Ind. Eng. Chem. Res. 2010, 49, 3217–3222
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Applicability of a Countercurrent Enzymatically Catalyzed Multistage Extractive Reaction Process for the Hydrolysis of Methyl Octanoate Przemyslaw Krause,* Roberto Macias, and Georg Fieg Institute of Process and Plant Engineering, Hamburg UniVersity of Technology, Schwarzenbergstrasse 95, 21073 Hamburg, Germany
An assessment of the minimum requirements for a countercurrent multistage extractive reaction process for the enzymatic hydrolysis of methyl octanoate with immobilized lipase in undiluted media was made. For simulation, UNIQUAC parameters for the quaternary reaction system containing methyl octanoate, octanoic acid, methanol, and water at 48.5 °C and atmospheric pressure and the equilibrium conversion, as a function of the initial water content, were determined. For a one-equilibrium stage reaction, good overall agreement, with respect to the liquid-phase compositions between experiments, and results from chemical and phase equilibrium calculations, based on Gibbs free-energy minimization, was obtained. A subsequently performed analysis of the influence of the stage number and solvent:feed ratio, using commercial flowsheet simulation software, showed that, in a countercurrent process, five equilibrium stages are required to obtain octanoic acid with a purity exceeding mass fractions of 0.95 at comparably moderate water excess. 1. Introduction The use of enzymes in the chemical industry is already established in several fields, and the number of biocatalytic processes is expected to increase rapidly.1,2 However, biotechnological processes may often be associated with heterogeneous systems and complex phase behavior.3 One example is the enzymatic hydrolysis with immobilized lipase in undiluted media, i.e., without additional organic solvents, because it typically requires two liquid phases and one solid phase. The thermodynamic limits of the process then are not only defined by chemical equilibrium but also by phase equilibrium. In such scenarios, simultaneous reaction and separation can offer advantages that cannot be achieved by conventional production routes.4 An interesting example can be taken from the synthesis of short-chain fatty acids via hydrolysis of their methyl esters applying the lipase B of Candida antarctica in a stirred reactor.5,6 Representatively, for short-chain fatty acids, which normally comprise fatty acids from plant oils containing 6-10 C atoms, the hydrolysis of methyl octanoate catalyzed by lipase immobilized on solid particles was within the scope of investigation. The schematic illustration of the solid-liquid-liquid reaction system is given in Figure 1. The composition of the two liquid phases is determined by the liquid-liquid equilibrium and the chemical reaction, which is defined for the hydrolysis of methyl octanoate as follows: C7H15COOCH3 + H2O / C7H15COOH + CH3OH In the past, this synthesis route was considered as an alternative within industrial practice to obtain fatty acids from excess methyl esters resulting from oil transesterification. In one of the (semi)batch process alternatives developed by Fieg et al.,5 continuous byproduct (i.e., methanol) removal is achieved by direct evaporation through the use of a dephlegmator to increase the fatty acid yield, while water is constantly provided to the reactor. Alternatively, in the procedure proposed by Masami et al.,6 many batch stages are performed to achieve high yield, * To whom correspondence should be addressed. E-mail: p.krause@ tu-harburg.de.
whereas the aqueous phase is separated from the reaction mixture, processed to remove methanol and replenished with water after each stage and subsequently used for the next reaction stage. The existence of two liquid phases in the reactor and the stage-to-stage replacement of the aqueous phase suggest that the use of a multistage extractive reaction (MSER) system might substitute the byproduct removal by evaporation. Hence, the goal of this work was to study the characteristics and thermodynamic limits of an extractive reaction process for implementation in the hydrolysis of a short-chain fatty acid methyl ester. Samant and Ng4,7 gave an extensive analysis of the basic principles of MSER processes, as well as some theoretical examples. They described the behavior of systems containing two liquid phases (where one was an inert solvent) and proceeded to analyze the advantages of using MSER units. For assessment of the industrial applicability of reaction equilibrium processes, a representation of achievable stream compositions without the necessity of performing the full process design provides a way to reach first conclusions. Graphical representations are one applied method. However, a graphical representation of these systems comprises the problem of depicting the reaction extent. For this reason, the procedure proposed by Ung and Doherty8 can be used to describe a multiphase system in which one or more chemical reactions are present. The major advantage of this transformation is the fact that the resulting coordinate system is independent of the reaction extent, thus reducing the degrees of freedom of the system and making a full representation in a two-dimensional graph possible. Considering the reaction system for the hydrolysis of fatty acid methyl esters described by Fieg et al.,5 water is utilized
Figure 1. Solid-liquid-liquid reaction system for enzymatic methyl octanoate hydrolysis.
10.1021/ie901789p 2010 American Chemical Society Published on Web 03/02/2010
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both as a reactant and as a solvent to remove the resulting methanol. The example arrangement examined within this work includes the following assumptions: the enzymatic reaction taking place at the catalyst surface is controlled by the composition of the organic phase; the aqueous phase is the solvent phase, which extracts the produced methanol and saturates the organic phase with water; the catalyst particles remain at each reaction stage; furthermore, the solubility of fatty acid and of methyl octanoate in the aqueous phase is negligible. For valid simulations of the extractive reaction process equilibrium stages, an adequate description of the system and verification of the results are required. Consideration of phase and reaction equilibrium is necessary since both occur simultaneously in the system. Therefore, experiments were performed to obtain the equilibrium conversion at different feed compositions and to complete missing liquid-liquid equilibrium (LLE) data. Equilibrium and process calculations were performed using software from Aspen Technology, Inc. 2. Experimental Section Two types of experiments were conducted. The first type was the determination of equilibrium conversion, as a function of the initial water content. The second type was the completion of missing thermodynamic data. 2.1. Determination of Equilibrium Conversion. Methyl octanoate from Dow Halterman Custom Processing (Kallo, Belgium) with 98% surface in the gas chromatography (GC) analysis was utilized for all experiments and further purified before use up to 99.8% (via GC). Novozym 435 from Novozymes A/S (Bagsvaerd, Denmark), which is immobilized Candida antarctica lipase B, was applied as catalyst for the hydrolysis reaction. As reference substances for analysis, octanoic acid (purchased with a purity of >99% (GC) from Merck Schuchardt OHG (Hohenbrunn, Germany)) and methanol (from Merck KGaA (Darmstadt, Germany) with a purity >99.9% (GC)) were used. Demineralized water and methyl octanoate were fed into a double-walled stirred reactor and 0.05 g of fresh Novozym 435 per gram of reaction mixture was added. The mixture was allowed to react at 48.5 °C until equilibrium was reached. Samples were taken after 6, 7, and 8 h of reaction time. After 6 h, no change in composition could be observed. Samples taken after 8 h reaction time were used for further evaluation. Samples from the organic phase were filtered to remove solid particles and analyzed by gas chromatography (Perkin-Elmer Model Clarus 500 fitted for flame ionization detection (FID); column, Supelco SPB-5; column length, 30 m; carrier gas, nitrogen) using the external standard method. Methanol and octanoic acid mass fractions in the organic phase were determined. This procedure was conducted at the following initial water mass fractions in the reactor: 0.10, 0.20, 0.34, 0.49, 0.66, 0.84, and 0.90. 2.2. Determination of Liquid-Liquid Equilibrium (LLE) Data. The mathematic description of the liquid-liquid phase equilibrium required additional experimental data for two ternary systems for UNIQUAC parameter determination: methyl octanoate and octanoic acid with both water and methanol. The same chemicals as those previously mentioned were used. Two experimental approaches to determine LLE data were applied. On one hand, experimental values were found using turbidity titration, following the method proposed by Maeda et al.9 Points on the binodal curve up to methanol mass fractions of 0.16 were determined. On the other hand, data were obtained for both ternary systems by experimental determination of the component
mass fractions in both phases at equilibrium at 48.5 °C for five methanol mass fractions in the aqueous phase between 0 and 0.1. Mixtures with 1:2 mass ratios of methyl octanoate or octanoic acid and water were prepared and appropriate amounts of methanol were added. Subsequently, the ternary systems were allowed to reach phase equilibrium. At equilibrium, samples were taken from each phase and analyzed by gas chromatography (Perkin-Elmer Model Clarus 500 fitted for FID; column, Supelco SLB-5 ms; column length, 15 m; carrier gas, nitrogen) using the internal standard method. Methanol mass fractions in both phases and ester or acid mass fractions in the organic phase were determined. 3. Results and Discussion The assessment of applicability of a MSER process to increase the achievable overall conversion of methyl octanoate during enzymatic hydrolysis was within the scope of this work. A countercurrent MSER cascade with N stages is considered. In each stage, the entering aqueous and organic phases are allowed to reach chemical equilibrium and phase equilibrium and are subsequently separated into aqueous and organic streams leaving the stage. Herein, fresh methyl octanoate is defined as feed F and demineralized water is defined as solvent S, whereas the resulting stage product streams are denoted as the raffinate Ri (organic phase) and the extract Ei (aqueous phase), where i is the stage number. 3.1. Phase Equilibrium and Chemical Equilibrium. For determination of the stream compositions entering and leaving the stages, a basis of calculation is required which includes information about the phase equilibrium and chemical equilibrium. Since only information about the equilibrium composition is necessary, the values can be obtained from minimization of the Gibbs free energy. This can be done applying the RGibbs model from Aspen Plus 2006.5 flowsheet simulation software (which is a product of Aspen Technology, Inc., Cambridge, MA). The RGibbs model uses Gibbs free-energy minimization to calculate simultaneous phase and chemical equilibria. It must be provided with an appropriately parametrized thermodynamic model for computation of the LLE and standard Gibbs free energies of formation. Methyl octanoate and octanoic acid exhibit wide miscibility gaps with water up to comparatively high methanol mass fractions. In addition, the solubility of methyl octanoate and octanoic acid in water is negligible at low methanol mass fractions. The liquid-liquid phase equilibrium was described by applying the UNIQUAC model. To obtain the binary interaction parameters, a LLE data regression was performed using literature data for two binary systems and data from our own experiments for two ternary systems. From these datasets, all binary interaction parameters (except for methyl octanoateoctanoic acid) were regressed. The methyl octanoate-octanoic acid pair is assumed to exhibit ideal behavior. For the methyl octanoate-water and octanoic acid-water binary systems, the required liquid-liquid solubility data were taken from literature reported by Korgitzsch.10 An evaluation of the given solubility data points results in a water mass fraction in the organic phase at 48.5 °C of 0.035 for octanoic acid-water and 0.006 for methyl octanoate-water. The solubility of the organic components in the aqueous phase is minor or not available, respectively. To account for the methanol influence, these data were completed with results from our own experiments for the methyl octanoate-water-methanol and octanoic acid-water-methanol ternary systems, because literature data were not available. For methanol contents exceeding the range of the reaction experi-
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Table 1. UNIQUAC Parameters for the Quaternary Reaction System Binary Interaction Parameters component i
component j
a(i,j)
a(j,i)
water octanoic acid methyl octanoate octanoic acid methyl octanoate methyl octanoate
methanol water water methanol methanol octanoic acid
-0.57678 -0.73918 -1.81081 -0.33880 -0.28215 0
0.84892 -0.62083 -0.69045 0.32746 -0.49210 0
Pure-Component Parameters
Figure 2. LLE data for the octanoic acid-water-methanol ternary system at 48.5 °C.
component
r
q
methanol methyl octanote octanoic acid water
1.4311 6.8214 6.2340 0.9200
1.432 5.816 5.312 1.400
Table 2. Standard Gibbs Free Energies and Standard Enthalpies of Formation for the Involved Pure Components
Figure 3. LLE data for the methyl octanoate-water-methanol ternary system at 48.5 °C.
ments and for the solubility of acid and methyl ester in the aqueous phase, which is relatively low (especially for the methyl ester), data were obtained using UNIFAC model calculations (with the LLE dataset) implemented in Aspen Properties 2006.5 (a product of Aspen Technology, Inc., Cambridge, MA). In Figures 2 and 3, the resulting LLE data for both ternary systems and the calculated binodal curves are depicted. For the sake of improving the readability, a logarithmic scale was chosen for the methanol mass fraction. A comparison of the tie-line slopes in both systems shows that methanol is better soluble in the acid than in the organic phase of the ester system. Determination of the UNIQUAC parameters was performed with Aspen Properties 2006.5, utilizing the incorporated parameter regression tool. For parameter determination, only the data pairs connected by tielines were utilized, which were obtained from experimental determination of the component mass fractions in both phases and UNIFAC calculations. Data from turbidity titration were used to confirm the results for the calculated courses of binodal curves, which were determined by applying the regressed UNIQUAC parameters. With respect to the binodal curve courses, good agreement between turbidity titration data and values from LLE experiments can be observed. The resulting parameters for calculations at 48.5 °C and atmospheric pressure for the UNIQUAC model implemented in Aspen Properties 2006.5 are given in Table 1. To describe the chemical equilibrium, standard Gibbs free energies of formation are required. Because the value of the Gibbs free energy of formation of methyl octanoate is not available in the literature, it was estimated using the Gani Group contribution method11 implemented in the Aspen Properties
component
∆Gf [kJ/mol]
∆Hf [kJ/mol]
octanoic acid methyl octanoate water methanol
-325.0 -283.0 -228.6 -162.3
-556.0 -620.5 -241.8 -200.9
2006.5 software. The remaining values were obtained from the integrated Aspen Physical Property System. The utilized standard Gibbs free energies of formation (∆Gf) and standard enthalpies of formation (∆Hf) for each of the components involved in the reaction system are shown in Table 2. All parameters and properties determined using this methodology were used for the subsequent equilibrium stage conversion calculations with Aspen Plus. To ensure the validity of the calculations, experimental and computed conversions were compared for a process that was comprised of one equilibrium stage. Therefore, the RGibbs model was set up for a single-stage process (N ) 1). Figure 4 shows the equilibrium mass fractions of octanoic acid (w°acid) and methanol (w°methanol) in the organic phase for a single-stage reaction of methyl octanoate and water, as a function of the initial water mass fraction in the reactor (wwater). A relatively high excess of water is required to achieve octanoic acid mass fractions in the organic phase exceeding 0.70 at equilibrium. Good overall agreement between experimental and simulated results using parameters specified in Tables 1 and 2 can be observed. Hence, the simulation provides a reliable representation of the thermodynamic limits of the system. An illustration of the liquid-liquid phase behavior of the quaternary reaction system is depicted in Figure 5. It shows
Figure 4. Experimental and calculation results for organic phase equilibrium mass fractions of octanoic acid and methanol at 48.5 °C and atmospheric pressure as a function of the initial water mass fraction in the reactor.
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Figure 6. Contour lines for fatty acid mass fractions (w°acid) in raffinate.
N ) 1-6 were set up. The desired mass fractions of acid in the raffinate stream RN were set as design specifications and the required solvent:feed ratios were determined at 48.5 °C for defined cases. Based on this procedure, contour lines can be drawn for different design specifications, which show the effect of the number of stages in the solvent:feed ratio. Figure 6 gives the resulting contour lines for octanoic acid mass fractions w°acid in the organic product stream RN (w°acid ) 0.60-0.95). From Figure 6, two major effects can be observed. On one hand, increasing the product purity via a simple increase of solvent flow in a N ) 1 stage process requires extreme solvent: feed ratios. On the other hand, to increase the stage number beyond N ) 5, the reduction of solvent flow required to achieve the desired purity is minor. To validate the found results, it is necessary to make a comparison with further experimental data, e.g., from literature, which is, however, scarce. Masami et al.6 investigated a cocurrent stepwise enzymatic hydrolysis of methyl octanoate. It contained separation and processing of the aqueous phase to remove methanol and replenish it with fresh water after each equilibrium stage before use in the subsequent reaction stage. To allow a comparison with the results of this work, a co-current five-stage process with fresh water feed at each stage in a 5:1 water:organic phase (pure methyl ester at stage 1) mass ratio was simulated, using the parameters from Tables 1 and 2, analogous to the experiments performed by Masami et al.6 For the simulation, Aspen Plus was used again, whereas the organic phase effluent of each RGibbs reactor model was passed to the subsequent stage, where fresh water was added. The calculated mass fractions of octanoic acid in the organic phase and results reported by Masami et al.6 are given in Table 4. The comparison of the fatty acid mass fractions at all five stages shows that both studies are in good agreement. Although the two studies were performed at different temperatures, a comparison of the results is acceptable, because it was found that the temperature dependence of the observed equilibrium mass fractions of
Figure 5. Liquid-liquid phase behavior of the quaternary system and influence of the initial water content on equilibrium composition.
the dependence of the organic and aqueous phase reaction equilibrium composition on the initial methyl octanoate mass fraction (wester) for the examined cases given in Figure 4. For wacid ) 0 (initial composition), the system consists only of methyl octanoate and water. Plotting the discrete equilibrium points from Figure 4 into Figure 5 identifies the compositions that satisfy both phase equilibrium and reaction equilibrium. The discrete organic phase equilibrium points are marked with a curve at the surface formed by the organic phase equilibrium composition, and they are connected to the corresponding aqueous phase compositions by tielines. The resulting data pairs are given in Table 3. It can be seen that, as the initial methyl octanoate mass fraction decreases, the methanol content in the system decreases and a higher octanoic acid mass fraction in the organic phase are achieved, whereas the methanol mass fraction in the aqueous phase is ∼5-8 times higher than the corresponding value in the organic phase. The position of the reaction equilibrium is very sensitive toward the methanol content and therefore is shifted to the methyl ester side already at comparatively low methanol mass fractions. 3.2. Design of the MSER Process. To realize the design of the MSER process, an analysis of the degrees of freedom must be performed. For the sake of simplicity, focus is kept only on the design of the MSER unit, without consideration of the external separation units and/or reflux flows. To solve the design problem, specifying the feed and solvent composition and the solvent:feed ratio (or desired yield), the problem is fully determined. Because of the technical implications of downstream processing, the mass fraction of octanoic acid in the raffinate stream was selected as the design specification, with feed and solvent compositions fixed to pure methyl octanoate and pure water. Using Aspen Plus flowsheet simulation software and the RGibbs model, MSER units with stage numbers ranging from Table 3. Calculated Tie-Line Data for Reaction Equilibrium Compositions
Equilibrium Initial wester [g/g] 0.101 0.164 0.344 0.508 0.662 0.796 0.897
Aqueous Phase waqester [g/g]
waqacid [g/g]
