velocity and a decrease in packing particle diameter for the two-phase reactions. The TG conversion rate did not increase significantly with the reaction temperature for the two-phase reactions, whereas for the single-phase reactions, the TG conversion rate increased significantly as the reaction temperature increased. The rate constant at two-phase conditions (largest velocity, smallest packing particle size, and maximum pressure gradient) was comparable to that obtained at single-phase conditions, indicating that the mass-transfer limitations for two-phase experiments can be effectively overcome using a liquid-liquid packed bed reactor. The diminished mass transfer was explained by the formation of a new interfacial area between the two liquid phases, caused by the droplets being momentarily deformed into an elongated nonspherical shape as they passed through the openings between the solid particles of the packed bed.

Introduction Biodiesel, as defined by the American Society for Testing and Materials (ASTM), is a fuel comprised of monoalkyl esters of long-chain fatty acids derived from vegetable oils or animal fats. Biodiesel is a domestic, clean-burning, renewable, liquid fuel that can be used in compression-ignition engines instead of petroleum diesel, with little or no modifications to the engine components. Moreover, because biodiesel is plant-derived, its use significantly reduces greenhouse gas emissions over a full life cycle.1 The dominant biodiesel production process, transesterification, involves the reaction of alkyl alcohol with vegetable or animal oils in the presence of a catalyst (acid or base) to yield monoalkyl esters (biodiesel) and glycerol. Commercial-scale biodiesel production uses base-catalyzed technologies requiring highly refined virgin vegetable oils as raw materials. Conversely, acid-catalyzed processes allow for the use of lower cost feedstocks, such as waste cooking oils, yellow grease, and animal fats, that contain significant amounts of free fatty acid (FFA). These lower grade fats and oils are a less expensive source of biomass than pure vegetable oils and are a promising alternative to vegetable oils for biodiesel production.2-4 This is especially important from the point of view of avoiding the use of agricultural land for the production of fuel instead of food. * To whom correspondence should be addressed. Telephone: (613) 5625920. E-mail: [email protected]. † University of Ottawa. ‡ EnPross, Inc. (1) Ma, F.; Clements, L. D.; Hanna, M. A. Biodiesel production: A review. Bioresour. Technol. 1999, 70, 1. (2) Kulkarni, M. G.; Dalai, A. K. Waste cooking oilsAn economic source for biodiesel: A review. Ind. Eng. Chem. Res. 2006, 45, 2913. (3) Ma, F.; Clements, L. D.; Hanna, M. A. Biodiesel fuel from animal fat. Ancillary studies on transesterification of beef tallow. Ind. Eng. Chem. Res. 1998, 37, 3768.

The transesterification reaction consists of a number of consecutive reversible steps. When methanol (MeOH) is the alcohol, 1 mol of triglyceride (TG) is converted sequentially to diglyceride (DG), monoglyceride (MG), and finally, fatty acid methyl ester (FAME) and glycerol (GLY). The overall reaction sequence for the transesterification reaction is shown below: Overall Reaction: 1TG + 3MeOH T 1GLY + 3FAME Stepwise Reactions:

(1)

1TG + 1MeOH T 1DG + 1FAME

(2)

1DG + 1MeOH T 1MG + 1FAME

(3)

1MG + 1MeOH T 1GLY + 1FAME

(4)

The reaction system is essentially heterogeneous because the nonpolar TG and polar MeOH phases are immiscible.3 The reaction occurs at the MeOH-TG interface, and the process involves simultaneous mass transfer and chemical reaction. The result is a slower reaction rate. One approach to overcome this reduced reaction rate is the use of a solvent, such as tetrahydrofuran, to compatibilize both phases.5 Conversely, one could also seek to maximize the interfacial surface area between the two phases to improve the performance of the reaction.6-8 (4) Zhang, Y.; Dube´, M. A.; McLean, D. D.; Kates, M. Biodiesel production from waste cooking oil: 2. Economic assessment. Bioresour. Technol. 2003, 90, 229. (5) Boocock, D. G. B.; Konar, S. K.; Mao, V.; Sidi, H. Fast one-phase oil-rich processes for the preparation of vegetable oil methyl esters. Biomass Bioenergy 1996, 11, 43. (6) Ataya, F.; Dube´, M. A.; Ternan, M. Variables affecting the mass transfer limitations in the biodiesel reaction. Energy Fuels 2008, 22, 679. (7) Ataya, F.; Dube´, M. A.; Ternan, M. Acid-catalyzed transesterification of canola oil to biodiesel fuel under single- and two-phase reaction conditions. Energy Fuels 2007, 21, 2450.

10.1021/ef800462t CCC: $40.75  2008 American Chemical Society Published on Web 08/30/2008

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Figure 1. Overall schematic of the experimental system.

Effective mixing methodologies are expected to increase the TG mass-transfer coefficient and the interfacial surface area of contact.9 Ensuring surface renewal and the shredding of surfactant species from the interface is also expected to decrease the TG induction period that precedes the reaction and, thereby, improve the overall biodiesel reaction rate.9 Thus, one could consider exposing greater interfacial surface areas by extruding one phase into another, as the two phases flow through the openings between the particles, such as would be found in a packed bed reactor. In this paper, acid-catalyzed transesterification reactions were performed as both agitated two- and single-phase reactions in an empty pipe and, for the first time, in a continuous flow packed bed reactor. The effects of liquid velocity, particle diameter, and reaction temperature on the TG reaction rate were studied. Experimental Section Experiments were conducted in the experimental system depicted in Figure 1. The reactor length was L ) 914 mm, and the reaction internal diameter was i.d. ) 12.6 mm. The reactor bed was either kept empty or packed with glass beads (Potter’s Industries, Inc.) that were held in place by a 0.457 mm opening stainless-steel screen (ESPI). The external tubing line inner diameter was i.d. ) 9.53 mm. A Triplex Plunger Piston Pump (Giant Industries) providing a discharge pressure of 17.2 MPa and capable of delivering 107 mL/s was used. The pump was controlled by a variable frequency drive (VFD) (WEG Electric). The packed bed was heated to the reaction temperature using a heating tape (Omega), and the temperature was recorded using thermocouples (Omega) at the points of entry and exit, in addition to six points along the length of the column. There were two pressure transducers (Omega) to measure the pressure drop through the reactor. (8) Ataya, F.; Dube´, M. A.; Ternan, M. Single-phase and two-phase base-catalyzed transesterification of canola oil to fatty acid methyl esters at ambient conditions. Ind. Eng. Chem. Res. 2006, 45, 5411. (9) Hanson, C. Recent AdVances in Liquid-Liquid Extraction; Pergamon Press Ltd.: Oxford, U.K., 1971.

The feedstock canola oil was the “no-name” brand marketed by Loeb grocery stores. The reaction mixture had a MeOH (Commercial Alcohols, Inc.) to canola oil ratio of 30:1. Sulfuric acid catalyst, H2SO4 (BDH, Inc.), at 4.74 wt % by weight of the canola oil was used for all reactions. HPLC analysis of the feedstock canola oil showed that it initially contained TG, DG, and FFA. The mass ratios of DG/TG and FFA/TG in the feedstock were determined to be 0.0135 and 0.0152, respectively. Experiments were carried out as agitated two-phase reactions and agitated single-phase reactions in both an empty pipe and a packed bed reactor. For the singlephase reactions, the homogeneous medium was obtained by adding HPLC-grade THF (Sigma-Aldrich) to the reaction vials at a volumetric ratio of 1.5:1 THF/total sample. The reaction mixture was placed in a 4000 mL capacity feed tank, and the total system contained a liquid volume that was maintained at 2250 mL for all experiments. The empty reactor volume was 114 mL. The liquid not in the reactor was in either the system piping or the feed tank. The reactor was operated in top-feed continuous flow mode down through the packed bed and then recycled to the feed tank and back again to the top of the bed, thereby constituting a continuous flow batch recycle reactor. The overall reaction time, including the time both inside and outside of the packed bed portion of the reaction system, was 6 h for each of the batch experiments. The reactants spent some time inside the packed bed reactor and some time outside the packed bed reactor in the feed tank and the piping. The time that the reactants spent inside the packed bed reactor is defined as the packed bed residence time. TG conversions and DG, MG, and FAME yields in the packed bed reactor were calculated by subtracting the amount of reaction that occurred during the time spent outside the packed bed reactor from the amount of reaction that occurred during the total elapsed reaction time. During the reaction, samples were taken periodically, subsequently allowed to phase-separate, and then placed in an iced water bath for a period of 24 h. A lower polar phase and an upper nonpolar phase were formed. Each of the polar and nonpolar phases was titrated using Ca(OH)2. The titration end point was determined using litmus paper. Both of the phases were then filtered through 0.2 µm polytetrafluoroethylene (PTFE) syringe filters and analyzed using

Transesterification of Canola Oil to FAME

Figure 2. Fractional TG conversion to all products versus packed bed residence time, for the two-phase reactions, at the total superficial velocities of 0.160 m/s (2), 0.481 m/s (0), and 0.801 m/s ([) at 25 °C. For all figures, dashed fitted lines represent experiments performed at 25 °C and solid fitted lines represent experiments performed at 35 °C.

HPLC according to the methods reported by Dube´ et al.10 A total of 0.04 g of sample was weighed into the HPLC vials and diluted with THF to make up 20 mg/mL sample solutions for HPLC analysis. The HPLC (Waters Corporation) consisted of a pump, a flow and temperature controller, a differential refractive index detector, and two 300 × 7.5 mm PLgel columns of 3 µm and 100 Å pore size (Varian, Inc.) connected in series. The system was operated using Waters Millennium 32 software. HPLC-grade THF was used as the mobile phase at a flow rate of 0.05 mL/min at 38 °C. The sample injection loop was 200 µL, and the injected sample volume was 20 µL. Calibration curves10 were generated for the standards (Sigma-Aldrich): triolein (TG), diolein (DG), monoolein (MG), methyl oleate (FAME), oleic acid (FFA), and glycerol (GLY). The areas under the peaks in the chromatograms for the product samples were used together with the calibration curves, to determine the moles of the constituents (TG, DG, MG, FFA, and FAME) present in the nonpolar phase of each sample. The amounts of product constituents were subjected to further computations to provide the conversions and yields of the components used in the analysis. The number of moles of TG in the feedstock (that corresponded to the measured amount of product) was calculated from the measured number of moles of the product constituents plus the stoichiometry of the reactions. The moles of DG and FFA in the feed were calculated from the moles of TG in the feed, the DG/TG feed ratio, and the FFA/TG feed ratio, respectively. The moles of MeOH in the feed was 6(moles TG + moles DG) in the feed. The moles of “MeOH in the product” and “glycerol in the product” were calculated using the moles of the product constituents plus stoichiometry of the reactions. With the above information on feed and product constituents, the conversions and yields were calculated. Standard IUPAC definitions were used for conversion (moles of feed converted/mole of feed) and yield (moles of feed converted to a product constituent/mole of feed).

Results and Discussion The reaction rate occurring outside the reactor (in the feed tank and piping) was identical to that occurring in the empty pipe reactor (or no packing experiment) at ambient temperature. Consequently, corrections were made to differentiate between the reactions occurring inside the packed bed reactor volume alone versus the remaining volume of the reaction system (feed tank plus piping). Results are shown in Figure 2 for the conversion of TG as a function of the packed bed residence time at different total (10) Dube´, M. A.; Zheng, S.; McLean, D. D.; Kates, M. A comparison of attenuated total reflectance-FTIR spectroscopy and GPC for monitoring biodiesel production. J. Am. Oil Chem. Soc. 2004, 81, 599.

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Figure 3. Fractional TG conversion to all products versus packed bed residence time, for the two-phase reactions, at the packing particle diameters of 1 mm ([), 2 mm (0), and 5 mm (2) at 25 °C.

superficial velocities for the two-phase transesterification of canola oil to FAME. By comparing the TG conversion profiles in Figure 2, one can observe that an increase in the total superficial velocity improved the rate of TG conversion. The pressure drop across the packed bed increased as the total superficial velocity, Vs, increased. The results in Figure 2 are consistent with other reports in the literature. The increase in the pressure drop across the packed bed is consistent with the prediction based on the Ergun equation.11 Merchuk et al.12 studied the efficiency of copper recovery from lean aqueous solutions with an organic solution of LIX-64N solvent in kerosene using several motionless mixers, packed beds, and empty pipes. Their results showed that the devices that proved to be the most efficient for copper recovery were the ones that provided the largest pressure drop. The main factor affecting effectiveness in the process was turbulence, which relates to energy consumption. The pressure drop provided the energy required to produce the turbulence in the system and created the interfacial area necessary for mass transfer. Leacock and Churchill13 measured the rate of mass transfer between isobutanol and water. The values obtained showed an increase of mean mass-transfer coefficients with both the isobutanol and water flow rates. Verma and Sharma14 investigated the mass-transfer characteristics of packed liquidliquid extraction columns with a variety of packing sizes at different superficial velocities of the dispersed and continuous phases. Data were also obtained for the co-current up-flow in spray columns, i.e., without packing. The values of kca and a for the spray columns were lower than those obtained for the stainless-steel packing under otherwise similar conditions. Therefore, the rate of mass transfer increases as the effective interfacial area and the volumetric mass-transfer coefficient increase. It can be postulated that the increase in the overall rate of TG conversion resulted from the increased rate of mass transfer caused by the increase in the total superficial velocity, Vs. Our results are consistent with the above findings. In Figure 3, results for the conversion of TG as a function of the packed bed residence time at the different packing particle diameters are shown. By comparing the TG conversion profiles, (11) de Nevers, N. Fluid Mechanic for Chemical Engineers, 2nd ed.; McGraw-Hill, Inc.: New York, 1991. (12) Merchuk, J. C.; Shai, R.; Wolf, D. Experimental study of copper extraction with LIX-64N by means of motionless mixers. Ind. Eng. Chem. Process. Des. 1980, 19, 91. (13) Leacock, J. A.; Churchill, S. T. Mass transfer between isobutanol and water in concurrent flow through a packed column. AIChE J. 1961, 7, 196. (14) Verma, R. P.; Sharma, M. M. Mass transfer in liquid-liquid extraction columns. Chem. Eng. Sci. 1975, 30, 279.

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Figure 4. Fractional TG conversion to all products versus packed bed residence time, for the two-phase reactions, at the temperatures of 25 °C (0) and 35 °C (9).

it can be seen that a decrease in packing particle diameter improved the rate of TG conversion. The pressure drop across the packed bed also increased as the packing particle diameters decreased. The Sauter mean diameter, d32, is the diameter of a sphere that has the same volume to surface area ratio as that of a dispersed phase droplet of interest. There are reports in the literature that suggest our results on the effect of packing particle diameter, in Figure 3, may be related to the d32 of dispersed phase droplets in our packed bed. Rigg and Churchill15 studied the behavior of immiscible liquids in the co-current flow through packed beds. Their results showed that an increase in the total superficial velocity of both phases resulted in a decrease in d32 for the range of packed bed particle diameters investigated. In addition, they showed that d32 decreased, for a specified total superficial velocity, with a decrease in the packing diameter. Duffy and Kadlec16 investigated turbulent liquid-liquid up-flow in packed beds. They showed that d32 decreased as the flow rate of the continuous phase increased and the particle diameter decreased. Given that the interfacial area increases as d32 decreases, one would expect that the rate of mass transfer should also increase. Thus, our observed increases in the overall rate of TG conversion probably resulted from the increased rate of mass transfer caused by the decrease in the packing particle diameter, dp. Results are shown in Figure 4 for the conversion of TG as a function of the packed bed residence time at different temperatures for the two-phase transesterification of canola oil to FAME. The results indicate that increasing the reaction temperature did not significantly improve the rate of TG conversion. These results are consistent with the work of Noureddini and Zhu,17 who examined the combined effects of mixing and kinetics on the transesterification of soybean oil and MeOH. The conversion of TG was measured as a function of time for the homogeneous single-phase transesterification of canola oil to biodiesel (see Figure 5). The effect of temperature can be seen by comparing the TG conversion profiles for reactions carried out at 25 and 35 °C. In contrast to the two-phase reactions, the results indicate that increasing the reaction temperature significantly improved the rate of TG conversion (compare Figures 4 and 5). (15) Rigg, R. G.; Churchill, S. W. The behavior of immiscible liquids in concurrent flow through packed beds. AIChE J. 1964, 10, 810. (16) Duffy, J. P.; Kadlec, R. H. The behavior of the dispersed phase in liquid-liquid concurrent flow through a packed bed. Can. J. Chem. Eng. 1975, 53, 621. (17) Noureddini, H.; Zhu, D. Kinetics of transesterification of soybean oil. J. Am. Oil Chem. Soc. 1997, 74, 1457.

Ataya et al.

Figure 5. Fractional TG conversion to all products versus total elapsed time, for the single-phase reactions, at the temperatures of 25 °C (0) and 35 °C (9).

Figure 6. ln(1/(1 - XTG)) versus total elapsed time, for the singlephase reactions, at the temperatures of 25 °C (0) and 35 °C (9).

The increase in TG conversion seen by comparing Figures 2-4 with those in Figure 5 was obtained by adding THF, which caused the two-phase reaction mixture to become a single-phase reaction mixture. The homogeneous single-phase reaction conditions resulted in the elimination of the interface between phases, and therefore, there can be no limitations because of interphase mass transfer. These observations are consistent with the idea that mass transfer influences the rate of reaction when an interface is present.5 Figure 6 is a plot of the ln(1/(1 - X)) as a function of the total elapsed time for the homogeneous single-phase reactions performed at 25 and 35 °C, respectively. For a first-order reaction, a plot of ln(1/(1 - X)) as a function of time will be linear, with a slope equal to the reaction rate constant, k. The reaction rate constants were determined to be 0.0429 h-1 or 1.19 × 10-5 s-1 at 25 °C and 0.0824 h-1 or 2.29 × 10-5 s-1 at 35 °C. The reaction rate constant obtained at 35 °C is approximately 2 times greater than that obtained at 25 °C. The experimental results are consistent with the notion that the rate of reaction doubles for a 10 °C increase in the reaction temperature.18 Previously, we reported a reaction rate constant for our batch experiments carried out under acid-catalyzed conditions at 20 °C.7 The reaction rate constant evaluated for the TG-DG reaction was 0.56 × 10-5 s-1 at 20 °C. The rate constant value obtained during our previous stationary batch experiments is directionally consistent with the rate constant values obtained during these continuous flow recycle batch experiments. (18) Fogler, S. H. Elements of Chemical Reaction Engineering, 3rd ed.; Prentice Hall: Upper Saddle River, NJ, 1999; p 72.

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Figure 7. Reaction rate constants for the fractional TG conversion versus 1/T × 103 for the single-phase reactions.

Figure 7 is a plot of the first-order reaction rate constants, k (for TG fractional conversion), versus 1/T for the agitated singlephase reactions. The single-phase reaction rate constants obtained from Figure 6 at 25 °C (0.0429 h-1 or 1.19 × 10-5 s-1) and 35 °C (0.0824 h-1 or 2.29 × 10-5 s-1), in addition to the reaction rate constant obtained from our previous study at 20 °C (0.0200 h-1 or 5.55 × 10-6 s-1), are plotted against the reciprocals of their respective reaction temperatures, and the slope was determined. The dependence of the reaction rate constant on temperature is given by the logarithmic form of the Arrhenius equation (eq 5), where A is the pre-exponential factor, E is the activation energy in J/mol or cal/mol, R is the gas constant in J mol-1 K-1 or cal mol-1 K-1, and T is the absolute temperature in Kelvin. The slope of the plot of log k versus 1/T can be determined, and the activation energy can be evaluated. (5)

Figure 8. ln(1/(1 - XTG)) versus packed bed residence time, for the two-phase reactions performed at Vs ) 0.481 m/s, dp ) 1 mm, and T ) 35 °C.

The activation energy with the H2SO4 catalyst, under singlephase conditions (see Figure 7), was determined from the slope to be E ) 63.9 kJ/mol. For the sake of comparison, Freedman et al.19 carried out experiments under acidic conditions at a BuOH/oil molar ratio of 30:1 for the acid-catalyzed butanolysis of TG to DG in the temperature range of 77-117 °C. One should note that a single-phase regime would be expected for the butanolysis reaction. Their calculated activation energy was evaluated as 1.49 × 104 cal/mol or 62.7 kJ/mol, which is similar to the single-phase value determined from our study. Previously, we reported rate constants for the transesterification of TG with MeOH using a NaOH basic catalyst.8 In those experiments, the activation energy with 1 wt % NaOH catalyst, under single-phase conditions, was determined to be 28.0 kJ/ mol. The fact that the activation energy with the H2SO4 catalyst reported in this study is greater than that for the 1 wt % NaOH catalyst single-phase experiments is consistent with transesterification being slower with acidic catalysts than with basic catalysts. Activation energy values for the TG-DG reaction

reported by others for two-phase experiments under basic conditions17,19,20 were larger (∼60 kJ/mol) than for our singlephase value of 28.0 kJ/mol. The smaller activation energy values for two-phase experiments could be explained by the two-phase experiments having both a mass-transfer resistance and a reaction kinetics resistance. In contrast, single-phase experiments would be expected to have no mass-transfer resistance. Figure 8 is a plot of ln(1/(1 - X)) as a function of the packed bed residence time for the heterogeneous two-phase reaction performed at a total superficial velocity, Vs ) 0.481 m/s, a packing particle diameter, dp ) 1 mm, and a reaction temperature, T ) 35 °C. The reaction conditions studied for this experiment resulted in the largest pressure drop across the packed bed reactor (5.2 MPa), which we will refer to as the maximum pressure gradient conditions (5.2 MPa/0.919 m ) 5.6 MPa/m) among our two-phase reactions. The TG conversion profile for this reaction was shown in Figure 3. For this case, k was determined to be 0.209 h-1 or 5.81 × 10-5 s-1. One observes that the apparent reaction rate constant obtained in

(19) Freedman, B.; Butterfield, R. O.; Pryde, E. H. Transesterification kinetics of soybean oil. J. Am. Oil Chem. Soc. 1986, 63, 1375.

(20) Darnoko, D.; Cheryan, M. Kinetics of palm oil transesterification in a batch reactor. J. Am. Oil Chem. Soc. 2000, 77, 1269.

E 1 log k ) log A 2.3R T

()

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Figure 8 for the two-phase conditions (5.81 × 10-5 s1) is larger than that obtained in Figure 6 from the single-phase conditions (1.19 × 10-5 s-1) at the same temperature. However, the concentrations of H2SO4 in both sets of experiments were not the same, and this necessitated the evaluation of effective reaction rate constants, keffective, from the apparent reaction rate constants, k. The TG conversion reaction is first-order in the TG concentration, [TG]. It is also first-order in the proton concentration or sulfuric acid concentration, [H2SO4].21 rate ) keffective[TG][H2SO4]

(6)

When the [H2SO4] is constant, eq 6 becomes rate ) k[TG] (7) The acid catalyst concentration was maintained at 4.74 wt % by weight of the canola oil for all experiments. However, the amount of TG added to the two-phase experiments was larger than the amount of TG added to the single-phase experiments. The reason was that the total reaction volume was maintained at ∼2250 mL for all experiments, and the volume necessary to achieve the homogenized medium was 1.5:1 THF/total sample. Therefore, the amount of TG added for the single-phase experiments was smaller than that in the two-phase experiments, and the concentrations of H2SO4 based on the wt % of the oil were approximately 2.6 times smaller in the single-phase experiments. This resulted in the differences observed in our data. The effective reaction rate constant values, corrected for the H2SO4 concentrations, were calculated. They showed that the value corresponding to the maximum pressure gradient twophase reaction condition approached that of the single-phase condition. Specifically, the keffective values for the single-phase reaction and two-phase maximum pressure gradient reactions were 0.297 and 0.285 mL mol-1 s-1, respectively. Therefore, comparable kinetic rate constants for TG conversion were obtained at both a single-phase reaction condition (no masstransfer limitations) and one two-phase reaction condition (superficial velocity ) 0.481 m/s, packing particle diameter ) 1 mm, and pressure drop across the packed bed ) 5.17 MPa). This is the first time that an overall two-phase transesterification rate constant has been reported to be essentially the same as the corresponding single-phase rate constant. It suggests that most of the mass-transfer resistance may have been removed at our maximum pressure gradient condition. The diminished mass-transfer resistance may be related to deformation of the dispersed phase droplets as they flow through the openings between the spherical solid particles in the packed bed. If the droplets are larger than the openings, they will temporarily take on an elongated nonspherical shape (perhaps a dumbbell shape22), when part of the droplet has finished flowing through the opening and part of the droplet has yet to enter the opening. Regardless of the shape, a droplet that is larger than the openings between the spheres that form the packed bed will be deformed (21) Vicente, G.; Martinez, M.; Aracil, J.; Esteban, A. Kinetics of sunflower oil methanolysis. Ind. Eng. Chem. Res. 2005, 44, 5447. (22) Batey, W.; Thornton, J. D. Partial mass-transfer coefficients and packing performance in liquid-liquid extraction. Ind. Eng. Chem. Res. 1989, 28, 1096.

and some of the fluid initially within the droplet interior will be forced to the droplet interface. New interfacial area will temporarily be created. The new interfacial area will be composed of fluid from the bulk of the droplet that does not have the concentration gradient that would normally be present at the interface between the dispersed droplets and the continuous liquid phase. Because the continuous liquid phase will also experience turbulent conditions as it flows through the openings, it is possible that some bulk concentration dispersed phase liquid may temporarily be in contact with some bulk concentration continuous phase liquid. If such a phenomenon occurred sufficiently frequently, then perhaps the two-phase reaction rate might become equivalent to a single-phase reaction rate. In some respects, this droplet shape deformation phenomenon is different from conventional mixing phenomena. Conclusions Canola oil transesterification experiments were performed as both two- and single-phase reactions in an empty pipe and a packed bed reactor. TG conversion was shown to increase with increasing superficial velocity and decreasing diameter of particles in the packed bed. The heterogeneous two-phase and homogeneous single-phase experiments were compared on the basis of their effective reaction rate constants. The rate constants at the maximum pressure gradient two-phase conditions were similar to those performed under single-phase conditions. The results indicated that the mass-transfer limitations for heterogeneous two-phase experiments can be effectively overcome using a liquid-liquid packed bed reactor. Acknowledgment. The authors acknowledge the BIOCAP Canada Foundation and the Natural Sciences and Engineering Research Council (NSERC) of Canada for financial support of this research.

Nomenclature FAME ) fatty acid methyl ester FFA ) free fatty acid MG ) monoglyceride DG ) diglyceride TG ) triglyceride THF ) tetrahydrofuran GLY ) glycerol MeOH ) methanol BuOH ) butanol H2O ) water H2SO4 ) sulfuric acid NaOH ) sodium hydroxide Ca(OH)2 ) calcium hydroxide X ) conversion of species R ) reaction rate of species k ) apparent reaction rate constant of species keffective ) effective reaction rate constant of species A ) pre-exponential factor E ) activation energy R ) gas constant T ) temperature EF800462T