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May 13, 2013 - Sumaiya Zainal Abidin , Nursofia Mohd Yunus , Siti Amirah Abdul Ghani ... Mohamed Kheireddine Aroua , Ahmad Shamiri , Mahtab Samadi...
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Esterification of Palm Fatty Acid Distillate Using Heterogeneous Sulfonated Microcrystalline Cellulose Catalyst and Its Comparison with H2SO4 Catalyzed Reaction Deepak D. Chabukswar, Parminder Kaur K. S. Heer, and Vilas G. Gaikar* Department of Chemical Engineering, Institute of Chemical Technology, Matunga, Mumbai-400 019, India S Supporting Information *

ABSTRACT: The kinetics of esterification of palm oil fatty acid distillate (PFAD) with methanol was investigated using heterogeneous carbonized microcellulose sulfonic acid as catalyst and compared with the sulfuric acid catalyzed reaction, considering liquid−liquid phase split during the progress of the reaction in both the cases. The solid catalyst was characterized for acidity, thermal stability, and surface area. The residual glycerides in the PFAD were hydrolyzed prior to the esterification using sulfuric acid as a catalyst. The esterification reaction was investigated for the effect of catalyst loading, temperature, and free fatty acids (FFA) to methanol ratio, on the conversion of the fatty acids. Sulfuric acid was a better catalyst than sulfonated microcrystalline cellulose, but the solid catalyst provides the ease of recovery. The sulfonated microcrystalline cellulose rendered good conversion and reusability for esterification. The process engineering aspects of the esterification reaction are also briefly discussed.

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INTRODUCTION Biodiesel, a mixture of fatty acid methyl esters (FAMEs), has attracted considerable attention as an alternative to petro-based transport fuels. The renewable nature of the vegetable oils feedstock and the lower sulfur value of the biodiesel resulting in the reduction of vehicular SOX emission and net carbon dioxide emissions are responsible for this increasing interest.1−4 The major feedstock for biodiesel comes from edible oils such as soyabean oil, palm oil, and nonedible oils such as jatropha and karanja oils. The production of biodiesel by alkali catalyzed transesterification of neat vegetable oils, which by far is the simplest method of preparing the fatty acid esters, has not been commercialized to a large extent because of the higher cost of raw materials and also because of the limited availability of the feedstock in developing countries despite bringing considerable land under cultivation for nonedible oil producing plants.5 The biodiesel manufacturing process has been under intensive study in the past decade for development of new technologies to enable the use of low cost waste oils as the feedstock as compared to more expensive neat vegetable oils. The biodiesel production from waste oils from different sources is a challenging job because of the presence of free fatty acids (FFAs) in such oils to varying degrees.6−8 The FFAs react with the alkali that is used as a catalyst in the transesterification process, to form soap that hinders subsequent phase separation of the methyl esters from glycerol formed in the reaction. The effectiveness of the alkali catalyst is significantly reduced in the presence of moisture generated by the neutralization reaction that reduces the overall conversion in a desired time frame and thus demands larger amounts of the catalyst in the process. The fatty acid salts also dissolve in the glycerol phase making subsequent glycerol recovery cumbersome. Completely dry conditions are the most essential to produce the fatty acid methyl esters by the transesterification of triglycerides in the shortest possible time as the reaction is extremely fast in the © XXXX American Chemical Society

absence of moisture and gets limited only by the mixing conditions. The conventional alkali catalyzed transesterification process, therefore, becomes inadequate to use cheaper raw materials having significant FFA content.8 The pretreatment of the waste oil to esterify the FFAs by acid catalyzed esterification becomes a prerequisite in the biodiesel production by conventional alkali catalyzed transesterification process. On the other hand, feed stocks containing of mostly FFAs with a limited content of triglycerides can be handled by direct esterification process. Palm fatty acid distillate (PFAD) is a byproduct of the palm oil refining process and mainly consists of a mixture of fatty acids with a small percentage of mono-, di-, and triglycerides and is comparatively far cheaper than palm oil as a starting material for methyl esters production. Biodiesel synthesis, by direct esterification of waste oils with methanol, has been reported by many researchers using homogeneous and heterogeneous acid catalysts.9−27 Homogeneous mineral acids such as H2SO4, phosphoric acid (H3PO4), and organic acids like p-toluene sulfonic acid (pTSA), trichloroacetic acid (TCA), and methanesulfonic acid (MSA) are routinely used as catalysts for the esterification reactions.10−14 Since the recovery or post-treatment of the catalyst remains the main concern in such homogeneous catalyzed reactions, low cost sulfuric acid remains still the catalyst of choice despite its corrosive nature, as it also can be easily neutralized. The organic acids after neutralization, on the other hand, add a significant load to the chemical oxygen demand (COD) of the waste streams that is not easy to deal with even by biological means. Received: November 10, 2012 Revised: April 4, 2013 Accepted: May 13, 2013

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solid with a high density of active sites, allowing highperformance catalysts to be prepared from natural products. The esterification of PFAD, with CSMC, is compared with that using Amberlyst-15 under similar conditions.

Heterogeneous catalysts offer major advantages of facile separation from the reaction mixture and reusability and substantial benefits from an environmental pollution point of view. Inorganic solid acids such as niobic acid, silica and zeolite supported Lewis acids, zirconium sulfate, and super acid catalysts have been extensively studied for the esterification reaction.16−20 Lipase catalyzed enzyme esterification has been also reported, but the reaction rates are much slower as the temperature and pressure conditions are too mild, usually not exceeding 40−45 °C as enzyme denaturation leads to the loss of its activity at higher temperatures. This is a serious drawback of the enzymatic reaction in industrial conditions because of poor volumetric productivity although enzymes can work very well in the presence of water formed in the reversible esterification reaction. Acidic ion-exchange resin catalysts, such as CT-175 cation exchage resin, Dowex monosphere 550A, Amberlyst-15, and Nafion NR50, in H+ form, have all been reported, sometimes even with supercritical methanol conditions to take advantage of higher rates of the reaction at higher operating temperatures.20−27 Many of these resin catalysts, however, retain the water of esterification that often decreases the reaction rates and conversions. The prior literature indicates that cheaper feedstock, driving reaction to completion, easily available and recyclable catalyst, and moderate reaction times are the major concerns of the biodiesel manufacturing process. Kinetic modeling of esterification of fatty acids using homogeneous and heterogeneous catalysts also has been performed by several authors.28−35 Most of these papers refer to a homogeneous liquid phase reaction using a large excess of methanol and, therefore, use of a pseudohomogeneous power law model or Eley−Rideal or Langmuir−Hinshelwood model to describe the kinetics of homogeneous and heterogeneous catalyzed reactions, respectively, is very common. However, generation of water during the esterification reaction and limited miscibility of methyl esters with methanol lead to a biphasic reaction medium in both of these cases. The two liquid phases attain equilibrium with respect to each other very quickly. Therefore, it becomes necessary to understand the thermodynamic framework of the reaction system, and the effect of phase equilibrium between the two liquid phases on the reaction kinetics of the esterification. Very few papers have analyzed the kinetics of the esterification considering the biphasic nature of the system.36,37 In this paper, PFAD is used for the production of FAMEs by acid catalyzed esterification, after hydrolysis of residual glycerides in the feed. The esterification reaction is conducted with concentrated sulfuric acid as the homogeneous catalyst for comparison with the esterification using a carbonized sulfonated microcrystalline cellulose solid acid (CSMC) as a heterogeneous catalyst. Sulfonation and incomplete carbonization of natural products results in a rigid carbon framework composed of small polycyclic carbon sheets in a threedimensional sp3-bonded structure.38−42 Microcrystalline cellulose is a very strong organic material consisting of polysaccharide linear chains of β (1 → 4) linked D-glucose units. Cellulose, as such, has hydrophobic surface and a high amount of amorphous regions. Microcrystalline cellulose material has a lower content of amorphous regions and a higher degree of crystallinity and is usually obtained by partially hydrolyzing cellulose with a mineral acid or by steam explosion process. Sulfonation of such carbonaceous materials, such as microcrystalline cellulose, is expected to afford a highly stable

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EXPERIMENTAL SECTION Materials. The palm fatty acid distillate (PFAD) was procured from Royal Energy Ltd., Mumbai, as a yellow color solid having melting point of 40 °C. Methanol (AR grade), sulfuric acid (98%), and Amberlyst-15 were procured from SD, Fine Chemicals, Mumbai. A crystalline powder sample of microcrystalline cellulose (MC; surface area = 105.1 m2/g, pore size = 1.3 nm) was supplied by Godavari Biorefineries Ltd., Mumbai. The HPLC grade acetonitrile and acetone were procured from Thermo Fischer Scientific, Mumbai, for the analysis. Catalyst Preparation. In a typical run, MC (10 g) was taken along with dichloromethane (50 cm3) in a reaction vessel and chlorosulfonic acid (5 cm3) was added dropwise into the reaction vessel over a period of 2 h, under continuous stirring conditions at 0 °C in an ice bath. The stirring was continued until evolution of hydrochloric acid fumes stopped completely. The reaction mixture was then filtered and the solid was washed thoroughly with methanol and then dried at 40 °C for 3 h in an oven. Amberlyst-15 catalyst was treated with H2SO4 solutions (0.1 mol/dm3) and then washed with deionized water up to neutral pH to remove free ions and then dried in an oven at 120 °C for 5 h before use. Catalyst Characterization. The acidity of the catalyst was determined by ammonia temperature programmed desorption (NH3-TPD). The catalyst (0.2 g) was charged into a U-shaped quartz sample tube and was heated at 180 °C under the flow of helium for degasification. After cooling to ambient temperature of 30 °C, ammonia was injected into the sample to saturate it. The thermal desorption of ammonia was recorded under the flow of helium and by raising the temperature at a rate of 5 °C/ min up to 180 °C. The thermal stability of the catalyst was checked by differential scanning calorimetry and thermal gravimetric analysis (DSC/TGA). The catalyst (2 mg) was heated in the sample compartment in the temperature range of 40−500 °C at a constant rate of 5 °C/min. The surface morphology of catalysts was studied by a scanning electron microscope (JEOL/JSM 6380 LA). The BET surface area was analyzed on Micrometrics unit (ASAP 2020 V3.01 H). The catalyst (0.5 g) was charged into U-shaped quartz sample tube that was put in a Dewar flask containing liquid nitrogen. A prescribed amount of nitrogen was passed for adsorption. The relative pressure difference was measured to estimate the absorbed quantity of nitrogen. Desorption of nitrogen was further carried out at room temperature of 30 °C. The absorption and desorption values were used in BET equation to estimate the surface area and pore size. The X-ray diffraction of CSMC was also carried out to determine its crystallinity. Hydrolysis of Residual Glycerides in PFAD. The hydrolysis reaction was conducted in a stirred reactor of volume 250 cm3 (4.5 cm I.D.) equipped with a six blade turbine impeller running at 1200 rpm. In a typical reaction, 100 g of PFAD was first melted in the reactor at 90 °C for 10 min. Water (25 g) and sulfuric acid 0.5% (w/w), as a catalyst were added to the melt. The reaction mixture was vigorously agitated B

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very good thermal stability. Beyond 270 °C, the material showed a gradual weight loss up to 500 °C indicating the loss of nongraphite carbon initially and then of the graphite carbon in the higher temperature range. The scanning electron microscope (SEM) images of MC showed an irregular network structure and several closely linked pores which after sulfonation and carbonization at 250 °C, disintegrated to some extent with increased pore size to 4.25 nm. Microcrystalline cellulose is a very strong organic material consisting of polysaccharide linear chains of β (1 → 4) linked D-glucose units. The carbonization of the polysaccharides results in a rigid carbon framework composed of small polycyclic carbon sheets in a three-dimensional sp3-bonded structure.38−40 Microcrystalline cellulose material has a higher degree of crystallinity and thus relatively better hydrophobicity. It was expected that the catalyst will have poorer water retention characteristics, helping in maintaining the catalytic activity in the presence of water generated during the reaction. The BET surface area and mean pore size of the CSMC catalyst are 105 m2/g and 4.25 nm, respectively, whereas MC had a BET surface area of 242 m2/g and pore size of 1.3 nm. As compared to reported D-glucose and sugar based catalysts,38−42 these values are higher and thus should be useful for dealing with reacting molecules that are bigger in size. The mesoporic structure also reduces the intraparticle diffusional limitations commonly associated with the molecular sieve based acid catalysts. The carbonization of carbohydrates, such as sugar, leads to the formation of polycyclic carbon rings. The surface area of the catalyst decreased as expected with increasing pore size. The carbonization of sulfonated MC causes an increase in pore size greatly that, however, allows easy migration of the reacting species inside the catalyst matrix. The natural sources thus produce carbon catalysts having different pore sizes after the carbonization process. The X-ray diffraction (XRD) spectrum of the CSMC exhibits one broad, strong diffraction peak attributable to crystalline carbon composed of graphite carbon sheets at an angle (2θ = 25°), oriented in a considerably random fashion (given in the Supporting Information). The infrared spectrum of MC showed a strong alcoholic −OH stretching at 3440 cm−1 attributed to the −OH of the polycyclic cellulose ring. The IR spectrum of CSMC shows development of two new strong symmetric and asymmetric stretching bands of sulfonate groups at 1156 and 1134 cm−1, respectively. The spectrum also shows formation of etheral rings, characterized by a new peak at 1100−1164 cm−1 (given in the Supporting Information). In the carbonization process, all the volatile impurities evaporated from the cellulosic matrix. Catalytic Activity of the Catalyst for Esterification. A comparative study of CSMC with strongly acidic resin Amberlyst-15 was done for the esterification reaction. The CSMC catalyst showed a slightly higher esterification activity than Amberlyst-15 at the same catalyst loading (Figure 1). The esterification of PFAD:methanol (1:3 mol ratio) with the CSMC catalyst gave 62% conversion in 3 h, but the conversion with Amberlyst-15 was 47%, in the same period. Since CSMC exhibited a remarkable activity for the esterification reaction, further studies were carried out by using only the CSMC catalyst. Effect of Reaction Temperature. Figure 2 shows the effect of temperature on conversion with the same molar ratio of methanol:FFA (3:1) using sulphuric acid and CSMC as catalyst, respectively. The FFA conversion increased from 26% at 40 °C to 52% at 50 °C and became 62% at 60 °C in 3 h of

at 1200 rpm for 30 min in an oil bath. After the reaction, the reaction mixture was kept in the oil bath for 5 min to separate into two phases. The organic phase was separated from the aqueous phase and further used for the esterification reaction as such. All esterification experiments were carried out with PFAD pretreated by hydrolysis. During the reaction, samples were withdrawn for the analysis of FFA by titration. The organic and aqueous phase samples were analyzed by HPLC equipped with a Hypersil C18 column and a RI detector. A mixture of acetonitrile:acetone in the ratio of 70:30 was used as a mobile phase at a flow rate of 0.3 cm3/ min for the analysis of glycerides. The aqueous samples were analyzed on the HPLC equipped with the RI detector. Deionized water was used as the mobile phase at a flow rate of 0.5 cm3/min. Esterification of PFAD. In a typical batch esterification reaction, PFAD (50 g) was melted in a reaction vessel at 60 °C for 20 min. The reaction was started by addition of methanol (18 g) and 98% sulfuric acid (about 5% w/w of PFAD) or the solid acid catalyst (about 3% w/w of PFAD). The reaction mixture was maintained at 60 °C in an oil bath for desired time period with vigorous agitation at 1200 rpm to eliminate the effect of the external mass transfer limitations, if any, on the conversion and/or rates. Samples were withdrawn from the liquid reaction mixture at specified intervals of time for the analysis of FFA, methanol, and sulfuric acid in the organic as well as the aqueous phases. In the case of solid catalyst, after termination of reaction, the reaction mixture was separated from the solid acid catalyst by decantation. The samples (0.5 g), withdrawn at specified intervals from the reaction mixture, were dissolved in methanol (20 cm3) to get a clear solution. The solution was then titrated with an alcoholic KOH solution (0.1 mol/dm3) using phenolphthalein as an indicator. The conversion was calculated from initial and final FFA concentrations of the reaction mixture. The sample was also analyzed on a HPLC equipped with a RI detector (Jasco 1090 series) using a Hypersil C18 column (l 25 cm, D 5 mm). A mixture of acetonitrile:acetone in the ratio of 70:30 was used as a mobile phase at a flow rate of 0.3 cm3/min for the analysis of methyl esters, methanol, and glycerides. Adsorption of Methanol/Water on Catalyst. The adsorption of methanol and water on the heterogeneous catalyst was separately studied by conducting the adsorption experiments of water and methanol from methyl ester solutions. Batch adsorption studies were carried out at 60 °C by equilibrating 3 g of CSMC in the stoppered conical flask with 50 cm3 of water (or methanol) in methyl ester solutions at known concentrations (0−2 wt %). Samples were prepared in 10 μL of acetonitrile:acetone (70:30) and injected on an HPLC equipped with an RI detector (Jasco 1090 series) using a Hypersil C18 column (l: 25 cm, D: 5 mm). The adsorbed amounts of water and methanol were estimated from HPLC analysis of residual concentration in the organic phase.

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RESULTS AND DISCUSSION Catalyst Characterization. The estimated acidity of CSMC from the NH3-TPD plot is 2.5 mmol/g as compared to 0.8 mmol/g for Amberlyst 15. The CSMC thus carries a higher number of protonic acidic sites. The reported acidity of another solid acid catalyst prepared in the same manner from sugar was in the range of 1.5−1.9 mmol/g.41,42 The DSC and TGA plots of CSMC (given in the Supporting Information) show a weight loss of only 3−4% up to 270 °C indicating its C

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Figure 3. Effect of FFA to methanol molar ratio on FFA conversion (temperature 60 °C, 3% catalyst loading): ◇ 1:4, □ 1:3, Δ 1:1.

60 °C the conversion of FFA was 54% for the molar ratio of 1:1 in 3 h whereas the conversion increased to 62% and thereafter to 68% with molar ratio of 1:3 and 1:4 in 3 h, respectively. The conversion increased in given time with the increasing molar ratio of FFA:methanol because, for the kinetically controlled reaction, the rate depends on the concentrations of the reactants and the equilibrium limited reaction is pushed toward more conversion in the presence of increased amount of one of the reactants. Higher temperatures and increasing methanol content of the reaction mixture thus increase the rate of the reaction and conversion in given time. Effect of Catalyst Loading. Figure 4 shows the effect of catalyst loading, taken as ratio of CSMC catalyst and PFAD in

Figure 1. Comparison of catalytic activity (temperature 60 °C, 1:3 FFA:methanol molar ratio, 3% catalyst loading): ◊ CSMC, □ Amberlyst-15.

Figure 4. Effect of CSMC catalyst loading on FFA conversion (temperature 60 °C, 1:3 FFA:methanol molar ratio): ◇ 7%, □ 5%, Δ 3%. Figure 2. (A) Effect of temperature on FFA conversion using H2SO4 as catalyst. (B) Effect of temperature on FFA conversion using CSMC catalyst (1:3 FFA:methanol molar ratio, 3% catalyst loading): ◇ 60, □ 50, Δ 40 °C.

mass, from 3 to 7% (w/w) on the FFA conversion. About 62% conversion of FFA was achieved with 3% (w/w) CSMC catalyst loading in 3 h. As the catalyst loading was increased, the FFA conversion increased to 74% for 5% (w/w) and 89% for 7% (w/w) catalyst loading, respectively, in the same time period. The reaction rate increased with the catalyst loading and the time required for the reaction to approach the equilibrium conversion in turn reduced. An increase in the catalyst loading increases the number of the acidic sites and hence increases the rate of conversion of FFA. The reusability of catalysts is an important aspect to be considered for any solid catalytic system. Figure 5 shows that CSMC catalyst could retain the activity efficiently better as compared to Amberlyst-15 catalyst that lost its activity

the reaction time when CSMC was used as the catalyst in single stage esterification. For the sulphuric acid catalyzed esterification, the FFA conversion was similar to that obtained with fresh CSMC catalyst, however, in much less time, 1 h. Effect of FFA to Methanol Molar Ratio. Figure 3 shows the effect of molar ratio of PFAD to methanol in the feed, from 1:1 to 1:4 at same CSMC catalyst loading (3% w/w of PFAD) and reaction temperature. As the initial concentration of methanol in the initial feed mixture was increased, the conversion of FFA also increased in the given time period. At D

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Table 1. Adsorption Constants for CSMC Catalyst and Kinetic Rate Constants for Esterification Reaction kinetic rate constants adsorption constants ×103 (m3/kmol)

temperature (K)

FFA

333 323 313

0.614 0.843 1.139

CSMC

H2SO4 k1 (m3/ kmol)2/ min 0.22 0.18 0.11

methanol

methyl ester

water

k1 × 10−7 (kg/ (kmol min))

0.407 0.774 1.043

0.283 0.193 0.11

0.453 0.142 0.092

4.3 1.0 0.25

methanol into the Amberlyst-15 catalyst without considering any effect of FFA and methyl ester which may not truly represent the actual adsorption process as the methyl ester phase is in large volume as compared to the aqueous phase. The adsorption constants in Table 1 thus show the partitioning of these components from the organic phase onto the solid catalysts. The adsorption of FFA from the methyl ester organic phase (determined from the kinetics data as discussed later), is higher as compared to that of methanol and/or water. It is thus expected that the adsorption of FFA on the catalyst takes place preferentially from the reaction medium of methyl ester due to very low equilibrium concentrations of methanol and water present in the organic phase. Phase Equilibrium Analysis of Esterification Reaction Mixtures. The starting reaction mixture of FFA and methanol forms a homogeneous mixture because of miscibility of the components. But as the reaction proceeds, aqueous methanol and methyl ester form two separate phases because of limited solubility in each other. This results in distribution of all the components of the mixture between the two phases. The organic phase consists mainly of unreacted FFA and methyl ester while the aqueous phase mostly consists of alcohol, water, and catalyst (in case of homogeneous catalyzed esterification), owing to the low solubility of FFA and methyl ester in the aqueous phase. The transfer of the alcohol and homogeneous catalyst to the aqueous phase results in drastic reduction in the rate of esterification reaction thereby, reducing the overall conversion of FFA in the given time. In order to determine the kinetics of such a reaction correctly it is necessary to consider a heterogeneous biphasic reaction system instead of following a commonly used pseudohomogeneous reaction phase approach, irrespective of the nature of the catalyst used for the reaction. Using both, homogeneous or heterogeneous, catalysts, the reaction phase will undergo liquid−liquid split, after formation of appropriate amounts of water and methyl esters as even methanol and methyl esters are not completely miscible. The presence of water formed in the reaction ensures transfer of significant amount of methanol from the ester phase into the aqueous phase. It is not a straightforward determination of the kinetic parameters as using a pseudohomogeneous power law model to represent the reaction. In a batch process, at the beginning, FFA, catalyst, and methanol form a homogeneous mixture. However, once the water concentration in the reaction phase crosses the saturation solubility, the phase split takes place. Since the exact time of the phase split for a batch reactor is not known, we had to solve the batch reactor kinetic equations (given below) with time by checking the phase stability of the multicomponent mixture of FFA, methanol, methyl ester, water, and catalyst. The thermodynamic analysis of the reaction

Figure 5. (A) Reusability of CSMC catalyst. (B) Reusability of Amberlyst-15 catalyst (temperature 60 °C, 1:3 FFA:methanol molar ratio, 3% catalyst loading): ◇ run 1, □ run 2, Δ run 3, ○ run 4.

significantly over four cycles. After each run, the catalyst was washed with methanol and used for the next run. Amberlyst-15 catalyst showed a good activity in the first run with conversion of 39% which, however, reduced to 37%, in the second recycle. After the fourth recycle, the conversion decreased to 26% (Figure 5B). The catalyst also showed some physical damage, probably due to higher stirring speed used for the reaction. The free acidity of the reaction mixture also increased due to leaching of acidic groups during the reaction. Higher stirring speed, retention of water, and swelling of resin resulted in significant loss of the activity of the Amberlyst-15 catalyst. With CSMC as catalyst, in the first run, the conversion of FFA was 62%, and in the fourth run, the conversion reduced to 53%. The CSMC catalyst contains a more hydrophobic surface and, therefore, shows a better ability to withstand water formed in the reaction. The CSMC catalyst thus shows a relatively better reusability as compared to Amberlyst-15. Adsorption of Methanol and Water. Heterogeneously catalyzed esterification reaction takes place via adsorption of FFA and methanol on the solid catalyst and ends with desorption of the products. In order to model the kinetics of this reaction, it was necessary to determine the equilibrium adsorption constants of these components. Thus, adsorption of water and methanol on the catalysts was studied from methyl esters separately and the adsorption constants obtained were thereby used in the kinetic analysis of the reaction as given below. The adsorption experiments were conducted from the methyl ester organic phase since the catalyst surface is wetted by the organic phase owing to its organic and hydrophobic nature. The adsorption constants thus obtained by fitting the adsorption data in Langmuir isotherm are given in Table 1. Tesser et al.29 have reported the partitioning of water and E

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The terms k1 and k2 are forward and backward rate constants which are related through the equilibrium constant K = k1/k2. It was assumed that the reaction occurs only in the organic phase and thus organic phase concentrations determined by the phase split calculations at each time interval were used for determining the reaction kinetics. Since parameters for sulfuric acid were not available for the estimation of the distribution of sulfuric acid, experimental values of sulfuric acid concentration were used to consider effect of the distribution of the catalyst on the reaction rate. The forward rate constant (k1) can thus be estimated by solving the phase and chemical equilibria, simultaneously, by numerical integration of batch reactor equation and minimizing the error with respect to the experimental conversions for the entire set. The Modified UNIFAC-LLE (Dortmund) parameters were used for an inhouse code using gPROMS (version 3.3, Process Systems Enterprise Ltd.) for thermodynamic calculations whereas the set of ordinary differential equations of batch reactor was solved using the Runga−Kutta 4 method. The experimental data as shown in Figure 2 for the H2SO4 catalyzed esterification of PFAD at three temperatures, 60, 50, and 40 °C, were used for the estimation of the rate constants at respective temperatures. The k1 values thus obtained are given in Table 1. The lines in Figure 2 indicate the predicted values using the fitted rate constants for the reaction. The activation energy and frequency factor were determined to be 30.1 kJ/mol and 1.2 × 104 (m3/kmol)2/min, respectively. The sharp change in the slope after a few minutes of reaction in the reaction at all temperatures indicates the phase-split of the reaction mixture and, thereby, transfer of a major amount of the homogeneous catalyst to the aqueous phase and the consequent reduction in the rate of the reaction. The reaction does not stop though but proceeds with a much lower rate. Analysis of Esterification Reaction Using Heterogeneous Catalyst. The effect of internal diffusional resistance on conversion was studied by performing experiments with catalysts of different particle sizes (given in the Supporting Information). No significant difference in conversion was observed with different particle sizes indicating minimum internal diffusional resistance. Thus, the reaction was considered to be kinetically controlled. The kinetic rate equation used for heterogeneous catalyzed esterification reaction was based on the Langmuir−Hinshelwood−Hougan−Watson (LHHW) model considering the following steps: (1) adsorption of fatty acid and methanol (eqs 4 and 5), (2) surface reaction (eq 6), and finally (3) desorption of methyl ester and water (eqs 7 and 8). Considering surface reaction to be the rate-controlling step, the final eq 9 is obtained as the rate expression.

mixture was thus performed by considering chemical and phase equilibria simultaneously. At each time interval, the stability of the liquid phase was checked during numerical integration of kinetic equations with initial guess values of the kinetic parameters. Once the liquid−liquid split was detected by thermodynamic calculations, the concentrations of each species in both, methyl ester and aqueous alcohol, phases were estimated. Palmitic acid can be used as a representative of FFA since it is a major component of PFAD, the others being, oleic and linoleic acids. A modified UNIFAC-LLE (Dortmund) method43 was used to calculate the activity coefficients for the components in both the phases and validated against the phase compositions reported for esterification of oleic acid.44 The experimental values showed that about 90% of the water produced gets distributed into the methanol and sulphuric acid layer at 1:3 molar ratio of oleic acid to methanol in the initial feed. The value predicted by modified UNIFAC-LLE (Dortmund) method was 93.6% which was quite in agreement with these experimental observations. The consideration of palmitic acid or oleic acid as a representative fatty acid did not show significant difference in the phase split results. After the phase split, concentrations only in the organic phase were considered responsible for determining the reaction rates. A reaction rate equation, first-order with respect to each reactant, was considered for the equilibrium reaction between methanol and FFA in the organic phase, while water solubilized in the organic phase was considered to be responsible for the reverse hydrolytic reaction, both reactions catalyzed by the acid present in the reaction phase. The effect of speed of agitation was studied by performing the reaction at different speeds of agitation. No effect of agitation speed on conversion was observed beyond 1000 rpm. (given in the Supporting Information) The reaction was, thus, considered to be kinetically controlled since all the experiments were performed at 1200 rpm which ensured elimination of mass transfer resistance. Kinetic Analysis of Esterification Reaction Kinetics Using H2SO4 as Catalyst. The equilibrium constant for the reversible esterification reaction (eq 1) was determined using eq 2 from the standard Gibbs free energy and heat of formation values, predicted using the Joback method.45 fatty acid + methanol ↔ methyl ester + water

(1)

⎛ −ΔG 0 ⎞ ⎡ ΔH 0 ⎛ T ⎞⎤ R R ⎜ ⎟exp⎢ 1 − 0 ⎟⎥ K = exp⎜ T ⎠⎦ ⎝ RT0 ⎠ ⎣ RT0 ⎝ ⎡ 1 ⎤ ⎡ T ΔCp ⎤ T ΔCp exp⎢ − dT ⎥exp⎢ dT ⎥ ⎣ T T0 R ⎦ ⎣ T0 RT ⎦

(2)

∫

∫

Since ΔCp ≈ 0, only first two factors are important. The predicted values of ΔG° and ΔH° for the components and the value of K are given in the Supporting Information. Equation 3 was used to represent the rate equation for homogeneous catalyzed esterification considering the reaction only in the organic phase because of very poor solubility of the fatty acid in the aqueous methanol and H2SO4 mixture.

KA

A + S ↔ AS KB

B + S ↔ BS Ks

AS + BS ↔ CS + DS

c c c ⎞ dC ⎛ rA = − A = k1cAc Bc E − k 2cCc Dc E = k1⎜cAc Bc E − C D E ⎟ ⎝ dt K ⎠

Kc

CS ↔ C + S

(3)

where, cA, cB, cC, cD, and cE are concentrations of fatty acid, methanol, methyl ester, water, and sulfuric acid, respectively.

KD

DS ↔ D + S F

(4) (5) (6) (7) (8)

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1 dNA W dt

(

k1ct 2 KAKBcAc B − =

K CKDcCc D Ks

medium was thus formed immediately after the initial very fast rate of the reaction, unlike in enzymatic esterification of oleic acid where the aqueous phase separated when the conversion reached to about 15% after a period of 6 h.37 The simulations show that the initial rate of reaction was high leading to a conversion of FFA of about 30% in the first few minutes owing to the homogeneous starting reaction mixture. This resulted in formation of about 9% (mol basis) of water. Water is almost insoluble with the fatty acid and its methyl ester. This was also observed from the graph of Gibbs free energy of mixtures of oleic acid−water and methyl oleate−water (given in the Supporting Information). Thus, water separated out immediately, leading to a biphasic reaction medium. If the amount of water in the initial reaction mixture is low or if the methanol content of the reaction mixture is high, the water may remain back in the reaction phase favoring the backward hydrolytic reaction, thus decreasing the net rate of ester formation. The presence of water in the reaction mixture can limit the conversion. But as soon as the phase splits, the water is transferred to another phase, thereby favoring more ester formation albeit at slower rates. In principle, the phase split helps in improving the extent of conversion as the reaction is driven more toward the product formation and overcomes the equilibrium conversion limitations set by thermodynamics. Although a relatively large amount of experimental data concerning the esterification reaction of fatty acids for alkyl esters production is found in the literature, only a few reports take into account the biphasic nature of the reaction into consideration. The thermodynamic phase analysis of the esterification of oleic and palmitic acids has been investigated at low pressures using simultaneous chemical and phase equilibria by Voll et al.36 But the authors did not discuss about the partitioning of the catalyst nor the kinetic parameters of the reaction were reported using the phase split. Aranda et al.14 also have reported the kinetics of the sulfuric acid catalyzed homogeneous esterification of palm fatty acids for biodiesel production but assuming homogeneous reaction, thus giving the first-order rate constant to be 106.1 h−1 at 433 K. In view of liquid phase split as discussed above, this rate constant remains only a fitted value. Aafaqi et al.35 studied the kinetics of esterification of palmitic acid with isopropanol using p-toluene sulfonic acid as catalyst. They determined the kinetics by considering esterification again as an elementary second-order reversible reaction. The forward rate constant was estimated to be 0.317 m3/(kmol h) at 373 K, considering that all the catalyst and methanol are completely miscible with the organic phase and thus available for the reaction. The present work clearly shows that formation of two liquid phases leads to distribution of the catalyst in the two phases leaving very low concentrations of the homogeneous catalyst in the organic phase. Since the reaction occurs in the organic phase, these rate constants ought to be quite higher owing to the low organic phase catalyst concentrations. In the current work, the estimated forward rate constant is 13.2 (m3/kmol)2/h at 333 K. The catalyst independent rate constant, therefore, becomes 4.3 m3/(kmol h) for the given H2SO4 concentration. This value is higher than those reported by previous researchers due to effective concentration of the reactive components after distribution of the components of the mixture in two separate liquid phases. Similarly, in the case of heterogeneous catalyzed reaction, though the studies in the literature use heterogeneous reaction models (solid−liquid), the effect of phase splitting was still not

)

(1 + KAcA + KBc B + K CcC + KDc D)2

(

k1′ct 2 cAc B − =

Article

cCc D Ke

)

(1 + KAcA + KBc B + K CcC + KDc D)2

(9)

where, A, B, C, D, and S, represent fatty acid, methanol, methyl ester, water, and catalyst site, respectively. AS, BS, CS, and DS represent the adsorbed species of each of these components. k1 is forward rate constant of the reaction (kg/(kmol min)) and k′1 = k1KAKB while ct is total catalyst sites available (2.5 kmol/kg). The Ki are equilibrium adsorption constants of each of the species in the reaction mixture, i.e. fatty acid, methanol, methyl ester, and water, respectively (m3/kmol), while Ks is surface reaction equilibrium constant and Ke is homogeneous reaction equilibrium constant. V is the volume of organic phase (m3), and W is the amount of catalyst (kg). The equilibrium adsorption constants of fatty acid and methyl ester were estimated by fitting the data of experimental conversion with respect to time in the LHHW model for expressing the rate of the reaction. The equilibrium adsorption constants of methanol and water were obtained from independent adsorption experiments on the catalysts as described previously. The fitted adsorption constants that were used for calculating the kinetics are given in Table 1. The reaction mixture was analyzed for liquid−liquid phase split at each time interval using modified UNIFAC-LLE (Dortmund) method as mentioned in the homogeneous catalyzed reaction while integrating the reaction rate expression at each time interval. The calculations similar to homogeneous system were performed to determine the forward rate constant of the reaction. The fitted conversions are shown in Figure 2B as solid lines. The forward rate constants thus obtained by fitting the experimental data of esterification of PFAD at different temperatures are given in Table 1. The activation energy and frequency factor were determined to be 123 kJ/mol and 8.8 × 1026 (kg/(kmol min)), respectively. In case of heterogeneous catalyst, the phase splitting does not result into a sharp change in the slope as observed for the homogeneous catalyst. This is because the phase splitting of reaction mixture does not result in distribution of the heterogeneous catalyst as in case of homogeneous catalyst owing to its solid nature. In addition to this, the splitting leads to transfer of water to the aqueous phase thereby reducing the rate of backward reaction in the organic phase thus, driving the reaction more toward the desired product. The kinetic rate constants determined by fitting the data were further used to predict the effect of mole ratio (Figure 3) and catalyst loading (Figure 4) on conversion of FFA. The predicted curves indicate that the estimated kinetic rate constants are able to predict the behavior of the heterogeneously catalyzed esterification system at different experimental conditions. According to the amount of water initially present in the reaction mixture and to the level of conversion achieved, reaction may proceed in a monophasic medium or in a biphasic medium.37 Despite having initial water content of 0.5 wt % of PFAD, the reaction proceeded to a significant extent and with a very high rate of reaction because of initial high concentrations of reactants until the phase split took place. The biphasic G

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Figure 6. (A) Conversion with time in crosscurrent flow of reaction phases at three temperatures. (B) Concentration profiles in crosscurrent flow three stage esterification.

considered. Esterification of palmitic acid with iso-propanol by using zinc ethanoate supported over silica gel as catalyst was carried out by Aafaqi et al.35 They have also used LHHW model considering surface reaction as the controlling mechanism for predicting the kinetics of the reaction. The concentrations used for the fitting were, however, the bulk overall pseudophase concentrations neglecting the phase separation. The kinetics of esterification of oleic acid using an acid ion exchange polymeric resin has been studied by Tesser et al.29 using the Eley−Rideal model. The experimental data have been interpreted with a kinetic model based on an ion exchange reaction mechanism that takes into account also the physical partitioning effects of the various components of the reacting mixture between the liquid phase internal and external to the resins but not the formation of two separate liquid phases. The current work considers the similar LHHW model but incorporates the effect of phase splitting on the reaction kinetics by considering phase equilibria of the reaction system. The organic phase was considered as the reaction phase in contact with the catalyst active sites. At equilibrium, the water produced in the reaction is mainly distributed into the aqueous phase, with a small amount remaining in the organic phase. The presence of water in a monophasic system would shift the reaction equilibrium to the reactants because of the hydrolysis reaction. Thus, consideration of a homogeneous/monophasic kinetics would lead to erroneous prediction of the rate constant. However, when two liquid phases are considered, the water formed in the reaction medium is not available in the organic phase, allowing the

reaction to proceed to products, thus representing the real scenario. Thus, prediction of liquid−liquid equilibrium must be used for reliable prediction of the chemical equilibrium of esterification reaction of fatty acids. Even for the heterogeneous catalyzed reactions, the separation of aqueous phase is an important factor, which probably was not observed because of significant water retention by the catalysts in the previous reports. The CSMC catalysts, being highly hydrophobic, retains far less water, as shown by the poor adsorption constant of water on the catalyst. An ideal catalyst of esterification would be a completely hydrophobic catalyst which would retain no water at all. In that case, the reaction can be taken to completion without any large excess of methanol. Even if the reaction equilibrium constant seems to be low, the phase separation drives the reaction toward more ester formation. Had water formed in the reaction not separated into a second liquid phase, the reaction would have faced significant equilibrium limitation. Essentially, it is the liquid split that helps in driving the reaction to higher conversions due to limited water solubility in the organic phase. Process Engineering Aspects of PFAD Esterification. A single stage esterification reaction does not give complete conversion of PFAD as desired even with using 200% excess methanol in a batch reactor. One can use a large excess of methanol to drive the reaction to almost complete conversion. Many laboratory batch kinetics studies are based on the molar ratio of methanol to FFA from 6 to 10, in some cases using still higher values, to drive the reaction to completion. But the entire excess amount of methanol ends up in the aqueous phase H

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Figure 7. Concentration profiles in countercurrent flow three-stage esterification.

third stages using the rate constants fitted earlier that show that the FFA conversion can go beyond 98% when operated in three stages. This crosscurrent approach still uses a significantly large amount of methanol for keeping the methanol to organic phase at 3 in each stage. If the amount of the methanol is reduced in the second and third stages to keep in line with reduced quantity of PFAD in the reaction phase, the rate goes down substantially, probably because of lower concentrations of methanol in the organic phase as methanol gets diluted by water formed in the reaction. The conversion of FFA to corresponding methyl esters thus can be increased to almost completion using sulfuric acid as the catalyst and fresh methanol in each stage. However, the total amount of methanol with respect to FFA can be still prohibitively higher in such a multistage reaction system as all the methanol ends up in the aqueous phase and requires recovery later by distillation. The cost of distillation, to recover almost unhydrous methanol for the esterification reaction, adds to the cost of the product and thus this ratio has to be brought down to reasonable level. In the countercurrent manner of conducting this reaction, however, in principle, the first stage gets fresh FFA feed, that is contacted with the aqueous phase coming from the second stage while the organic phase from the second phase is contacted in the third stage reactor with fresh methanol feed. The limited miscibility of methyl ester product with methanol allows phase separation of methanol from the final product in the third stage, and it can be charged into the second reactor after the separation. All the three reactors thus work with twophase systems, the first reactor giving aqueous methanol as another product stream of the process that can be sent for the methanol recovery. The organic phase reaching to the third reactor attains almost 90% conversion and meets the entire amount of methanol charged in the system where the operating ratio of methanol to unreacted FFA in the organic phase is very high and thus ensures FFA conversion in excess of 98%. The first reactor has the highest feed concentration of FFA and the reaction is driven toward the ester formation to a significant extent even when the entire amount of water generated in the reaction in present in this stage. A typical concentration profile in experimentally simulated countercurrent flow of two

which has to be separated. This methanol is recovered by distillation in subsequent operations for recycle. In order to obtain relatively anhydrous methanol, as required for the esterification reactions, a high reflux ratio is commonly employed. This distillation operation adds a significant energy cost to the recovery operation and also to the product cost. The phase split, on one hand, helps in driving the reaction toward formation of more esters in the organic phase where only the soluble amount of water can facilitate the hydrolytic reaction. A common engineering principle of separating one of the products from the equilibrium limited reaction mixture can be exploited to reduce the excess amount of methanol used for the reaction. The phase split, because of not only the immiscibility of water with the organic esters and FFA but also the partial immiscibility of methanol with the methyl esters, can be judiciously exploited. We decided to, therefore, investigate a process where methanol and FFA are brought in contact with each other in crosscurrent and countercurrent manners. The process was simulated in a series of batch reaction operations of three stages. The crosscurrent operation is more common where fresh methanol and acid catalyst are added to the reaction phase to improve the rate and conversion of the reaction. At each temperature, the reaction with sulfuric acid as catalyst was carried in three steps by adding fresh methanol and the catalyst in each step, keeping the reaction time of 1 h the same for each stage. The organic phase was separated from the aqueous phase and then treated with fresh methanol and catalyst in subsequent step. Figure 6 shows typical concentrations of the components in different phases at each stage. The FFA conversion increases from 81% at 40 °C to 98% at 60 °C after the third step in total reaction time of 3 h. The breaks in FFA conversion profiles at 1 h intervals in Figure 6 are because of the fresh methanol addition. The second and third stages of the reaction start with two phases as methanol is partially miscible with the ester phase. Thus, a significant percentage of the homogeneous catalyst gets transferred to the methanolic phase immediately, and therefore, no sudden change was observed in the second and third stages of reaction as was seen in the first stage. The lines in Figure 6 are predicted conversions in the second and I

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reactants is given in Figure 7, with the methanol to FFA molar ratio of 3:1. It is clear that even with the lower molar ratio for the entire system, it is possible to approach the conversion in excess of 98%. For still higher conversions, the organic phase will have to be dried completely before it is fed to the third reactor. The reduced methanol feed to the system also means less methanol to be distilled in the distillation column bringing down the energy penalty by a factor more than two as compared to commonly used crosscurrent operation. The less load on the distillation column reduces the capital cost of the unit as a smaller diameter column is required for the distillation. We have also considered the removal of glycerides from the PFAD by hydrolysis as pretreatment of the PFAD feed. We conducted the hydrolysis step to convert the residual glycerides in the feedstock to FFAs which are further converted to methyl esters by the acid catalyzed esterification. The hydrolysis reaction, if conducted with dilute H2SO4, then the FFA feed is available for the esterification system directly without any drying step as it anyway is contacted with aqueous methanol in the first stage of the esterification. The major advantage of the hydrolysis step is the absence of glycerol in the aqueous acidic methanol stream from the esterification process. The glycerol, if present, makes methanol recovery from this aqueous stream difficult due to its decomposition under acidic conditions in the presence of homogeneous acid catalyst like H2SO4 unless it is neutralized. It is necessary to distill a major amount of methanol before the neutralization of the catalyst acid, as FFAs solubilized in the aqueous methanol can also undergo reaction with the alkali. On distilling out the major amount of methanol from the aqueous phase, without increasing the temperatures of the solutions in excess of 100 °C, the FFAs separate out from the aqueous acidic solutions. After separation of the FFA from the acidic phase, and then neutralization of the homogeneous acid catalyst, the remaining amount of methanol, if any, can be distilled out. Even a small amount of fatty acid carried into aqueous methanol solutions can otherwise give rise to soap while neutralizing the mineral acid and then causing foaming in the recovery column that makes the column operation difficult. H2SO4 still remains the cheapest catalyst for the esterification process, as the other homogeneous catalysts like aromatic/ aliphatic sulfonic catalysts on neutralization give rise a significant organic load on microbial wastewater treatment process. The development of carbonized sulfonated MC is likely to provide a better solution in such multistage reaction system. In principle the multistage esterification, either crosscurrent or countercurrent, can also be used for the heterogeneous catalyst, provided the activity is retained in presence of large amount of water in the reaction phase.

Article

ASSOCIATED CONTENT

S Supporting Information *

The characterization of CSMC catalyst using NH3-TPD, DSCTGA, SEM, XRD, and FT-IR, esterification crosscurrent and countercurrent flow block diagram, effect of particle size of CSMC catalyst on conversion, Gibbs free energy of oleic acid− water and methyl oleate−water mixtures, and the values of standard Gibbs free energy, heat of formation, and thermodynamic equilibrium constant. This information is available free of charge via the Internet at http://pubs.acs.org/.

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AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Phone: 091-2233612013. Fax: 091-22-33611020. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS We would like to acknowledge financial support provided by University Grant Commission (UGC).

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REFERENCES

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CONCLUSIONS Carbonized sulfonated microcrystalline cellulose (CSMC) gave better conversion and water tolerance in esterification reaction of PFAD with methanol as compared to Amberlyst-15 catalyst. The kinetic data analysis considering phase equilibrium simultaneously with reaction equilibrium allows identification of true kinetic parameters of the esterification reaction. The presented methodology represents the two phase esterification reaction system in a more realistic way. Also a countercurrent approach of two reaction phases has been demonstrated for homogeneous catalyst to take the reaction to almost completion without using a large excess of methanol that is more attractive for the industrial adoption of the process. J

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K

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