Novel Procedure for Coproduction of Ethyl Acetate and n-Butyl Acetate

Article pubs.acs.org/IECR

Novel Procedure for Coproduction of Ethyl Acetate and n-Butyl Acetate by Reactive Distillation Hui Tian, Huidong Zheng, Zhixian Huang, Ting Qiu, and Yanxiang Wu* College of Chemistry and Chemical Engineering, Fuzhou University, Fuzhou 350108, China S Supporting Information *

ABSTRACT: A novel energy saving procedure for the coproduction of ethyl acetate and n-butyl acetate by reactive distillation is simulated by Aspen Plus in this paper. Its feasibility is theoretically analyzed on the basis of the physical properties of ethyl acetate/n-butyl acetate system. The new procedure not only significantly reduces energy consumption, but also eliminates the acetic acid purification column, as compared with the individual processes. The outputs of ethyl acetate and n-butyl acetate can be adjusted flexibly by changing the feed molar ratio of ethanol to n-butanol in a certain range. When the feed molar ratio is less than 1.0, the purity of the products is the same as that for those obtained from the individual procedures. Also, energy-saving increases as the feed molar ratio of ethanol to n-butanol decreases. When the feed molar ratio of ethanol to n-butanol is 0.5, 20.40% energy-saving is achieved, as compared with the individual procedures. At last, the total annual cost (TAC) of the individual and coproduction procedure is calculated, and the TAC of the coproduction is 3 795 912 USD/Y (U.S. dollars per year) less than the individual’s 5 190 921 USD/Y, when 10 000 t EtAc and 26 363 t n-BuAc are produced. tained. Calvar5 investigate the reaction kinetics of the esterification of HAc with EtOH catalyzed by Amberlyst-15. A packed bed reactive distillation column filled with Amberlyst15 has been employed to obtain EtAc, and the influence of feed composition and reflux ratio have been analyzed. However, the concentration of EtAc is less than 60 wt %. Hanika6 studied the synthesis of n-BuAc by a pilot-plant RD column, and a computer simulation is performed to evaluate the experimental data. However, because of the low purity of the product, two additional distillation towers have to be used to purify n-BuAc. Based on computer simulation, Steinigeweg7 carried out a pilotplant experiment to study the RD production of n-BuAc, which use a strongly acidic ion-exchange resin (Amberlyst-15, Rohm and Haas Co.) as the catalyst, and the thermodynamic and kinetic of the system were also investigated; however, the purity of n-BuAc was only 96%. The processes of EtAc and n-BuAc have many disadvantages, such as high energy consumption, low conversion, byproducts, and so on. Therefore, a novel procedure for EtAc/n-BuAc coproduction is necessary to eliminate these disadvantages. In the past decade, increasing attention has been paid to reducing energy consumption. The aim of this study is to present a new energy-saving procedure for EtAc and n-BuAc coproduction by RD. At present, many factories produce EtAc and n-BuAc in several respective sectors, for example, British Petroleum Co. and Showa Denko K.K. The new procedure can be achieved on the existing equipment and the process configuration, by only a simple reconstruction. This new technology can produce EtAc and n-BuAc simultaneously, with a flexible production capacity, and significantly reduce the energy

1. INTRODUCTION The concept of reactive distillation (RD) was first reported by Backhaus in 1921.1 RD has become an attractive unit operation for the following reasons: (i) enhanced conversion and selectivity for the equilibrium reaction, (ii) breaking the composition of azeotropic mixtures, and (iii) reduced energy consumption and amount of equipment. 2 RD has received increasing attention, especially for equilibrium-limited and consecutive reactions. Ethyl acetate (EtAc) and n-butyl acetate (n-BuAc) are common solvents that are often used in the manufacture of lacquers and coatings and in other industries. Generally, they are produced industrially from ethanol (EtOH) or n-butanol (n-BuOH) with acetic acid (HAc), as shown in eqs 1 and 2. CH3COOH + CH3CH 2OH ↔ H 2O + CH3CH 2OOCH3 HAc

H 2O

EtOH

EtAc

(1)

CH3COOH + CH3CH 2CH 2CH 2OH HAc

n‐BuOH

↔ H 2O + CH3CH 2CH 2CH 2OOCH3 H 2O

n‐BuAc

(2)

There are many reports in literatures giving details of EtAc and n-BuAc synthesis by RD. By line diagrams and the validated theoretical predictions obtained through performing a set of RD experiments in a glass-tray column with 80 bubble cap trays, Kenig3 examined the feasibility of EtAc synthesis by RD. However, excess HAc had to be injected into the column to achieve a high conversion of EtOH, which increases the energy consumption of the system. Lai4 used a complex two-column section, which consisted of one RD column (10 reactive trays plus structured packing for separation) and one downstream stripper joined by a decanter in the middle, to produce high purity EtAc, and high-purity EtAc (up to 99.5 wt %) is ob© 2012 American Chemical Society

Received: Revised: Accepted: Published: 5535

September 20, 2011 March 1, 2012 March 29, 2012 March 29, 2012 dx.doi.org/10.1021/ie202154x | Ind. Eng. Chem. Res. 2012, 51, 5535−5541

Industrial & Engineering Chemistry Research

Article

column as the feed material in order to use energy efficiently in the new energy-saving coproduction procedure. In addition, as the most common solvent, many factories produce EtAc and n-BuAc in several respective sectors, and the new procedure could be achieved after simple reconstruction of the existing equipments and the process configuration. Based on the studies of RD for producing EtAc,3,15−18 the RD column has three basic elements: rectifying section on the top, reactive section in the middle, and stripping section at the bottom. Therefore, Figure 1a shows the process of individual production of EtAc, which was optimized in a previous investigation.19 HAc was fed at the top of the reactive section of column I at room temperature. EtOH was preheated to 348.15 K and then introduced to the bottom of the reactive section; the HAc/EtOH feed molar ratio was 1.5−3, and the reflux ratio was 2. The overhead vapor was condensed to ambient temperature and was introduced to a decanter for phase separation. The water formed by the reaction was removed from the phase separator. Because only a small amount of water with a high water content (>97 wt %) was removed from the phase separator as aqueous phase, the energy consumption of the rectification in this stream can be neglected. The major part of the organic phase was refluxed back to the column. The rest of the organic phase withdrew as the feed of the EtAc purification column II, and EtAc (>99 wt %) was obtained at the bottom of the column II. The bottom product of RD column I was HAc (approximately 90 wt %) and water, which was purified and fed back to RD column I as raw material. The HAc purification therefore increased the energy consumption in EtAc production. The flow-diagram for n-BuAc production is shown in Figure 1b, and the process was optimized in a previous investigation.20 HAc and n-BuOH were mixed at a molar ratio of 1 and then introduced to a fixed-bed reactor. The effluent from the fixed-bed reactor, which almost reached chemical equilibrium, was fed to the RD column II at the top of the reactive section. The overhead vapor of RD column II condensed to ambient temperature and was introduced to a decanter for phase separation. Most of the water in the column was removed from the phase separator, and the organic phase was refluxed back to the column. Because the mass fraction of water was more than 97 wt % in the water stream, the energy consumption of the rectification in this stream was neglected. The bottom products of RD column II were n-BuAc and very little water, which were introduced to column V for n-BuAc purification. n-BuAc (>99 wt %) would be obtained at the bottom of the column V. The flow-diagram for coproduction of EtAc and n-BuAc is shown in Figure 1c. RD column I was used to synthesize EtAc. The overhead vapor products were EtAc/ H2O/EtOH, EtAc/H2O, and H2O/EtOH azeotropic mixtures; excess HAc, as the high-boiling product, was obtained as the bottom product. The bottom product of column I was mixed with n-BuOH and then introduced to the fixed bed. The water formed by the reaction and entrained by HAc was removed from the top product in column IV, and the organic phase was refluxed back to the column IV. The outputs of EtAc and nBuAc could be adjusted by changing the feed molar ratio of EtOH to n-BuOH in the coproduction procedure, and the HAc purification column was eliminated compared with the individual processes.

consumption. Simulation for the new procedure is performed by Aspen Plus. The RADFRAC module, based on the rigorous equilibrium stage model, is used to describe the multistage vapor−liquid separation in the columns. The total annualized energy consumption and product purities obtained by the coproduction and individual procedures are compared in this paper.

2. FEASIBILITY ANALYSIS AND PROCESS DESCRIPTION The compositions of the azeotropes of EtAc and n-BuAc systems are shown in Tables 1 and 2. When EtAc is produced inTable 1. Compositions and Temperatures of the Azeotropes of EtAc System at Atmospheric Pressurea

a

azeotrope

H2O (mol %)

EtOH (mol %)

EtAc (mol %)

T (K)

EtOH-H2O EtAc-EtOH EtAc-H2O EtAc-EtOH-H2O

9.62 0.00 31.23 29.14

90.38 46.22 0.00 12.23

0.00 53.78 68.77 58.63

351.25 344.95 343.55 343.38

Data from refs 11−13.

Table 2. Compositions and Temperatures of the Azeotropes of n-BuAc System at Atmospheric Pressurea

a

azeotrope

n-BuOH (mol %)

n-BuAc (mol %)

H2O (mol %)

T (K)

n-BuAc-H2O n-BuOH-H2O n-BuOH-BuAc n-BuOH- n-BuAc- H2O n-BuOH- n-BuAc-HAc

0.00 24.76 78.73 4.78 20.32

28.87 0.00 21.27 24.01 22.04

71.13 75.24 0.00 71.21 0.00

364.34 366.11 389.30 363.81 394.45

Data from refs 7, 11, 14.

dustrially, excess HAc, as the highest-boiling component, is obtained as the bottom product; the boiling points of EtAc/H2O/ EtOH, EtAc/H2O, and H2O/EtOH azeotropes are similar, and these are obtained as the top product. In the esterification, when 1 mol ester is produced, so is 1 mol water. However, Table 1 shows that water accounts for 31% of the EtAc/H2O azeotrope. The reason is that if all of the water withdrew from the RD tower, a large amount of EtAc would reflux to the RD tower, which would increase the energy consumption. In previous investigations, when all the water withdrew from the top of the RD column, the conversion of the EtOH would decrease obviously. So, allowing some of the water withdraws from the bottom of the RD column is rational for the EtAc production. However, it decreases the concentration of HAc in the bottom product, and it increases the energy consumption in HAc purification. Because of the pinch point in the HAc/H2O system, the energy consumption in HAc purification is huge. The energy consumption is therefore still high when RD is used for EtAc production. However, Table 2 shows that water accounts for 71% of the n-BuAc/H2O azeotrope, and the ability of n-BuAc to include water is better than that of EtAc. The n-BuAc is used as the entrainer in the azeotropic distillation for HAc purification,8−10 which explains the n-BuAc have a better ability to take away water. Based on this, a new procedure for EtAc and n-BuAc coproduction is developed for saving energy. The bottom product of the EtAc system is injected into the n-BuAc RD

3. MODELING AND SIMULATION 3.1. Kinetic Model. Reactions 1 and 2 had been studied when they were catalyzed by Amberlyst 36-Wet ion-exchange 5536

dx.doi.org/10.1021/ie202154x | Ind. Eng. Chem. Res. 2012, 51, 5535−5541


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Figure 1. Flow-diagram of coproduction and individual productions of EtAc and n-BuAc: (a) individual EtAc production; (b) individual n-BuAc production; (c) EtAc and n-BuAc coproduction (I. EtAc RD column; II. EtAc purification column; III. HAc purification column; IV. n-BuAc RD column; V. n-BuAc purification column).

Table 3. Parameters of the Kinetic Model parameters system

k+ (L·mol−1 ·min−1·g−1)

k− (L·mol−1·min−1·g−1)

Ea+ (kJ·mol−1)

Ea− (kJ·mol−1)

EtAc system n-BuAc system

131 137 609 000

81 389 188 880

57.96 52.83

60.55 39.29

resin previously. The pseudohomogeneous model (eqs 3−5) has been proposed for the esterification maturely,14,21 as follows: r = ωcat(k+C HAcC EtOH(n‐BuOH) − k −C H2OC EtAc(n‐BuAc))

(3) 5537

⎛E ⎞ k+ = A+exp⎜ a + ⎟ ⎝ RT ⎠

(4)

⎛E ⎞ k − = A −exp⎜ a − ⎟ ⎝ RT ⎠

(5)



Article

where r is the reaction rate (mol·L−1·g−1·min−1), C is the concentration (mol·L−1), ωcat is catalyst weight (g), k is the kinetic constant (L·mol −1·min−1·g −1), A is preexponential factor ((L·mol−1·min−1·g−1), Ea is apparent activation energy (kJ·mol−1), and R is gas constant (J·mol·K−1). The parameters are given in Table 3.19,22 The chemical reaction is assumed to occur only in the liquid phase in the reactive section. Different catalyst concentrations and the efficiency of the catalyst in the RD column are considered, and the liquid holdup in the reactive

section is also determined and transformed into the simulated scale in the previous investigation.23 3.2. Thermodynamic Model and the Simulation. RD was an extremely complex process, which contained not only the mass transfer process but also the separation process and the reaction process. In this paper, the simulation of the synthesis of EtAc and n-BuAc in a RD column was carried out by Aspen Plus 11.1. The equilibrium stage model had been often applied with great success for the simulation of RD column.24 Therefore, the RADFRAC module, based on the rigorous equilibrium stage model containing the mass balance, phase equilibrium, summation and energy balance (MESH) equations, was used to describe the multistage vapor−liquid separation in the RD. To simplify the calculation, pressure drop along the column was neglected according to some research,25,26 and the whole equipment was supposed to be operated at atmospheric pressure. The selection of the physical property was very important due to the highly nonideal nature of the quaternary system, and the phase equilibrium of EtAc and n-BuAc systems was complex because of the existence of the azeotropes. Because of the nonideal vapor−liquid equilibrium (VLE) and possible vapor−liquid− liquid equilibrium (VLLE) for the quaternary system, the UNIQUAC activity coefficient model for n-BuAc system and the NRTL activity coefficient model for EtAc system were adopted,27 and the model equations and parameters came from the Aspen Plus database. The Hayden−O’Conell second virial coefficient model with association parameters was used to account for the dimerization and trimerization of HAc in the vapor phase.28 The binary interaction parameters came from the built-in Aspen Plus databanks. The azeotropes for EtAc and n-BuAc systems calculated at the atmospheric pressure are listed in Tables 4 and 5. The calculated azeotropic compositions and temperatures are compared with the literature values in Tables 1 and 2, and the maximal error is less than 3%.

Table 4. Calculated Results of the Compositions and Temperatures of Azeotropes at Atmospheric Pressure of EtAc System calculated data azeotrope

H2O (mol %)

EtOH (mol %)

EtAc (mol %)

T (K)

EtOH-H2O EtAc-EtOH EtAc-H2O EtAc-EtOH-H2O

9.85 0.00 31.23 28.89

90.15 45.87 0.00 14.23

0.00 54.13 68.77 56.88

351.33 344.58 343.52 343.10

Table 5. Calculated Results of the Compositions and Temperatures of Azeotropes at Atmospheric Pressure of n-BuAc System calculated data azeotrope

n-BuOH (mol %)

n-BuAc (mol %)

H2O (mol %)

T (K)

n-BuAc-H2O n-BuOH-H2O n-BuOH-BuAc n-BuOH- n-BuAc- H2O n-BuOH- n-BuAc-HAc

0.00 22.05 78.5 3.24 21.10

28.3 0.00 21.5 28.47 18.66

71.70 77.95 0.00 68.29 0.00

364.07 366.91 390.00 363.98 395.08

Table 6. Specifications of the Columns column total stages rectifying sect. reactive sect. stripping sect. diam. (m) liquid holdup m3/m3(reactive sect.)23 feed stage feed temp. (K) reflux ratio

I

II

III

IV

V

23 1−7 stages 8−13 stages 14−22 stages 1.367 0.237 HAc 8 EtOH 14 HAc 298.15 EtOH 348.15 2.0

27 2−12 stages

42 2−21 stages

22 2−11 stages

13−26 stages 0.563

22−41 stages 1.232

13

22

23 1−7 stages 8−13 stages 14−22 stages 0.832 0.237 8

12

298.15

383.15

358.15

393.15

1.2

3.0

12−21 stages 0.372

1.0

Table 7. Results of the Coproduction Simulation feed molar ratio of EtOH to n-BuOH stream H2O (wt %) EtOH (wt %) EtAc (wt %) HAc (wt %) n-BuOH (wt %) n-BuAc (wt %) conversion of EtOH conversion of n-BuOH

0.5

1.0 n-BuAc

EtAc 0.1 0.6 99.3

0.2 0.1 99.7 99.49% 99.73%

n-BuAc

EtAc 0.2 0.8 99.0

0.3 0.2 99.5 99.31% 99.62%

5538

2.0 n-BuAc

EtAc 0.1 0.8 99.1

0.3 0.2 99.5 99.27% 99.67%



Article

3 to 1.5), the EtOH conversion decreases slightly, because the reflux of the top product of column II increases the EtOH conversion, and the conversion of n-BuOH is not decreased as more water carried by the feed HAc stream. No EtAc and EtOH are taken to the column III, and the purity of product n-BuAc is desired. The conversion of EtOH and n-BuOH are both greater than 99%, and there is no EtOH and EtAc in the n-BuAc product, which all mean that the new energy saving procedure is feasible. 4.2. Energy Consumption. To compare the energy consumption, we define the RD columns for the individual and coproduction technologies have the same configurations and specifications, as shown in Table 6. Based on the simulation, the total annualized energy consumed by different units is calculated and listed in Table 8. The product purities of EtAc and n-BuAc are the same in both the individual procedure and coproduction procedure, as shown in Table 7. As the feed molar ratio of EtOH to n-BuOH decreases from 2.0 to 0.5 (the feed molar ratio of HAc to EtOH changes from 3 to 1.5), the energy-saving increases from 17.31% to 20.40%. The energy consumption significantly increases as the the feed molar ratio of EtOH to n-BuOH decreases, but the feed molar ratio of EtOH to n-BuOH is not better as smaller. As the the feed molar ratio of EtOH to n-BuOH decreases, the capacity of EtAc decreases, and the capacity of n-BuAc increases. However, the quantity demanded of EtAc in the market is usually greater than n-BuAc.29 So, the quantity demanded of the market should be considered, when the suitable feed molar ratio of EtOH to n-BuOH is decided. When 10 000 t EtAc is produced, so is 26 363 t n-BuAc, with the feed molar ratio of EtOH to n-BuOH 0.5, and it is acceptable both for the energy saving and the market demand. Under this operation condition, an energysavings of 20.40%, compared with the individual procedures, is achieved using the novel coproduction technology. Therefore, the new energy saving procedure not only saves energy but also saves the HAc purification column III compared with the individual processes, and the feed molar ratio of EtOH to n-BuOH can be adjusted with changes in the market demand.

Therefore, the selections of the thermodynamics model were reliable. The specifications of the RD columns are listed in Table 6, and the configuration of the RD column and the operating conditions are all optimized.19,20 EtOH and n-BuOH are all defined as the analytical reagents. The diameter of the columns is calculated when the production capacity of EtAc is 10 000 t/Y, with the feed molar ratio of EtOH to n-BuOH = 0.5.

4. DISCUSSION 4.1. Product Purities. The variations of product purities with the changing of feed molar ratio of EtOH to n-BuOH are shown in Table 7. The streams in Table 7 are marked in Figure 1c. As is well-known, the stoichiometric ratio of reactants significantly affects the reaction conversion and the products purities, so investigation of the feed molar ratio of EtOH to n-BuOH is necessary, especially for the new procedure. Based on the previous optimization, it is found that as the feed molar ratio of EtOH to n-BuOH increases from 0.5 to 2 (the feed molar ratio of HAc to EtOH changes from Table 8. Energy Consumption of Individual and Coproduction Technologies feed molar ratio of EtOH to n-BuOH EtAc (t/year) n-BuAc (t/year) individual column I/MW column II/MW column III/MW column IV/MW column V/MW total/MW coproduction column I/MW column II/MW column IV/MW column V/MW total/MW energy saving/%

0.5 10 000 26 363 3530 752 4926 11 395 605 21 208 3530 752 11 965 635 16 882 20.40

1.0 10 000 13 181 3957 791 2874 5698 303 13 623 3957 791 5983 318 11 049 18.89

2.0 10 000 6590 4366 812 1893 2849 151 10 071 4366 812 2991 159 8328 17.31

Table 9. TAC of the Individual and Coproduction Procedure individual

coproduction

TAC (USD/Y) column cost reboiler cost condenser cost catalyst cost paking cost steam cost cooling water cost total utility cost annualized total capital cost total TAC column cost reboiler cost condenser cost catalyst cost paking cost steam cost cooling water cost total utility cost annualized total capital cost total TAC

I 378 811 128 736 99 837 7289 37 982 515 355 3087 518 442 614 673 378 811 128 736 99 837 7289 37 982 515 355 3087 518 442 652 655

II 75 648 26 983 23 987

III 563 492 170 982 148 726

68 983 109 787 6289 116 076 126 618

52 987 719 161 3453 722 614 883 200 5 190 921

75 648 26 983 23 987

IV 140 731 148 726 110 983 2703 12 983 1 663 589 7367 1 670 956 403 143 140 731 148 726 110 983 2703 12 983 1 746 805 7523 1 754 328 416 126

68 983 109 787 6289 116 076 195 601

V 26 910 10 893 8734 3098 88 326 336 88 662 46 537 26 910 10 893 8734 3098 92 706 344 93 050 49 635

3 795 912 5539



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4.3. Economic Analysis. To compare the complete economic investment, the total annual cost (TAC) of the RD columns is calculated. The TAC includes the utility cost and the annualized capital cost. The utility cost includes steam for the reboiler, cooling water for the condenser, and catalyst cost. The capital cost includes the column shell, internal packing, reboiler, and condenser. A capital charge factor of 3 years is assumed in the calculation. The depreciation rate of the column, reboiler, condenser and packing is 10%, and the depreciation rate of the catalyst is 33.3%. The cost calculation can be seen from the work by Douglas.30 The TAC is calculated when 10 000 t EtAc is produced, so is 26 363 t n-BuAc with the feed molar ratio of EtOH to n-BuOH 0.5. The results of TAC are shown in Table 9. Because of the saving of column III, the TAC of the coproduction is 3 795 912 USD/Y (U.S. dollars per year) less than the individual’s 5 190 921 USD/Y, when 10 000 t EtAc and 26 363 t n-BuAc are produced. Therefore, the coproduction procedure has a better economic investment than the individual procedure.

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NOMENCLATURE RD = reactive distillation EtAc = ethyl acetate EtOH = ethanol n-BuAc = n-butyl acetate n-BuOH = n-butanol HAc = acetic acid TAC = total annualized cost (USD/Y) r = reaction rate (mol·L−1·g−1·min−1) C = concentration (mol·L1) ωcat = catalyst weight (g) k = kinetic constant (L·mol−1·min−1·g−1) A = preexponential factor ((L·mol−1·min−1· g−1) Ea = apparent activation energy (KJ·mol−1) R = gas constant (J·mol·K−1) REFERENCES

(1) Backhaus, A. Continuous Process for the Manufacture of Esters. U.S. Patent 1,400,849, 1921. (2) Baur, R.; Higler, A. P.; Taylor, R.; Krishna, R. Comparison of Equilibrium Stage and Nonequilibrium Stage Models for Reactive Distillation. Chem. Eng. J. 2000, 76, 387. (3) Kenig, E. Y.; Bider, H.; Gorak, A. B.; Bling, T.; Adrian, H. S. Investigation of Ethyl Acetate Reactive Distillation Process. Chem. Eng. Sci. 2001, 56, 6185. (4) Lai, I. K.; Liu, Y. C.; Yu, C. C.; Li, M. J.; Huang, H. P. Production of High-Purity Ethyl Acetate Using Reactive Distillation: Experimental and Start-up Procedure. Chem. Eng. Process. 2008, 47, 1831. (5) Calvar, N.; Gonzalez, B.; Dominguez, A. Esterification of Acetic Acid with Ethanol: Reaction Kinetics and Operation in a Packed Bed Reactive Distillation Column. Chem. Eng. Process. 2007, 46, 1317. (6) Hanika, J.; Kolena, J.; Smejkal, Q. Butyl Acetate via Reactive DistillationModeling and Experiment. Chem. Eng. Sci. 1999, 54, 5205. (7) Steinigeweg, S.; Gmehling, J. N-Butyl Acetate Synthesis via Reactive Distillation: Thermodynamic Aspects, Reaction Kinetics, Pilot-Plant Experiments, and Simulation Studies. Ind. Eng. Chem. Res. 2002, 41, 5483. (8) Donald, F. O. Acetic Acid Recovery Methods. Chem. Eng. Process. 1958, 54, 48. (9) Kirbaslar, I. S. Liquid−Liquid Equilibrium of the Water−Aceitic Acid−Butyl Acetate System. Braz. J. Chem. Eng. 2002, 19, 243. (10) Nagata, I.; Gotoh, K.; Tamura, K. Association Model of Fluids Phase Equilibrium and Excess Enthalpies in Acid Mixture. Fluid Phase Equilib. 1996, 124, 31. (11) Cheng, N. L. Handbook of Solvents; Chemical Industry Press: Bei Jing, 2007. (12) Nishith, V.; Prodromos, D. Dynamics and Control of an Ethyl Acetate Reactive Distillation Column. Ind. Eng. Chem. Res. 2001, 40, 833. (13) Calvar, N.; Domíngue, A.; Tojo, J. Vapor−Liquid Equilibria for the Quaternary Reactive System Ethyl Acetate + Ethanol +Water + Acetic Acid and Constituent Binary Systems Ethyl Acetate + Ethanol, Ethyl Acetate + Acetic Acid and Water + Acetic Acid at 101.3 kPa. Fluid Phase Equilib. 2005, 235, 215. (14) Hanika, J.; Kolena, J.; Smejkal, Q. Butyl Acetate via Reactive Distillation: Modelling and Experiment. Chem. Eng. Sci. 1999, 54, 5205. (15) Wu, K. C.; Lin, C. T. Catalytic Processes for the Preparation of Acetic Esters. U.S. Patent 5,998,658, 1999.

ASSOCIATED CONTENT

S Supporting Information *

Simulation details of EtAc/n-BuAc coproduction and relevant discussion. This information is available free of charge via the Internet at http://pubs.acs.org.

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ACKNOWLEDGMENTS

The authors acknowledge the financial support for this work from the National Natural Science Foundation of China (20676023).

5. CONCLUSIONS Because the water-including ability of n-BuAc is better than EtAc, a new procedure for the coproduction of EtAc and n-BuAc by RD is investigated and presented in this paper. The Aspen Plus simulator is employed to simulate the processes, and the RADFRAC module, based on the rigorous equilibrium stage model is adopted. The reaction kinetics of the synthesis of EtAc and n-BuAc are investigated in the previous investigation, and the kinetic constants for the pseudohomogeneous model based on activities are derived. In the new procedure, the outputs of EtAc and n-BuAc can be adjusted by changing the feed molar ratio of EtOH to n-BuOH. The influences of the feed molar ratio of EtOH to n-BuOH on product purities and energy consumption are investigated. When the feed molar ratio of EtOH to n-BuOH is less than 1.0, the product purities are the same as those obtained by using the individual technologies; this proves the feasibility of the coproduction technology. The energy-savings increases slightly as the feed molar ratio of EtOH to n-BuOH decreases. Compared with the individual technologies, the coproduction technology can save 20.40% energy when the feed molar ratio of EtOH to n-BuOH is 0.5. This novel technology can produce EtAc and n-BuAc simultaneously, with a flexible production capacity, and can significantly reduce energy consumption, and the HAc purification column can be saved. The TAC of the individual and coproduction procedure is calculated, and the TAC of the coproduction is 3795912 USD/Y less than the individual’s 5190921 USD/Y, when 10 000 t EtAc and 26 363 t n-BuAc are produced. Therefore, the coproduction procedure has a better economic investment than the individual procedure.

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AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Notes

The authors declare no competing financial interest. 5540



Article

(16) Klöker, M.; Kenig, E. Y.; Górak, A.; Markusse, A. P.; Kwant, G.; Moritz, P. Investigation of Different Column Configurations for the Ethyl Acetate Synthesis via Reactive Distillation. Chem. Eng. Process. 2004, 43, 791. (17) Reepmeyer, F.; Repke, J. U.; Wozny, G. Time Optimal Start-up Strategies for Reactive Distillation Columns. Chem. Eng. Sci. 2004, 59, 4339. (18) Kolena, J.; Lederer, J.; Moråvek, P.; Hanika, J.; Smejkal, Q.; Skåla, D. Method of Producing Ethyl Acetate and an Equipment for Carrying out this Method. U.S. Patent 6,693,213, 2004. (19) Huang, Z. X.; Tian, H.; Qiu, T.; Wu, Y. X. Simulation Studies of Reactive Distillation Processes for Synthesis of Ethyl Acetate. IEEE ICCSIT Meeting, Chengdu, July 2010; Paper 560. (20) Qiu, T.; Huang, Z. X.; Cheng, C. B.; Wu, Y. X. Synthesis of nButyl Acetate by Catalytic Distillation. Chem. React. Eng. Technol. 2008, 24, 545. (21) Pöpken, T.; Götze, L.; Gmehling, J. Reaction Kinetics and Chemical Equilibrium of Homogenously and Heterogeneously Catalyzed Acetic Acid Esterification with Methanol and Methyl Acetate Hydrolysis. Ind. Eng. Chem. Res. 2000, 39, 2601. (22) Qiu, T.; Huang, Z. X.; Cheng, C. B.; Wu, Y. X. Kinetics of Synthesis of n-Butyl Acetate over Cation-Exchange Resin Catalyst. Chem. React. Eng. Technol. 2009, 25, 354. (23) Chen, Y. B. Fu Zhou University, Fu Jian province at China, personal communication, 2007. (24) Lee, H. Y.; Huang, H. P.; Chen, I. L. Control of Reactive Distillation Process for Production of Ethyl Acetate. J. Process Control 2007, 17, 363. (25) Zheng, Y. X.; Flora, T. T. N.; Garry, L. R. Catalytic Distillation: A Three-Phase Nonequilibrium Model for the Simulation of the Aldol Condensation of Acetone. Ind. Eng. Chem. Res. 2001, 40, 534. (26) Almeida-Rivera, C. P.; Swinkels, P. L. J.; Grievink, J. Designing Reactive Distillation Processes: Present and Future. Comput. Chem. Eng. 2004, 28, 1997. (27) Gangadwala, J.; Radulescu, G.; Kienle, A.; Sundmacher, K. Computer Aided Design of Reactive Distillation Processes for the Treatment of Waste Waters Polluted with Acetic Acid. Comput. Chem. Eng. 2007, 31, 1535. (28) Li, C. S.; Daisuke, H.; Suzuki, K. N-Butyl Acetate Synthesis via Reactive Distillation: Thermodynamic Process Design. Comput. Appl. Chem. 2009, 1124, 1585. (29) Luyben, W. L.; Yu, C. C. Reactive Distillation Design and Control; Wiley-VCH: New York, 2008. (30) Douglas, J. M. Conceptual Design of Chemical Process; McGrawHill: New York, 1988.

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Novel Procedure for Coproduction of Ethyl Acetate and n-Butyl Acetate

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