Reactive-Distillation Process for Direct Hydration of Cyclohexene to

Article pubs.acs.org/IECR

Reactive-Distillation Process for Direct Hydration of Cyclohexene to Produce Cyclohexanol Bor-Chang Chen, Bor-Yih Yu, Yuan-Lin Lin, Hsiao-Ping Huang, and I-Lung Chien* Department of Chemical Engineering, National Taiwan University, Taipei 10617, Taiwan ABSTRACT: Cyclohexanol is an important precursor in synthesizing intermediates of Nylons and also plasticizers. In this paper, a plantwide reactive-distillation process for the production of cyclohexanol by direct hydration of cyclohexene with water has been studied. Design with excess water is needed to economically increase conversion of cyclohexene to 99.9%. The optimal design of the proposed flowsheet also makes use of the natural liquid−liquid splitting of the binary cyclohexanol−water azeotrope. Realistic fresh cyclohexene feed with impurity of cyclohexane was also considered. This inert is designed to leave the system in the organic outlet stream of a decanter on top of the reactive-distillation column. The reboiler duty requirement of the proposed design flowsheet is compared with existing processes in open literature. It is found that significant energy savings can be realized by using this design.

1. INTRODUCTION Cyclohexanol is an important precursor in synthesizing the intermediate of Nylon-6, and it is being produced more than 1 million tons per year worldwide. The traditional process is to hydrogenate benzene to produce cyclohexane, and then cyclohexane can be oxidized by air to form a mixture containing cyclohexanol and cyclohexanone.1,2 However, the drawbacks of this process are the low selectivity of cyclohexanol, the explosion risk, high energy requirement, and numerous sideproduct formations. The alternative routes to produce cyclohexanol are by either direct or indirect hydration of cyclohexene. There are research articles that have discussed the indirect hydration route by reacting cyclohexene with formic acid to form a reactive intermediate, cyclohexyl ester, and then hydrating this intermediate to produce cyclohexanol and formic acid.3,4 The formic acid can then be recycled and used in the first reaction. The overall process includes two reactive-distillation (RD) columns and two decanters. This process was shown to exhibit multiple steady-states which may hinder the operation and control of this system. The direct hydration reaction is simpler by reacting cyclohexene with water to produce cyclohexanol. Previous studies of the process via this route showed that the reaction rate is slow and the reaction is also limited by low equilibrium conversion of this reversible reaction.5 A common way to overcome the equilibrium conversion of a reversible reaction is to use reactive-distillation technology. A nice review paper6 and two books7,8 summarize well the technology and industrial applications of this important process intensification technology. Steyer et al.9 suggested using this technology to increase the cyclohexene conversion. However, because of design at only the stoichiometric feed ratio and with low mutual solubility between cyclohexene and water, the conversion of the process is still quite low. A more recent paper10 also studied this direct hydration system with stoichiometric feed ratio. To increase the catalyst effectiveness, a side reactor configuration is proposed and evaluated. However, the proposed configuration will have © XXXX American Chemical Society

problems handling the cyclohexene feed stream containing impurities of cyclohexane. This study aims to improve the process design of this direct hydration process by further considering the excess-water design for the RD system in order to further improve the cyclohexene conversion. With the excess water design, the separation system will also need to be carefully designed so that excess water can be recycled back to the reactive-distillation column. An overall optimized reactive-distillation process will be established in this paper with a realistic cyclohexene feed having impurities (cyclohexane). Comparisons of total reboiler duties with existing cyclohexanol processes will also be made to demonstrate the superiority of the proposed design.

2. REACTION KINETICS AND PHASE EQUILIBRIUM The reaction of direct hydration of cyclohexene (ENE) with water to produce cyclohexanol (NOL) is showed below:

This reaction is exothermic, heterogeneous, and catalyzed by solid catalyst Amberlyst 15 with density of 770 kg/cm. The kinetic expression is in Langmuir−Hinshelwood form which is shown below: ⎛ ⎞ ene H 2O K ads K ads ⎟ r = ⎜⎜mcatk het H 2O ene nol 2 ⎟ (1 + aeneK ads + a H2OK ads + anolK ads ) ⎠ ⎝ ⎛ ⎞ ⎜⎜aenea H O − 1 anol ⎟⎟ 2 Keq ⎝ ⎠

(2)

Received: December 14, 2013 Revised: March 31, 2014 Accepted: April 5, 2014

A

dx.doi.org/10.1021/ie404226m | Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

Industrial & Engineering Chemistry Research

Article

The reaction rate is activity-based and the kinetics parameters are obtained from the literature (Steyer and Sundmacher11) as below. ⎡ −93687 ⎤⎛ kmol ⎞ ⎜ ⎟ = 7.7083 × 109 exp⎢ ⎣ RT ⎥⎦⎜⎝ kg s ⎟⎠ cat

(3)

⎡ ⎛1 1 ⎞⎤ ⎟⎥ Keq = 4.2907 exp⎢3389.38⎜ − ⎝T ⎣ 298.15 ⎠⎦

(4)

k het

Table 2. Boiling Point and Azeotropic Temperature Ranking at Atmospheric Pressurea NRTL data

The adsorption constant Kads, for H2O, ENE, and NOL are at 19.989, 0.056839, and 0.77324, respectively (cf. Steyer and Sundmacher11). Because there is liquid−liquid split at the reactive-distillation column trays, this reaction expression is assumed to occurr at both liquid phases. Since the catalyst Amberlyst 15 can only be operated below 120 °C, there is also a high-temperature constraint in the reaction section in the following simulations. As for predicting the phase equilibriums of this system, there are three components in the direct hydration reaction. However, since ENE and the inert cyclohexane (ANE) are difficult to be separated, some amount of cyclohexane is realistically considered as an impurity in the cyclohexene feed stream. Therefore, there are a total of four components in this process. The NRTL (nonrandom-two-liquid) thermodynamic model is used for modeling the vapor−liquid and vapor− liquid−liquid equilibrium in the system. The model parameters are obtained from the literature12 and are shown in Table 1.

a

cyclohexene

comp. j aij aji bij (K) bji (K) cij comp. i

cyclohexene

cyclohexene

cyclohexanol

water

cyclohexane

0 0 429.205 0.115 809 0.802 522 cyclohexanol

0 0 1705.00 2609.45 0.267 206 cyclohexanol

0 0 5.109 61 7331.87 0.831 053 cyclohexane

comp. j

water

cyclohexane

water

aij aji bij (K) bji (K) cij

0 0 160.782 1318.19 0.359 706

0 0 2.397 65 489.733 0.993 301

0 0 2122.93 3012.81 0.258 799

component

T (°C)

composition

T (°C)

composition

H2O/ANE H2O/ENE ANE ENE H2O/NOL H2O NOL

69.41 70.71 80.78 82.88 98.74 100.02 160.84

(0.30, 0.70) (0.32, 0.68)

69.60 70.80 80.74 82.90 98.10 100.03 161.10

(0.30, 0.70) (0.31, 0.69)

(0.94, 0.06)

(0.93, 0.07)

All azeotropes are heterogeneous.

3. PROPOSED DESIGN FLOWSHEET It seems that the conceptual design of this reactive-distillation process is rather simple. However, because of low mutual solubility between the two reactants (cyclohexene and water), the conversion of the reactant is low for the design with stoichiometric feed ratio. One way to increase the conversion of the reaction is to improve the solubility of cyclohexene and water by introducing a new cosolvent component. There are several cosolvents proposed in the open literature.13−16 However, adding a new component will further complicate the separation problem. For example, Qiu et al.17 proposed a reactive-distillation flowsheet for this system by adding a cosolvent of 1,4-dioxane. The recycling of this cosolvent into the reactive-distillation column was designed via the organic phase from the top decanter. However, with realistic feed assumption of containing inert (cyclohexane) in the feed stream, this design flowsheet will not work because the organic phase should be drawn out of the system to prevent inert accumulation inside the system. An alternative design could be to only recycle portions of the organic phase back to the RD column and to purge the rest to avoid cyclohexane accumulation. However, the flow rate of the purge stream needs to be quite high because of significant amounts of cyclohexane contained in the feed stream. This will result in large loss of the cosolvent. Another alternative way to increase the conversion of a reaction is to feed one reactant in excess so that the other reactant can be more completely consumed. It is trivial to choose water as the excess reactant because of the low price and more importantly because it is easier to separate from cyclohexanol. With excess water feed, the conceptual design of the process is to have the excess water to go out from the bottom of the reactive-distillation column. Water will not go out of the system from the top of the reactive-distillation column because all the aqueous phase in the decanter is designed to recycle back to the reactive-distillation column for further reaction. The separation of the cyclohexanol−water mixture at the column bottom is easy because the binary azeotrope is heterogeneous (see Table 2 and Figure 1). From the discussion in Wu et al.,18 binary azeotrope can easily be separated with one decanter and two strippers. Since the composition of the aqueous phase is already quite pure, one of the strippers is even not needed in the proposed design flowsheet. The resulting optimal design of the proposed flowsheet is shown in Figure 2. In the design flowsheet, a decanter configuration on top of the reactive-distillation column commonly used in industry is adopted in the system. From

Table 1. NRTL Model Parameters of This System comp. i

experimental data

The boiling point and azeotropic temperature ranking at atmospheric pressure are shown in Table 2. There are multiple azeotropes in the system. The heaviest and the lightest components are the product cyclohexanol and the azeotrope of water and cyclohexane, respectively. The residual curve map and liquid−liquid envelope at 40 °C are shown in Figure 1. From Table 2 and Figure 1, the product (NOL) can be obtained from the bottoms of the reactive-distillation column. The top product should approach the azeotrope of water and cyclohexane. From Figure 1, it is shown that, in this azeotrope after condensing in a decanter, a natural liquid−liquid separation will occur. The aqueous phase (water reactant) should be recycled back to the RD. The organic phase will contain mostly the impurity (ANE) which the system should be designed to draw out to prevent impurity accumulation inside the system. B



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Figure 1. RCM and LLE of the studied four-component system.

Figure 2. Proposed design flowsheet for the production of cyclohexanol via the direct hydration route.

the figure, it is illustrated that the reactive-distillation column is designed to be the same as a regular distillation column with top the part of the column including a total condenser and a reflux stream. The liquid distillate product is then naturally separated into two liquid phases in a decanter with the aqueous phase totally recycled and the organic phase totally drawn out of the system. For the bottom stream of the reactive-distillation column, since cyclohexanol and water contain a minimum-boiling heterogeneous azeotrope, another decanter can be placed at the column bottom with the aqueous phase recycled back to the RD column. The organic phase can further be purified to obtain pure cyclohexanol at the bottom of a stripper. The top stream of the stripper can be recycled back to the bottom decanter.

In all the design studies followed, the software package of Aspen Plus was used to conduct the rigorous simulation. RADFRAC module with rigorous mass component and energy balance in each tray was used to represent the RD column as well as the stripper. In each tray of the reactive section in the RD column, the kinetic expression shown in the previous section 2 can be easily put in via a FORTRAN subroutine to link with the RADFRAC module. The VLE or VLLE calculations were also performed in each tray of the RD column and the stripper. Detection of possible liquid-splitting in each tray was built in for the RADFRAC module using the Gibbs’ tangent plane analysis method by Michelson.19 There are six variables in the flowsheet needed to be determined including: water/ENE feed ratio, the pressure of RD column, number of stages for the reaction section (Nrxn), the rectifying section (Nrec) and the stripping section (Nstr) of C



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Figure 3. Reactive-distillation flowsheet with stoichiometric feed ratio.

equipment. For the differential revenue calculation, the unit price of the NOL product was assumed to be $1.52/kg. An exhaustively iterative sequential optimization procedure was used to find the optimal design of the proposed flowsheet by minimizing TAC. Although time-consuming, all combinations of the design variables were investigated to obtain the minimization of TAC. The main reason for not using an automated optimization procedure (such as the one provided in Aspen Plus) is because of often-encountered convergence problems for simulation runs. In our experience, especially for the flowsheet with multiple recycle streams, we may need to disconnect the recycle stream(s) first to get the flowsheet converged and then manually connect the recycle streams oneby-one to obtain the results from a run with no convergence problems. The optimization procedure is described in the following steps: (1) Guess the water/ENE feed ratio. (2) Guess RD column pressure (3) Guess Nrxn. (4) Guess Nrec and Nstr. (5) Gauss NT2. (6) Change reflux ratio and reboiler duty of the RD column and also the reboiler duty of the stripper until three product specifications can be met. The three specs are the conversion of ENE is set at 99.9%; the ANE purity is set to be higher than 99.5%; and the product NOL purity is set at 99.9 mol %. (7) Go back to (5) and change NT2 until the TAC is minimized. (8) Go back to (4) and change Nrec and Nstr until the TAC is minimized. (9) Go back to (3) and change Nrxn until the TAC is minimized. (10) Go back to (2) and increase the RD column pressure until the maximum temperature in reactive section of the RD column still does not exceed 120 °C. (11) Go back to (1) and change the water/ENE feed ratio until the TAC is minimized.

the RD column, and also the total stages of the stripper (NT2). The heavier water and the lighter cyclohexene/cyclohexane feed streams are assumed to be fed at the top and bottom of the reaction section, respectively. The fresh feed of lighter cyclohexene/cyclohexane mixture is assumed to be 80 mol % ENE and 20 mol % impurity ANE with the flow rate of 100 kmol/h and as saturated liquid feed. Because the production rate will vary among various cases, the differential revenue to a base case from selling the product will be included in the total annual cost (TAC) calculation for fair comparison. The TAC is defined below: TAC = operating cost +

capital cost payback period

+ (revenue base case − revenueeach case)

(5)

The operating cost includes the steam for the two reboilers, the cooling water for the condensers, and the catalyst cost in the RD column. For the steam used in two reboilers, lowpressure steam (50 psig at 147 °C) is used in the RD column, and medium-pressure steam (150 psig, 185 °C) is used in the stripper. The unit prices of the low-pressure steam at $6.60/ 1000 kg, medium-pressure steam at $10.50/1000 kg, and cooling water at $0.02/m3 were adopted from Table 23.1 of Seider et al.’s book.20 As for the catalyst volume in each tray of the reactive section for the RD column, it is assumed to occupy half of the total holdup. The total holdup for each reactive tray can be calculated from the column diameter (from Aspen traysizing) and height. We assumed the active area calculated from the column diameter is discounted to 90% to allow for downcomer, etc. Catalyst cost is assumed to be $1.59/kg with its density assumed to be 770 kg/m3.21 The capital cost includes the column shell and trays for the RD column and the stripper, reboilers, and condensers. The formulas for the installed capital costs of all process equipments can be found on pages 572−575 in Appendix D of Douglas’ book.22 In the formulas, the Marshall & Swift equipment cost index was selected as 1533.3. The payback period was assumed to be 3 years for calculating the annual costs of all installed D



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The hierarchy of the iterative optimization procedure was arranged so that the design variable for the outermost iterative loop is the most sensitive one in terms of TAC changes, and the design variable for the innermost iterative loop affects the TAC the least. Step 7 is the innermost iterative loop in the overall procedure. A new optimum NT2 will be searched after changing Nrec and Nstr in the second innermost iterative loop. The resulting optimal water/ENE feed ratio is 2.0. Below this ratio, the conversion of ENE cannot be higher than 99.0 mol %. The optimal operating pressure of the RD column is at 3 atm. Above this pressure, the high temperature constraint of 120 °C will be violated in the reaction section of the reactive-distillation column. As can be seen in Figure 2, the RD column has 27 stages and two feed streams are fed at stages 5 and 21, respectively. For the stripper, the optimal total number of stages is five stages. The Aspen Plus simulation file for this optimized case is available to interested readers upon request. Note that the recycle stream from the top of the stripper back to the decanter contains more significant amounts of ENE (mole fraction of 0.217) although the mole fraction of ENE at the RD column bottom stream is only at 0.00014. This is because of the recirculation loop at the top of the stripper causing small amounts of light ENE at the RD bottom stream to accumulate into a more significant balanced amount in the recycle loop. An alternative design is to purge out some materials from the top of the stripper. However, any loss of the cyclohexanol product from this purge stream will cause a significant increase in the TAC at the differential revenue term and will not be economically favorable.

Figure 4. Conceptual design flowsheet of the Asahi process.

phase going to a distillation column. The cyclohexanol product is obtained from the bottom of this distillation column. The unreacted cyclohexene is recycled back to the reactor section from the top of this column. The Asahi process was chosen in Imam et al.23 as a benchmark reference case to evaluate various developed process concepts. Since the design details of the optimum Asahi process are not available in open literature, Imam et al.23 made several assumptions to estimate the energy potential of design alternatives as compared to this benchmark process. The feed to the distillation column was assumed to contain 12 wt % cyclohexanol, and a large number of stages (100) were also assumed. The benchmark condition is given in Table 3 of Imam et al.23 for a cyclohexanol product purity of 99% with more than 99% recovery. The resulting energy requirement for a production rate of 100 kilotons/year was 7.5 MW. Note that the distillation column in the case of Imam et al.23 only included cyclohexanol/cyclohexene separation. In reality, the organic phase may also contain some water. Also in reality, the feed should contain some impurity of cyclohexane, and this impurity should also be designed to leave the system. The above two considerations will make the separation more complex. The other remark is that the actual total stages of the column may not be this high to trade-off the capital cost with the operating cost. This will make the energy requirement even higher than 7.5 MW for the production rate of 100 kilotons/ year. Since the production rate of our proposed design is not the same as 100 kilotons/year, the proposed design flowsheet in Figure 2 was scaled up so that direct comparison of the reboiler duty requirements can be made. A new steady-state simulation is shown in Figure 5 by keeping the same design variables as in Figure 2 and to let the production rate be exactly the same as the benchmark condition in Imam et al.22 The production rate of 103.55 kmol/h in Figure 5 corresponds directly to 100 kiloton/year in the Asahi benchmark process. Table 3 shows the direct comparison of the proposed flowsheet in Figure 5 to the benchmark Asahi process. It is found that the savings in the total reboiler duty of 46.76% can be made by using the proposed design flowsheet. 4.3. Indirect Hydration Route. The limited information on an optimal indirect hydration process from the literature (Katariya et al.4) is shown in Figure 6. There are two RD columns and two top decanters. At the first RD column, cyclohexene reacts with formic acid to produce the intermediate cyclohexyl ester, and then at second RD column, this intermediate reacts with water to form the product cyclohexanol. The operating pressures of the two RD columns were set to below atmospheric pressure to prevent formic acid decomposition at a higher operating temperature.

4. COMPARISONS TO OTHER DESIGN ALTERNATIVES 4.1. Direct Hydration with Stoichiometric Feed. Process simulation of the RD column with a stoichiometric feed ratio was also conducted. For this case, the bottom product of the RD column is designed to obtain NOL. However, high conversion of ENE cannot be achieved with even a very high number of reactive trays. The flowsheet with highest obtainable conversion is shown in Figure 3. The specifications are the purity with composition of 99.0 mol % for both the top ANE stream and the bottom NOL product stream. The conversion of ENE is lower at 99.4%. The reflux ratio, the reboiler duty, and the column stages are all much higher than that in Figure 2. Thus, with the stoichiometric feed ratio, the economics of this alternative design flowsheet are much inferior to the one in Figure 2. There are also two issues that need to be mentioned which make the simulation in Figure 3 not realistic. First, the column needs to be operated at 5 atm for the purpose of increasing the reaction rate. At this higher pressure, the temperature of the reaction section is actually over the temperature constraint of the catalyst. Second, the reaction holdup in each reactive tray has to be set at a very large value which is inconsistent with the tray sizing results obtained by Aspen Plus. 4.2. ASAHI PROCESS Mitsui and Fukuoka5 developed a conventional process for the production of cyclohexanol via direct hydration of cyclohexene, utilizing a zeolite catalyst of type HZSM-5. The conceptual design of this process is illustrated in Figure 4. A very large amount of catalyst must be employed in a slurry-type reactor. The separator as shown in Figure 4 is actually a decanter with the aqueous phase recycled back to the reactor and the organic E



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Figure 5. Proposed design flowsheet with production rate the same as in Imam et al.23

comparison of the reboiler duty requirements can be made. Figure 7 displays the simulation result of the proposed design flowsheet with exactly the same feed condition as in Katariya et al.4 Again, design variables are kept the same as the ones in Figure 2. Notice that the feed temperatures of two fresh feed streams in Figure 7 are lower at 20.85 and 41.0 °C, respectively for ENE/ANE and water feed streams. This is to simulate the exact feed condition as that in Katariya et al.4 Notice also that the two column diameters are extremely small because of very low feed flow rates selected in their study. The results of the comparison can be found in Table 4. It is found that savings in the total reboiler duty of 32.86% can be made by using the proposed design flowsheet as compared to the condition in Katariya et al.4 To be fair for the comparison, the two bottom temperatures in Figure 6 are lower than the two bottom temperatures in our proposed design. Therefore, there may be a steam grade difference for the required steam letting the savings in steam cost not be as high as 32.86% for the proposed design flowsheet.

Table 3. Comparison between Asahi Benchmark Process and Proposed Design Flowsheet (production rate =100 kt/y) Asahi benchmark processa (Figure 4) product column NOL product purity number of stages reboiler duty (MW) total duty (MW) comparison (%) a

0.990 100 7.50 7.50 −

proposed design flowsheet (Figure 5)

−

RD column

stripper

− − −

− 27 2.363

0.999 5 1.630

3.993 −46.76%

Data from Table 3 of Imam et al.23

Various feed compositions with a cyclohexane impurity were studied. The one close to our feed composition was the one with a cyclohexene feed flow rate of 2.67460 × 10−3 mol/s with additional 20% cyclohexane entering the system with the cyclohexene feed. Since the flow rate of the feed is not the same between the cases in Figures 2 and 6, the proposed design flowsheet in Figure 2 was scaled-down so that direct

Figure 6. Conceptual design flowsheet via indirect hydration route. F



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Figure 7. Proposed design flowsheet with feed condition the same as in Katariya et al.4

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Table 4. Comparison between Indirect Hydration Process and Proposed Design Flowsheet (cyclohexene feed rate = 2.6746 × 10−3 mol/s) indirect hydration processa (Figure 6)

NOL product purity number of stages reboiler duty (W) total duty (W) comparison (%) a

− 20 130

(1) Fisher, W. B.; VanPeppen, J. F. Cyclohexanol and Cyclohexanon. In Kirk-Othmer Encyclopedia of Chemical Technology. John Wiley & Sons, Inc.: New York, 2000. (2) Musser, M. T. Cyclohexanol and Cyclohexanone. In Ullmann’s Encyclopedia of Industrial Chemistry. Wiley-VCH Verlag GmbH & Co. KGaA: Weinheim, Germany, 2000. (3) Steyer, F.; Freund, H.; Sundmacher, K. A Novel Reactive Distillation Process for the Indirect Hydration of Cyclohexene to Cyclohexanol Using a Reactive Entrainer. Ind. Eng. Chem. Res. 2008, 47, 9581−9587. (4) Katariya, A.; Hannsjorg, F.; Sundmacher, K. Two-step Reactive Distillation Process for Cyclohexanol Production from Cyclohexene. Ind. Eng. Chem. Res. 2009, 48, 9534−9545. (5) Mitsui, O.; Fukuoka, Y. Process for Producing Cyclic Alcohol. U.S. Patent 4,588,846, 1986, Asahi Kasei Kogyo Kabushiki Kaisha. (6) Malone, M. F.; Doherty, M. F. Reactive Distillation. Ind. Eng. Chem. Res. 2000, 39, 3953−3957. (7) Sundmacher, K., Kienle, A., Eds. Reactive Distillation: Status and Future Directions; Wiley-VCH Verlag CmbH & Co. KgaA: Weinheim, Germany, 2003. (8) Luyben, W. L.; Yu, C. C. Reactive Distillation Design and Control. John Wiley & Sons, Inc.: Hoboken, NJ, 2008. (9) Steyer, F.; Qi, Z. W.; Sundmacher, K. Synthesis of Cyclohexanol by Three-phase Reactive Distillation: Influence of Kinetics on Phase Equilibria. Chem. Eng. Sci. 2002, 57, 1511−1520. (10) Ye, J.; Li, J.; Sha, Y.; Lin, H.; Zhou, D. Evaluation of Reactive Distillation and Side Reactor Configuration for Direct Hydration of Cyclohexene to Cyclohexanol. Ind. Eng. Chem. Res. 2014, 53, 1461− 1469. (11) Steyer, F.; Sundmacher, K. Cyclohexanol Production via Esterification of Cyclohexene with Formic Acid and Subsequent Hydration of the Ester-reaction Kinetics. Ind. Eng. Chem. Res. 2007, 46, 1099−1104. (12) Steyer, F.; Sundmacher, K. VLE and LLE data for the System Cyclohexane + Cyclohexene + Water + Cyclohexanol + Formic Acid + Formic Acid Cyclohexyl Ester. J. Chem. Eng. Data 2005, 50, 1277− 1282. (13) Panneman, H. J.; Beenackers, A. A. C. M. Solvent Effects on the Hydration of Cyclohexene Catalyzed by a Strong Acid Ion-exchange Resin. 1. Solubility of Cyclohexene in Aqueous Sulfolane Mixtures. Ind. Eng. Chem. Res. 1992, 31, 1227−1231. (14) Panneman, H. J.; Beenackers, A. A. C. M. Solvent Effects on the Hydration of Cyclohexene Catalyzed by a Strong Acid Ion-exchange

proposed design flowsheet (Figure 7)

RD column 1 RD column 2 0.9927 20 502.1 632.1 −

RD column

stripper

− 0.999 27 5 273.4 151.0 424.4 −32.86%

Data from Table 4 and Table 5 of Katariya et al.4

5. CONCLUSIONS In this study, a reactive-distillation process for producing cyclohexanol by direct hydration of cyclohexene is developed. The design of excess reactant water is necessary in order to increase the conversion of cyclohexene. A decanter/stripper design configuration can be used to separate excess water with cyclohexanol product by means of natural liquid−liquid split in the decanter. The energy requirement of the proposed design flowsheet is also compared with other design alternatives in open literature. It is found that significant energy savings can be made by using this proposed design flowsheet.

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REFERENCES

AUTHOR INFORMATION

Corresponding Author

*Tel: +886-2-3366-3063. Fax: +886-2-2362-3040. E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS Financial support from the National Science Council of the R.O.C. under Grant No. NSC 100-2221-E-002-115-MY3 is greatly appreciated. G



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Resin. 2. Effect of Sulfolane on the Reaction Kinetics. Ind. Eng. Chem. Res. 1992, 31, 1425−1433. (15) Panneman, H. J.; Beenackers, A. A. C. M. Solvent Effects on the Hydration of Cyclohexene Catalyzed by a Strong Acid Ion-exchange Resin. 3. Effect of Sulfolane on the Equilibrium Conversion. Ind. Eng. Chem. Res. 1992, 31, 1433−1440. (16) Panneman, H. J.; Beenackers, A. A. C. M. Solvent Effects in the Liquid Phase Hydration of Cyclohexene Catalyzed by a Macroporous Strong Acid Ion-exchange Resin. Chem. Eng. Sci. 1992, 47, 2635−2640. (17) Qiu, T.; Huang, C. H.; Li, C. G.; Zhang, X. W.; Wang, X. D. Study of Feasibility of Reactive Distillation Process for the Direct Hydration of Cyclohexene to Cyclohexanol Using a Cosolvent. Ind. Eng. Chem. Res. 2013, 52, 8139−8148. (18) Wu, Y. C.; Chien, I. L.; Luyben, W. L. Two Stripper/Decanter Flowsheet for Methanol Recovery in the TAME Reactive-Distillation Process. Ind. Eng. Chem. Res. 2009, 48, 10532−10540. (19) Michelsen, M. L. The Isothermal Flash Problem, Part I: Stability. Fluid Phase Equilib. 1982, 9, 1−19. (20) Seider, W. D.; Seader, J. D.; Lewin, D. R.; Widagdo, S. Product and Process Design Principles Synthesis, Analysis, and Evaluation; Wiley: Hoboken, NJ, 2010. (21) Lee, H. Y.; Lee, Y. C.; Chien, I. L.; Huang, H. P. Design and Control of a Heat-Integrated Reactive Distillation System for the Hydrolysis of Methyl Acetate. Ind. Eng. Chem. Res. 2010, 49, 7398− 7411. (22) Douglas, J. M. Conceptual Design of Chemical Processes; McGraw Hill: New York, 1988. (23) Imam, R. A.; Freund, H.; Guit, R. P. M.; Fellay, C.; Meier, R. J.; Sundmacher, K. Evaluation of Different Process Concepts for the Indirect Hydration of Cyclohexene to Cyclohexanol. Org. Process Res. Dev. 2013, 17, 343−358.

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Reactive-Distillation Process for Direct Hydration of Cyclohexene to

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