Triethyl Citrate Synthesis by Reactive Distillation - Industrial

Ind. Eng. Chem. Res. 2008, 47, 1017-1025

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Triethyl Citrate Synthesis by Reactive Distillation Aspi K. Kolah, Navinchandra S. Asthana, Dung T. Vu, Carl T. Lira, and Dennis J. Miller* Department of Chemical Engineering & Materials Science, Michigan State UniVersity, East Lansing, Michigan 48824

A continuous reactive distillation process is proposed for the synthesis of triethyl citrate from citric acid and ethanol in the presence of macroporous Amberlyst 15 ion-exchange resin catalyst. The process design, developed using ASPEN Plus simulation software, is based on laboratory kinetic and thermodynamic studies and pilotscale reactive distillation experiments. Pilot-scale experiments were carried out in a 5-m glass reactive distillation column; catalyst effectiveness was then determined from ASPEN Plus simulation of pilot-scale experiments using a user-written reaction kinetic module based on activity coefficients. Because citric acid esterification kinetics are slow, complete conversion could not be obtained in the pilot-scale column. Using parameters determined from simulation of a pilot-scale column experiment, design of a reactive distillation column to completely convert citric acid to triethyl citrate was carried out. Optimum column performance occurs at low reflux ratios (L/D < 0.1) to avoid water reintroduction and at moderately elevated pressure (2.5 bar) to increase temperature and enhance kinetic rates without leading to product degradation. The effect of ethanol feed position and values of reflux and boilup ratios were also examined. A large number of reactive stages is required because of the slow reaction of diethyl citrate to triethyl citrate. As a final step, the design of a complete commercial-scale process to produce 25 million lbs/y triethyl citrate, with the reactive distillation column as the core component, was carried out. Three different process schemes were examined. In the first scheme, only a reactive distillation column is used. The second uses a prereactor followed by a reactive distillation column. In the third and preferred scheme, a prereactor followed by a simple distillation column to remove water is placed ahead of the reactive distillation column. In each configuration, triethyl citrate product yield is maintained above 98.5 wt %, with the main byproduct being diethyl citrate. Comparison of stream compositions and equipment design parameters is provided for the three schemes considered. 1. Introduction Organic acid esters, produced by the reaction of organic acids and alcohols, can be entirely biorenewable or “green” chemicals that replace petroleum-based solvents. Triethyl citrate (TEC), synthesized via esterification of citric acid (2-hydroxy-1,2,3propanetricarboxylic acid, herein (CA)) with ethanol and tributyl citrate, synthesized via esterification of CA with n-butanol, are such “green” solvents that are finding expanded uses as nontoxic plasticizers, medical products, printing ink coatings, cosmetics, etc. These biodegradable materials are also suitable as food additives, including whipping agents for dried egg whites, food flavorings, or food packaging materials. Citrate esters rapidly metabolize in the body via liver and blood serum enzymes to liberate the citrate ion, which is disposed through the usual biochemical pathways. Citric acid esterification with ethanol to form TEC proceeds sequentially through a series-parallel reaction scheme (Figure 1) involving monoethyl citrate (MEC) and diethyl citrate (DEC), as described in our recent paper1 characterizing CA esterification kinetics. Prior information on the kinetics of CA esterification with ethanol or n-butanol is confined mainly to the Chinese and German patent literature and is detailed in our earlier work.1 The esterification of CA is a thermodynamically limited reaction and, thus, proceeds only to partial completion in a conventional reactor. Continuous removal of one of the products of the reaction mixture is therefore required in order to drive the reaction to completion. We propose to do this using continuous reactive distillation. * To whom correspondence should be addressed. Tel.: 1-517-3533928. Fax: 1-517-432-1105. E-mail: [email protected].

Figure 1. Esterification of citric acid.

1.1. Overview of Reactive Distillation. Reactive distillation (RD) is a multifunctional process where chemical reaction and distillation occur simultaneously in a single vessel. The numerous advantages arising from the synergistic interaction of unit operations in a single unit over the conventional, sequential operation of a reactor followed by a distillation column include overcoming thermodynamic limitations of chemical reactions (as in methyl acetate and MTBE production), improving selectivity via rapid removal of products from the reaction zone to limit secondary conversion, “reacting away” azeotropes present in conventional distillation, utilizing the exothermic heat of reaction for vaporization, and avoiding hot spot formation. These advantages lead to lower capital and operating costs as well as reduced environmental impact. Shortcomings of reactive distillation include greater complexity in design and process control, often involving multiple steady states, and high nonlinearity due to the strong interaction between process variables. Moreover, the operating window of reactive distillation processes must be compatible with the thermal stability of the catalyst in order to ensure long catalyst

10.1021/ie070279t CCC: $40.75 © 2008 American Chemical Society Published on Web 01/23/2008

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life. Reactive distillation also requires a significant research and development effort in order to develop and successfully scaleup a promising process. Reactive distillation was first reported by Backhaus.2 The technique has witnessed expanded application over the last 20 years since its commercial application to MTBE and methyl acetate production. Reviews on reactive distillation are available from Sharma and Mahajani,3 Mahajani and Chopade,4 Taylor and Krishna,5,6 and Hiwale et al.7 Reactive distillation in the presence of solid acid catalysts, also commonly referred to as catalytic distillation, has come a long way since its inception by Spes.8 Industrially applied reactive distillation processes are rather limited in number currently, especially in comparison to potential applications.9 The degree of complexity in RD processes increases with use of solid catalysts, but the development of modern structured column packings makes the use of such catalysts viable. Katapak-S structured packing (Sulzer), used in this study, is a well-known and highly versatile packing consisting of vertically oriented enclosures filled with catalyst particles. Fluid dynamic properties and use of Katapak-S in reactive distillation systems have been described in the open literature by Moritz and Hasse,10 Hanika et al.,11 Van Baten et al.,12 Gotze et al.,13 Van Baten and Krishna,14 Smejkal et al.,15 Ratheesh and Kannan,16 Kolodziej et al.,17 and Schmitt et al.18,19 1.2. CA Esterification via Reactive Distillation. Synthesis of organic acid esters by reactive distillation is well-established, but in most applications, the ester has either the highest volatility of the reagents present (e.g., methyl acetate) or the lowest volatility, with water as the most volatile component.18 In these cases, recovery of 100% pure ester is straightforward via optimization of column operating conditions. Triethyl citrate (TEC) production via reactive distillation does not fit into either of these categories, since it has a volatility that is lower than ethanol and water but higher than CA (which is essentially nonvolatile). Therefore, it is only possible to isolate the pure product if complete conversion of CA and intermediate products MEC and DEC are achieved within the reactive distillation column. The primary challenge is therefore to achieve sufficiently rapid esterification kinetics so as to ensure complete conversion to the desired product TEC. Previous experimental work on similar esterification systems has been described by Bock et al.20 for the synthesis of isopropyl myristate and more recently from the authors’ laboratory by Asthana et al.21,22 for synthesis of ethyl lactate. In both of these systems, excess ethanol is used, which distills along with water as the top product. Omata et al.23 have described a reactive distillation system for synthesis of fatty esters where an immiscible twophase water-alcohol mixture distills as the top product. There has been no prior study on the application of reactive distillation for citrate esters formation other than our recent paper on citrate esterification kinetics in the presence of ion-exchange resin catalysts.1 Therefore, the present work has been carried out to develop a favorable reactive distillation configuration for high conversion of CA and high TEC yield. Experimental results are presented from a continuous pilot-scale reactive distillation system for CA esterification experiments operating at 1 bar pressure. Simulation of the experimental pilot-scale reactive distillation column to obtain high yield of TEC has been performed using the ASPEN Plus process simulation software. Effects of important design variables have been studied for the pilot scale reactive distillation column. Three process configurations have been presented for the plant scale design of a reactive distillation column.

2. Experimental Details 2.1 Materials. Anhydrous CA crystals were obtained from Aldrich Chemicals. Absolute ethanol (99% purity) and HPLC grade water were obtained from J. T. Baker. The strong acid cation exchange resin catalyst Amberlyst-15 (Rohm and Haas, Philadelphia, PA) was obtained in H+ form and was used without modification in the reactive distillation column. Purity of all chemicals was checked by gas chromatography or HPLC. 2.2 Analysis. The identities of CA and its ethyl esters (MEC, DEC, and TEC) in reaction mixtures were first confirmed by GC-MS analysis of their trimethylsilyl (TMS) derivatives. Citric acid (CA), MEC, DEC, and TEC were quantitatively analyzed on a Hewlett-Packard 1090 HPLC using a reversed phase C18 column (Novapak, 3.9 mm x 150 mm) held at 40 °C. Water/ acetonitrile (ACN) mixtures, buffered at pH ) 1.3, were used as the mobile phase (1.0 mL/min) in a gradient mode (0% ACN (t ) 0) to 60% ACN (t ) 20 min) to 90% ACN (t ) 25 min) to 0% ACN (t ) 28 min)), and species were quantified by UV detection (Hitachi L400H) at a wavelength of 210 nm. Citric acid (CA) and TEC were identified and quantified by comparing HPLC retention time and peak area with their respective calibration standard. Standards for MEC and DEC could not be obtained commercially. On a mass basis, the response factor values for CA and MEC were found to be same; therefore, MEC and DEC were each assigned the same response factor value as corresponding to TEC. Using these response factor values, the carbon balance for each reaction sample, based on CA and its esters, was in the range of 100 ( 10%. Ethanol and water from reaction samples were analyzed by gas chromatography (Varian 3700 with TCD detection; 20 mL/ min He as a carrier gas) using a packed stainless steel column (3.25 mm X 4 m) containing Porapak-Q as the stationary phase. The column temperature program involved initially holding at 140 °C for 2 min, heating to 220 °C at 20 °C/min, and holding at 220 °C for 6 min. 2.3. Reactive Distillation Column Description. Continuous reactive distillation experiments were performed in a pilot-scale Pyrex glass column with an inner diameter of 50 mm and total height of 5.5 m.22 The column consists of a 3.7 m long reactive section packed with Katapak-S structured packing elements (Sulzer Chemtech Ltd, Winterthur, Switzerland) containing approximately 76 kg/m3 column volume of 0.7 mm diameter Amberlyst-15 beads, a strong acid macroporous cation exchange resin. The nonreactive enriching section of height 0.8 m above the reactive zone and the nonreactive stripping section of height 1.0 m below the reactive zone contain empty Katamax structured packing (Koch-Glitsch, Ltd.). The column is equipped with an electronic reflux splitter to control reflux ratio, a total condenser with chiller capable of achieving a condenser temperature of -20 °C, and a reboiler with an overflow outlet to maintain a constant reboiler volume of 1.0 L and allow product withdrawal. Two pumps feed solutions to the column at controlled rates. One feed port is located toward the top of the column (F1) and the other port located toward the bottom of the column (F2). The column has ten ports along its length to allow internal temperature measurements, introduction of feed, and sample withdrawal. The columns are wrapped with electric heating tapes controlled by surface thermocouples and Omega controllers and further insulated using glass wool in order to obtain near-adiabatic operating conditions. 2.4. Reactive Distillation Column Operation. Since pure CA is a crystalline solid at room temperature, it was introduced to the column as a solution in ethanol. The acid feed, composed

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of 23 wt % CA in anhydrous ethanol, was fed near the top of the enriching zone (0.2 m from top of column), while preheated ethanol, either in liquid or vapor form, was fed 1 m above the reboiler at the bottom of the reactive zone. The reflux ratio (L/ D) was set to zero, although a small amount of internal reflux was noted experimentally (L/D < 0.05). The reboiler duty was held constant for all experiments. The column operating pressure was limited to 1 bar. In typical operation, the column was started by turning on the external heating tapes and reboiler heater, and starting the feed pumps at the specified feed rates. Steady state was generally achieved after a quantity of bottoms product equivalent to three reboiler volumes (or about 3.0 L) had been collected. Samples were then collected from the distillate and bottoms streams for product analysis, and steady-state flow rates of feed, bottoms, and distillate streams were measured by timed filling of graduated cylinders. In some instances, temperature profiles in the column were recorded. 3. Results and Discussion 3.1. Reactive Distillation Experiments. Four reactive distillation experiments are reported here for esterification of CA with ethanol. In batch kinetic experiments,1 we observed that CA esterification was relatively slow at 80 °C (the normal boiling point of ethanol). This kinetic limitation dictates that relatively low conversion of CA and low citrate ester yields can be expected in the experimental pilot-scale columnsa significant limitation in the reactive distillation experiments. In an initial esterification experiment (Run 1), the column was operated such that the reboiler temperature reached 235 °C, indicating that there was neither ethanol nor water in the reboiler. At these high temperatures, significant secondary reaction byproducts were formed that included citraconic acid.27 The HPLC analysis (not shown) of the reboiler indicated that more than 20 products were present, including nine in large quantities. We thus concluded that a feasible reactive distillation process for TEC formation requires enough ethanol in the reboiler to maintain a temperature that avoids secondary reactions. Ethanol in the bottoms stream can be easily separated from TEC product by simple distillation. Results of experimental runs 2-4 are presented in Table 1; compositions are taken directly from HPLC and GC analysis without normalization. In run 2, the reboiler heating rate was adjusted such that the reboiler liquid phase contained 29 wt % ethanol at steady state and reboiler temperature was below 110 °C. Figure 2 shows the HPLC analysis of the reboiler effluents no products of secondary reactions were detected. HPLC analysis of CA and its esters has not been reported previously and is thus illustrated in Figure 2sthe split peak for DEC represents the 1,2- and 1,3-diethyl esters, with the larger peak the 1-3 isomer. Run 2 appeared to achieve steady state after 16 h of operation, although we could not verify via multiple samples that steady state was achieved. From this run, we conclude that it is desirable to design and operate CA esterification such that the reboiler effluent contains approximately 30% ethanol. In run 3 (Table 1), carried out at significantly higher feed rates than those in run 2, steady-state operation was achieved. HPLC analysis of the reboiler effluent is very similar to that obtained in run 2, indicating that no secondary reaction products were formed. A high concentration of ethanol was observed in the reboiler effluent along with a lower conversion of CA than observed in run 2.

Table 1. Results of Pilot-Scale Reactive Distillation Experiments experiment

run 2a

citric acid feed (F1) citric acid conc. (wt %) 23 citric acid rate (mol/min) 0.0084 ethanol rate (mol/min) 0.11 temp (°C) 25 rate (mol/min) temp (°C) distillate Temp (°C) bottoms Temp (°C) CA conversion (%)

ethanol feed (F2) 0.34 78 78 87 85

run 3b

run 4c

23 0.023 0.34 25

23 0.023 0.34 25

0.32 78 78 82 41

0.32 84 78 91 61

Ethanol Water

distillate composition (wt %) 98.1 98.4 1.2 0.8

98.2 1.0

citric acid MEC DEC TEC ethanol water

bottoms composition (wt %) 9.6 13.5 30.2 9.8 30.8 2.7 8.0 0.3 29.0 73.6 0.0 0.0

18.2 24.6 14.0 2.2 40.4 0.0

a Steady state was not verified by multiple samples; results are for sample at 16 h of operation. b Steady state was achieved; four samples were collected between 6 and 10 h of operation (10 h shown). c Steady state was achieved; four samples were collected between 6 and 16 h of operation (16 h shown).

Figure 2. HPLC analysis of reboiler composition from run 2.

Run 4 (Table 1) was carried out at similar feed conditions to run 3, except that the ethanol feed was superheated to 84 °C. Because the results show the highest conversion to citrate esters, steady-state operation, and a reasonable quantity of ethanol in the bottoms stream, we have used pilot-scale experimental run 4 as a basis for determining catalyst performance parameters and for further process simulations. 3.2. Simulation of Pilot-Scale Reactive Distillation. Run 4 of the reactive distillation experiments was simulated using the RADFRAC module of the ASPEN Plus process modeling software. RADFRAC simulates reactive distillation by considering phase equilibria simultaneously with chemical reaction, assuming either that chemical equilibria is achieved on each stage or that reactions proceed via a specified kinetic rate. In the latter case, an estimate of liquid residence time or liquid holdup on each stage of the distillation column is required. Details of the RADFRAC algorithm are described by Venkataraman et al.28 It should be noted that ASPEN Plus is supported by a strong physical and chemical properties database, including hydrodynamics of structured column packings similar to the Katapak-S material used in our laboratory column. The design parameters used in the ASPEN Plus column model are shown in Table 2. The ASPEN Plus molecular library contains all species involved in this system except MEC and

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Table 2. Design Parameters Used in ASPEN Plus Simulations pilot-scale exp simulation feed ratio EtOH:citric acid

pilot-scale parametric study simulation (base case)

27.5

plant-scale simulation (scheme 1, RD column only)

19.3

14.6

70 2.6 0.0013 0.0053

70 2.6 4.7 19.6

78 2.7 0.019

78 2.7 49.04

top feed-F1

• temperature (°C) • pressure (bar) • citric acid feed rate (kmol/h) • ethanol feed rate (kmol/h)

25 1.0 0.0014 0.0197

• temperature (°C) • pressure (bar) • ethanol feed rate (kmol/h)

78 1.1 0.019

total number of stages (N) feed stages: F1 F2 column operating pressure (bar) column pressure drop (bar) reactive stages reflux ratio boilup ratio Murphree stage efficiency (stages 2 to N) liquid holdup (stage 2 to (N - 1)) (vol %)

10

60

120

On 2 Above 9 1.0 0.05 3 to 8 0.01 3.35 0.5 8

Above 2 On 58 2.5 0.1 3 to 58 0.01 5.8 0.5 8

Above 2 On 118 2.5 0.1 3 to 118 0.01 4.8 0.5 6

• type • height equivalent to theoretical stage (m) • fractional approach to maximum capacity heat of reaction

catalytic packing Kerapak 0.6 0.135 0

Kerapak 0.6 0.1 0

Kerapak 0.6 0.07 0

bottom feed-F2

column

Table 3. Pre-exponential Factors and Activation Energies for Rate Constants in Activity-Based Kinetic Model of Resin-Catalyzed Esterification Reactions in a Pilot-Scale Reactive Distillation Column1 self-catalyzed reaction

resin-catalyzed reaction

reaction

k°self (1/s)

EA (kJ/kmol)

k°cat (kgsol/(kgcat s))

EA (kJ/kmol)

CA + EtOH f MEC + H2O MEC + H2O f CA + EtOH MEC + EtOH f DEC + H2O DEC + H2O f MEC + EtOH DEC + EtOH f TEC + H2O TEC + H2O f DEC + EtOH 2 EtOH f DEE + H2O

8.4 E+6 1.3 E+6 9.8 E+6 3.6 E+6 5.0 E+6 1.3 E+6

70800 70800 72000 72000 72400 72400

2.5 E+8 3.9 E+7 1.4 E+9 5.2 E+8 1.5 E+7 4.0 E+6 2.5 E+8

76900 76900 83100 83100 73200 73200 102000

DEC, so these compounds were defined using the group contribution method. All necessary physicochemical properties used in the simulations were taken as the default values from ASPEN Plus. The catalytic packing for the simulation was chosen as Kerapak in ASPEN Plus, since Katapak-S is not available in the ASPEN Plus data base and the two packings are similar. Our experimental evaluation of CA esterification clearly showed that reaction is slow at column operating conditions and that solution behavior is significantly nonideal. Hence, we wrote and inserted into the ASPEN Plus simulation a subroutine incorporating the activity-based kinetic model for CA esterification, based on UNIFAC, that we developed in our earlier work.1 In addition to both resin-catalyzed and self-catalyzed esterification reactions, the kinetic model includes the formation of diethyl ether (DEE) from ethanol. The self-catalyzed reactions must be included because they constitute a significant fraction of the overall reaction taking place in the column. In the simulation of run 4 in the pilot-scale column, CA conversion was fitted by a single parameter defined as catalyst effectiveness29 (ηcat). The catalyst effectiveness accounts for incomplete catalyst wetting in the column and for mass transfer limitations that may arise. Catalyst effectiveness is multiplied by the preexponential factor of each rate constant determined in batch studies,1 where intrinsic reaction kinetics are observed, to give effective preexponential factors and thus rate constants

for the catalyst in the column. A value of ηcat ) 0.75, the explanation for which is given in the Supporting Information, was found to best describe the experimental results. The kinetic parameters used in the column simulation for both self-catalyzed reactions and Amberlyst 15 cation exchange resin catalyzed reactions are given in Table 3. Using these constants, results of the simulation are in reasonable agreement with the experimental data (Figure 3). The average deviation in species composition for CA and its esterification products is about 20%, an acceptable result considering the complexities of the reactive distillation process and the analytical challenges associated with the lack of a pure chemical standard for MEC and DEC. The validity of the simulation is further supported by good agreement between experimental and predicted reboiler and condenser temperatures. 4. Extended Pilot-Scale Column Simulation Using rate constants (Table 3) calculated from a catalyst effectiveness of 75% as determined from the pilot-scale simulation of run 4 (Table 1), the effect of various column design and operating parameters such as number of reactive stages, ethanol feed position, column pressure, reflux ratio, and boilup ratio on the performance of a pilot-scale reactive distillation column to achieve complete CA conversion and high TEC yield (defined as mole TEC formed per mole CA fed) was investigated.


Figure 5. Predicted effect of number of reactive stages on TEC yield. Conditions: reflux ratio 0.01; boilup ratio 5.8; column pressure 2.5 atm.

Figure 3. Comparison of simulated (model) and experimental stream compositions (in weight percent) and temperatures for run 4 of citric acid esterification in the pilot-scale reactive distillation column.

Figure 6. Predicted effect of ethanol feed stage position on TEC yield and ethanol content in reboiler. Conditions: 60 stage column with 56 reactive stages; reflux ratio 0.01; boilup ratio 5.8; column pressure 2.5 atm.

Figure 4. Liquid-phase composition profile for base case simulation of pilot-scale reactive distillation column.

Parameters for the base case simulation of this parametric study are given in the second column of Table 2. These parameters are different from those used for simulating the pilot-scale experimental results in two aspects: a 50% CA in ethanol solution was used (vs 24% in run 4) and the column pressure was taken as 2.5 bar to increase overall column temperature and thus esterification rate to the highest possible values while still avoiding secondary degradation reactions. Figure 4 shows the liquid phase weight fraction profile inside the column for the base case simulation. The ethanol concentration is high everywhere in the column because of the high molar excess of ethanol used; the increase in ethanol concentration at stage 58 reflects the ethanol feed location. Citric acid concentration decreases quite rapidly within the first few stages below its point of introduction, followed by a decline in MEC concentration. DEC concentration decreases very slowly, indicating that the conversion of DEC to TEC is the slowest reaction and is responsible for the large number of stages required to achieve high TEC yield. The temperature profile in the reactive section of the column is essentially constant at the ethanol bubble point temperature, exhibiting a slight minimum because of the introduction of less volatile CA at the top of the column and increased ester content at the bottom of the reaction zone. 4.1. Effect of the Number of Reactive Stages. The reactive stages for the parametric simulation start at stage 3 and go up

to stage N - 2, where N is the total number of stages in the column. Figure 5 shows the effect of increasing the number of reactive stages from 16 (in a 20 stage column) to 116 (in a 120 stage column) on TEC yield when the column is operated at 2.5 atm with a reflux ratio of 0.01 and a boilup ratio of 5.8. The reactive stages are assumed to have a Murphree stage efficiency of 0.5. There is a rapid increase in TEC yield as the number of stages is increased from 16 to 56 (in a 60 stage column), followed by a further marginal increase in yield as the number of stages is increased to 116 (in a 120 stage column). The increase in number of reactive stages increases the catalyst loading and hence the extent of the slow chemical reaction encountered in the present work. 4.2. Effect of Ethanol Feed Position. Figure 6 shows the effect of changing ethanol feed position from stage 40 to 58 in a 60 stage reactive distillation column having reactive stages from 3 to 58 and operating at 2.5 atm, reflux ratio of 0.01, and boilup ratio of 5.8. Since CA is not volatile, its feed position is kept at stage 2. As the ethanol feed position is lowered from 40 to 58, it is observed that TEC yield increases from 96.5 to 98.5 wt % with little change in the reboiler ethanol concentration. This leads to the conclusion that in order to obtain optimum performance for citrate ester formation, the ethanol feed should be positioned right at the bottom of the reactive zone. 4.3. Influence of Column Operating Pressure. The effect of pressure on the performance of the reactive distillation column has been studied in the pressure range of 1-4 bar absolute for a 60 stage column containing 56 reactive stages, operating at a reflux ratio of 0.01 and boilup ratio of 5.8. As pressure increases, so does the reactive zone temperature and therefore the reaction rates. As shown in Figure 7, the yield of TEC improves up to a pressure of 2.5 bar, above which higher diethyl ether (DEE)

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Figure 7. Predicted effect of column operating pressure on TEC yield, DEE content in distillate, and average reactive zone temperature (°C). Conditions: 60 stage column with 56 reactive stages; reflux ratio 0.01; boilup ratio 5.8.

Figure 9. Predicted effect of boilup ratio (V/B) on TEC yield, ethanol content in reboiler, and DEE content in distillate. Conditions: 60 stage column with 56 reactive stages; reflux ratio 0.01; column pressure 2.5 atm.

5. Simulation of Commercial-Scale TEC Production

Figure 8. Predicted effect of reflux ratio on TEC yield and ethanol content in reboiler. Conditions: 60 stage column with 56 reactive stages; boilup ratio 5.8; column pressure 2.5 atm.

concentrations are seen in the column. The enhanced formation of DEE results in higher water concentration, leading to reverse reaction (hydrolysis) that reduces TEC yield. The maximum achievable TEC yield at these conditions was 98.5 wt % at 2.5 bar. 4.4. Effect of Reflux Ratio. The influence of reflux ratio (L/D) over the range 0.01-0.5 on TEC yield and ethanol concentration in the reboiler was investigated for the otherwise base-case column conditions. It is seen in Figure 8 that TEC yield is highest at the lowest reflux ratio of 0.01 and decreases sharply for L/D > 0.2. A reflux ratio close to zero suggests that water removal out the top of the column is critical for effective column performance; in this mode, the column is essentially operating as a reactive stripping column. This mode of operation is feasible because the CA feed point is above stage 2, essentially at the top of the distillation column. The ethanol concentration in the reboiler increases nearly linearly with reflux ratio increasing from 0.01 to 0.5, as shown in Figure 8. 4.5. Influence of Boilup Ratio. The effect of changing boilup ratio (V/B) is given in Figure 9 for the base-case column. The yield of TEC is practically unchanged above a boilup ratio of 3, and as expected, the ethanol concentration in the reboiler decreases from 55.2 to 17.3 wt % as boilup ratio increases from 2 to 10. DEE in the distillate decreases from 7.4 to 3.5 wt % as boilup ratio increases from 2 to 10; increasing boilup ratio decreases the concentration of ethanol in the lower, higher temperature region of the column, therefore reducing the extent of formation of DEE.

Three process flow schemes have been proposed and analyzed to evaluate the feasibility of commercial scale production of 25 million lbs TEC/y. The three schemes are illustrated in Figure 10 and are described below. In all cases, the feed to the process is a 50 wt % solution of CA in ethanolsthis corresponds to the solubility limit of CA in ethanol at 70 °C. We have not included the premixing and heating tank for CA dissolution in ethanol in the proposed process scheme. A TEC yield >98.5% has been chosen as a design criterion in each of the three process flow schemes. Scheme 1. This process consists of a stand-alone reactive distillation column for TEC formation. Scheme 2. This process consists of a fixed-bed prereactor for initial conversion of CA and ethanol to TEC, followed by a reactive distillation column for completion of TEC formation. Scheme 3. This process consists of the same two unit operations as scheme 2, with the addition of a regular distillation column following the prereactor to remove product water as its azeotrope with ethanol. The bottoms stream from this distillation column is then fed to the reactive distillation column for completion of TEC formation. Table 2 (column 3) gives the column operating parameters for the simulation of the RD column in scheme 1. For scheme 1, the CA flow in the upper port of the column (F1) was initially fixed at 4.7 kmol/h and the ethanol flow rate in F1 was 19.6 kmol/h (Table 4, configuration 1). The initial ethanol flow in the lower port of the column (F2) was set at 49.04 kmol/h. Achieving a TEC yield of 98.5% is not possible using these process flow rates, even with 120 actual stages in the reactive distillation column and operation at 2.5 bar total pressure. We thus increased the ethanol feed rate F2 to 71 kmol/h (configuration 2 in Table 4) and achieved the desired TEC yield of >98.5%. To complete the process simulation, the bottom stream from the reactive distillation column of configuration 2 in scheme 1, which contains mainly ethanol and TEC, was fed to a simple distillation column (not shown in Figure 10) containing 14 stages and operating at 0.2 atm total pressure. The bottom stream from this TEC purification column contains 1.1 wt % DEC and 98.9 wt % TEC with negligible ethanol; ethanol in the distillate is recycled to the process. The distillate stream from the reactive distillation column was fed to a simple distillation column (not shown in Figure 10) containing 15 stages in order to separate DEE from the ethanol-water mixture. The distillate from this column contains 98.8 wt % DEE and along with 0.4 wt % ethanol and 0.8 wt % water. The bottoms stream from this


Figure 10. Different reactive distillation configurations for ASPEN Plus simulations. (1) Reactive distillation column with nonreactive rectifying, nonreactive stripping, and reactive middle section. (2) Plug flow prereactor followed by reactive distillation column. (3) Plug flow prereactor followed by a simple distillation column and reactive distillation column. Reactive sections are distinguished by shaded areas.

distillation column contains 2.94 wt % DEE in ethanol; this stream can be sent to an ethanol purification train for recycling. In process scheme 2, the CA-ethanol feed mixture initially enters a fixed-bed, plug flow prereactor. The goal of adding the prereactor is to allow the esterification reactions to approach equilibrium prior to the reactive distillation column, thus reducing column size. The prereactor in the simulation operates at 100 °C and 2.3 bar total pressure: at these conditions, a conversion close to the equilibrium value is obtained for a reactor space time of approximately 27 hsan admittedly large value that could be substantially reduced by increasing reactor temperature. The outlet from the prereactor is directed to the reactive distillation column operating at 2.5 bar pressure. As seen in Table 4, the number of stages required in the reactive distillation column for scheme 2 is 120, essentially the same value required for scheme 1. This is because the reactions of CA and MEC are rapid relative to the conversion of DEC to TEC; therefore, column size is almost entirely dictated by the kinetics of TEC formation from DEC. The bottom stream from the reactive distillation column is fed to a simple distillation column, where a purified TEC stream containing 1.5 wt % DEC and 98.5 wt % TEC was obtained. It is worth noting that the addition of the prereactor in scheme 2 has essentially the same effect as the higher ethanol F2 rate in Scheme 1sboth enhance TEC yield to a modest degree. For scheme 3 (configuration 1), a prereactor is used at the same conditions as in scheme 2, but the effluent from the prereactor is directed to a simple distillation column of 10 stages operating at 1.0 bar pressure. In this column, about 90% of the water produced in reaction is removed as its azeotrope with ethanol in the distillate stream. The bottom stream from this column is then fed to a reactive distillation column having 60 stages and operating at 1.6 bar pressure. The bottom stream from the reactive distillation column is directed to a simple distillation column (not shown in Figure 10) having 14 stages and operating at 0.2 bar pressure in order to separate ethanol from TEC. The bottom stream from this column contains 1.40 wt % DEC and 98.60 wt % TEC. In a variation of scheme 3, an azeotropic ethanol-water mixture (82 mol % ethanol, 18 mol % water) is used as the solvent for CA in F1 (configuration 2). With the same prereactor

and intermediate distillation column, an 80-stage reactive distillation column gives the desired TEC yield. The use of the azeotropic ethanol-water mixture significantly reduces processing costs, as less anhydrous ethanol would need to be produced via recycling. Use of the azeotropic ethanol-water mixture was not attempted for schemes 1 or 2, because the RD column size is already very large for those process configurations. Analysis of these commercial-scale configurations indicates that, if used alone, a large reactive distillation column (up to 120 stages) and excess ethanol is required to achieve high TEC yields from CA, mainly because the reaction of DEC to TEC is slow. The addition of a prereactor allows a reduction in the quantity of ethanol required to achieve the desired conversion, but does not allow the reactive distillation column size to be reduced. (We attempted to obtain column convergence in ASPEN Plus with columns containing more than 120 stages, but were unsuccessful.) The addition of a prereactor and an intermediate distillation column (scheme 3) to remove some of the product water considerably reduces the required size of the reactive distillation column, and is thus the preferred design configuration. The use of an azeotropic ethanol-water mixture as the initial feed solvent, which can be recycled from the intermediate distillation column distillate stream, would further reduce ethanol purification costs while only increasing the reactive distillation column size marginally. It can be seen from the present work that there are significant challenges associated with the synthesis of TEC using reactive distillation. The slow reaction rates of esterification, particularly in forming TEC from DEC, require a large number of reactive stages. The unfavorable thermodynamics of TEC formation1 require a large mole ratio of ethanol to CA, leading to large recycling streams and high ethanol purity costs. Formation of DEE results in additional separation requirements and loss of ethanol and is alleviated only by maintaining low temperatures which result in slower kinetics. Despite these limitations, reactive distillation is clearly a viable approach to overcome the thermodynamic limitations of citrate esters formation. As a point of reference, we have calculated the equilibrium composition obtained from an equilibrium reactor using the feed streams (4.7 kmol/h CA and 68.64 kmol/h ethanol) to the given process schemes in Figure 10 at

1024


Table 4. Comparison of Stream Compositions for Commercial-Scale Synthesis of TEC scheme 1

scheme 3

config. 1 config. 2 scheme 2 config. 1 config. 2 feed F1: citric acid (kmol/h) ethanol (kmol/h) water (kmol/h) feed F2: ethanol (kmol/h)

4.7

4.7

4.7

4.7

4.7

19.6 0.0 49.04

19.6 0.0 71.0

19.6 0.0 49.04

19.6 0.0 49.04

19.6 4.3 49.04

100 2.3 6 3

100 2.3 6 3

100 2.3 6 3

prereactor outlet composition (wt %) 2.6 2.6 16.5 16.5 N.A. N.A. 29.1 29.1 15.0 15.0 28.2 28.2 8.5 8.5

4.0 19.2 26.0 10.2 29.2 11.5

• operating tem (°C) • pressure (bar) N.A. • length (m) • diameter (m) CA MEC DEC TEC EtOH water • no. of stages • column pressure (bar) • feed stage • diameter (m) CA MEC DEC TEC EtOH water

TEC EtOH water DEE

N.A.

intermediate distillation column N.A 10 N.A. N.A. 1 above 5 0.70

10 1 above 5 0.66

intermediate distillation column bottom composition (wt %) 3.9 6.4 N.A. N.A. N.A. 24.6 30.8 43.5 41.7 22.4 16.4 1.0 0.5 4.6 4.3

• no. of stages • no. of reactive stages • column pressure • feed stages: feed F1 feed F2 • reflux ratio • boilup ratio • column diam (m) CA MEC DEC TEC EtOH water DEE

prereactor

reactive distillation column 120 120 120 116 116 116

60 56

80 76

2.5 above 2

2.5 above 2

2.5 above 2

1.6 above 2

1.6 above 2

on 118 0.01 4.8 2.23

on 118 0.01 6.3 2.55

on 118 0.01 4.1 2.07

on 58 0.01 3.0 1.90

on 78 0.01 3.0 1.90

reactive distillation column bottom composition (wt %) 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 1.8 0.8 1.1 1.0 69.5 70.4 70.3 70.4 28.7 28.8 28.6 28.6 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0

0.0 0.0 0.5 70.9 28.6 0.0 0.0

reactive distillation column distillate composition (wt %) 0.0 0.0 0.03 0.14 0.12 80.5 86.3 82.6 88.7 87.4 12.8 8.9 12.4 9.6 10.6 6.7 4.8 5.0 1.4 1.9

100 °C. The equilibrium values are 0.07 wt % CA, 1.53 wt % MEC, 9.93 wt % DEC, 18.9 wt % TEC, 64.34 wt % ethanol, and 5.26 wt % water. Thus a single reactor would give only about 60% yield of TEC, requiring further reaction and significant separation to achieve yields comparable to those found via reactive distillation. The three schemes proposed in Figure 10 represent only our initial, systematic attempts to develop an efficient process for citric acid esterification. Many other process concepts and configurations are possible to produce high TEC yields. Alternate schemes may include using reactive pumparounds

using either a reactive distillation column24 or a simple distillation column,25 pervaporation assisted esterification,26 the use of trays instead of structured packing in the reactive distillation columns, the use of semibatch reactive distillation, and the use of traditional multiple reaction/separation systems. Each of these process schemes, however, will have its own challenges, and it appears that the reactive distillation approach described here is viable and has advantages over other approaches. 6. Conclusions Reactive distillation can be used to esterify CA with ethanol to produce TEC. Both pilot-scale experimental studies and process simulation using ASPEN Plus software show that there are no intrinsic barriers to achieving yields of TEC close to the theoretical limit. The process is most strongly limited by the slow conversion rate of DEC to TEC. In the preferable design of a commercial-scale process that includes a prereactor, an intermediate distillation column for water removal, and a reactive distillation column (scheme 3), about 60 actual stages are required in the reactive distillation column in order to obtain a TEC yield above 98.5%. The maximum operating pressure of the column is limited by the operating temperature of the ion-exchange resin catalyst and by the formation of DEE, which becomes important as temperature rises above 120 °C. Further, ethanol must be carried into the bottoms stream of the RD column to maintain temperatures below about 150 °C, above which CA and its esters decompose. The reactive distillation column is best operated at very low reflux ratios (

Triethyl Citrate Synthesis by Reactive Distillation - Industrial

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