Design of a Reactive Distillation Column for Direct Preparation of

Department of Chemical and Biochemical Engineering, College of Chemistry and Chemical Engineering, Xiamen ... Publication Date (Web): November 9, 2009...
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Ind. Eng. Chem. Res. 2009, 48, 10779–10787

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Design of a Reactive Distillation Column for Direct Preparation of Dichloropropanol from Glycerol Zheng-Hong Luo,* Xiao-Zi You, and Jie Zhong Department of Chemical and Biochemical Engineering, College of Chemistry and Chemical Engineering, Xiamen UniVersity, Xiamen 361005, China

In the present study, a reactor design for the direct preparation of dichloropropanol (DCP) from glycerol (GLY) is presented. As an optimum design, a reactive distillation (RD) column in which the second chlorination of GLY occurs during the direct preparation of DCP from GLY has been introduced. Using a simple equilibrium mathematical model of RD, a pilot plant scale column is designed. Furthermore, the whole preparation process of DCP from GLY used the optimum RD column as the central unit is simulated by using an advanced software tool, namely ASPEN PLUS. The pilot plant realization of the whole preparation process is presented and its results are compared with the theoretical prediction. On the basis of the pilot plant data and simulated results, a new technology of DCP production has been designed. 1. Introduction Dichloropropanol (DCP), a mixture of isomers including 2,3-dichloro-1-propanol (2,3-DCP) and 1,3-dichloro-2-propanol (1,3-DCP), is an intermediate in the production of epichlorohydrin (ECH),1 an important intermediate in chemical industries.1-3 A direct process for the preparation of DCP has been developed recently based on chlorination of glycerol (GLY) with hydrochloric acid (HCl),2,3 and it finds wide application in the chemical industries. There are many reports on investigation of the direct preparation processes;1-20 however, most of them focused on purification of GLY or on separation/recycling of the catalysts.4-18 Results indicate that chlorination of GLY with HCl is reversible, and a universal reaction mechanism can be obtained as shown in Scheme 1.4-10 A similar mechanism (Scheme 1) was proposed by Tesser et al. for this process,19 which is the same as shown in our research.20 Two reversible steps of chlorination reactions are presented, and the catalysts, e.g. acetic acid (HAc) and heteropolyacid (HPA), are required in the reactions.4,5,19,20 At least one of the products must be continuously removed from the reaction mixture to complete the corresponding reactions. The conventional chlorination is usually carried out in an equilibrium reactor, and, the products, byproduct, and the residual reactants are separated with rectification columns.21 In principle, reactive distillation (RD) can be performed on equilibrium reactors, and it seems to be a simple, energy-saving process with lower investment and operating costs.21,22 Therefore, simplification of the DCP preparation processes is made possible by the introduction of the RD techniques, and furthermore, the process will be environmentally friendly. No detailed studies have been reported on the direct preparation process of DCP from GLY with RD columns. Krafft et al. used RD columns in their research, but no feasible operation conditions were provided.15 Our previous work could not obtain pilot-plant data or perform practical analyses, either.23 However, in this work, the RD columns are successfully integrated into the direct preparation process of DCP from GLY, and the industrial feasibility is described. By using a simple equilibrium mathematical model of RD

columns, the designs of a pilot-plant scale RD column and a new process of the DCP production are achieved. 2. Reaction Description and Process Design According to our previous work,20,23 the direct preparation reaction of DCP from GLY follows the SN2 mechanism and mainly consists of two steps of reversible chlorination, which can also be seen in Scheme 1. In addition, our previous results show that the first step of chlorination is much faster than the second one.20,23 Namely, monochlorohydrin (MCP), a mixture of isomers including 3-chloro-1,2-propanediol (3MCP) and 2-chloro-1,3-propanediol (2-MCP), was produced very fast in the first step of reaction of GLY with HCl, while DCP was obtained much slower in the following reaction of MCP with HCl. A flow diagram of the two-step reactions can be obtained by introducing the RD column as shown in Figure 1.23 Namely, the first step of the reactions is accomplished in a reactor (as E-2 in Figure 1) and the second step occurs in the RD column (as E-4 in Figure 1), which is different from the processes of the current technology.12,14-19 The GLY-HAC mixture from the mixer (E-1) and the gaseous HCl are fed into a primary reactor (E-2). The temperature of the reactor is maintained at 100 °C. The equilibrium mixture from E-2 and the gaseous HCl are introduced into the RD column (E-4) from its top and bottom, respectively. The RD column is packed with ceramic, with the packing factor 0.7 and the packing height 2.25 m. The distillate from the RD column has a composition close to the DCP-water(H2O)-HCl ternary azeotrope. The bottom product that mainly consists of DCP and MCP from the RD column is fed into the distillation column (E-6) where heavy Scheme 1. Chlorination Network of GLY with HCl

* To whom correspondence should be addressed. E-mail: luozh@ xmu.edu.cn. Tel.: +86-592-2187190. Fax: +86-592-2187231. 10.1021/ie900933b CCC: $40.75  2009 American Chemical Society Published on Web 11/09/2009

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Figure 1. Schematic representation of the new preparation process of DCP: E-1 mixer; E-2 continuous stirred tank reactor; E-3 and E-5 heater; E-4 reactive distillation column; E-6 distillation column. Table 1. Kinetic Constants at T ) 100 °C with HAC as the Catalyst19

HAC

K1R

K2R

K3R

K4R

34619 ( 4012

2342 ( 198

1576 ( 302

17 ( 4

impurities are separated at the bottom of the column and then recycled to the RD column. The distillate of E-6 is DCP (55 mol %) and H2O and can be directly used as the reactant in the process of producing ECH. Since the yield of DCP is of the main interest in this process, the separation of E-6 with distillation columns is not discussed in detail. This process design includes one chlorination reactor (E-2, CSTR), one RD column (E-4), and one purified distillation column (E-6). And the following study concentrates on the design of the RD column with the purification process of distillation discussed briefly. HAC is selected as the chlorination catalyst with the feed rate of 0.05625 kmol/h. The feed rates of GLY into E-1 and HCl into E-2 and E-4 are 0.917, 1.375, and 0.917 kmol/h, respectively. 3. Design of the RD Column (E-4) 3.1. Mathematical Model for the RD Column. A mathematical model, considering the heat of reaction, is developed based on the mass and heat balances as well as the vapor-liquid equilibrium at all RD stages. Several assumptions are made in this work:24-28 (1) The operation is at the steady-state conditions. (2) The reactions take place entirely in the bulk liquid. (3) The vapor phase fugacity coefficients are unity due to the low pressure. (4) To simplify the calculations, the pressure drop along the column is neglected, and therefore, the whole system is considered to be operated under the atmospheric pressure. The stages are numbered from the top of the RD column (condenser) to the bottom (reboiler). And thus, the MESH equations (mass balance, phase equilibrium, summation equation, and enthalpy balance) can be obtained for all stages (see the Appendix and refs 24 and 25). In this paper, ASPEN PLUS simulator (RADFRAC) is used to solve the MESH equations.25-30 Here, we only describe the selections of the reaction kinetic equations and thermodynamic model. More detailed relevant model equations and assumptions may be found elsewhere.19,25-30

Corresponding reaction kinetic equations selected in this work are from the comprehensive chlorination kinetics model proposed by Tesser et al.:19

[

r1 ) CC K1CHCGLY -

]

K1 C C KE1 R W

r2 ) K2CCCHCGLY

[

r3 ) CC K3CHCR -

(2)

]

K3 C C KE3 Rγ W

r4 ) K4CCCHCR

(1)

(3) (4)

The accurate descriptions of the thermodynamic properties and phase behavior in reactors, separation units, etc., are the most important aspects in the process simulation. Choosing the model with the appropriate property for thermodynamic calculations is a important work. There are two categories of thermodynamic models, namely, the activity-coefficient category and the equation-of-state (EOS) category.31 In this work, the nature of the species (GLY, HAC, HCl, 2-MCP, 3-MCP, et al.) involved and the atmospheric pressure conditions in the whole process suggest the use of an activitycoefficient category. ASPEN PLUS contains several properties and thermodynamic models of activity coefficient specifically and can be used to describe the chemical species behavior. Among them, the NRTL (nonrandom two liquids) equation can be widely used in chemical engineering thermodynamics and shows excellent performance in predicting the thermodynamics properties of multicomponent systems only with the data of binary systems. Therefore, the NRTL equation was used to describe the thermodynamic properties of the chemical components in this work, which is the same as other works.32-35 For detailed information on the NRTL equation used in this work, the reader is refered to refs 32-35. In this work, all pure-component parameters involved in the process were taken directly from the database contained in ASPEN PLUS.36 Moreover, the selections in this work also come from the work of Tesser et al. The values of kinetic equation parameters are listed in Table 1.19 3.2. Optimization of the RD Column. Using the model described above, it is the operational conditions of a pilot plant scale column are optimized theoretically. In this simulated system, the molar yield of DCP (nDCP/nGLY0), the molar

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Figure 2. Experimental and simulated mole-fraction profiles of the liquid phase in the RD column. (Simulated conditions: stage number )16; operational pressure ) 1 atm; reflux ratio ) 3; feed-stage positions of MCP and HCl ) the 2nd and 15th stages; nHCl0/nGLY0 )1.5; distillation rate ) 1.7 kmol/h; residence time ) 6000 s).

conversion of GLY(1 - nGLY/nGLY0) and the separation efficiency of RD are considered. Generally, the stage number of the RD column is mainly influenced by stuffing. As described eariler, the working RD column in our work is packed with ceramic, with the packing factor 0.7. Typically, the corresponding stage number of the RD column does not exceed 18.26-30,37-39 The effect of the stage number is simulated, and the results show that the RD column model does not work when the stage number is 17 or 18. Consequently, the stage number is determined to be 16. 3.2.1. Size of Reactive Stages and Nonreactive Sections. Within the RD column, the reaction has an important effect on the volatilities. If no reaction occurs, the volatilities exhibit a smooth and predictable profile. In the RD column, reactions may occur at all stages.21,28-30,39 Concentration profiles of the chemical species are simulated using the model above, and results are shown in Figures 3 and 4. Figure 2 illustrates the simulated concentrations of the main chemical species in the liquid phase of the reaction system along the column, which shows that a stepwise change on the concentration curves. Furthermore, according to Figure 2, it can be found that the chemical reaction causes remarkable changes in all components profiles in the feed stage position. The mole fraction of water in overhead distillate is more than 98%, and the mole fraction of 3-MCP and 1,3-DCP in the bottom is more than 95% (mol/mol). As described in Figure 2, the reaction is concentrated in the fifth to eleventh stages due to the fast changes of the concentrations of the main chemical species along these stages. Figure 3 shows the simulated total concentrations of the main chemical species in the reaction system along the columns, which proves further that the reaction is concentrated in the fifth to eleventh stages. Moreover, it also shows that there are two reaction peak zones resulting from the high concentration of reactants in the two feed stages and the proper material ratio. 3.2.2. Effect of the Reflux Ratio. The above model is also used to simulate the effect of the reflux ratio on the yield of DCP. As shown in Figure 4, it can be found that with the increase of the reflux ratio, the simulated yield of DCP increases

Figure 3. Simulated total mole-fraction profiles in the RD column. (Simulated conditions: stage number ) 16; operational pressure ) 1 atm; reflux ratio ) 3; feed-stage positions of MCP and HCl ) the 2nd and 15th stages; nHCl0/nGLY0 )1.5; distillation rate ) 1.7 kmol/h; residence time ) 6000 s).

to about 1.0 and then decreases after that. In order to obtain the maximal yield of DCP, the optimum reflux ratio should be 1.0. However, as to the whole preparation process of DCP, the low reflux ratio leads to insufficient product separation and traces of the heavy component, which limits the conversion of GLY and contaminates the top product of the RD column. In addition, at a high reflux ratio, the reactants are separated so effectively it limits the conversion of GLY. The optimal reflux ratio is an eclectic solution between these situations, although it has been shown that the optimal reflux ratio is 1.0. Thus, it is considered that 1.5 is better.26-29 3.2.3. Effect of the Material Mole Ratio. In order to investigate the optimum material ratio for the designed RD column, the sensitivity analysis of the material mole ratio is simulated as described in Figure 6. If the high yield of DCP in the RD column is consider solely, the obtained optimal material ratio is about 1.24. A high yield of DCP can be achieved if the losses of 3-MCP from the bottom of the RD column are recycled after being separated in the whole preparation process of DCP

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Figure 4. Simulated yield of DCP vs reflux ratio of the RD column. (Simulated conditions: stage number ) 16; operational pressure ) 1 atm; nHCl0/nGLY0 ) 1.5; feed-stage positions of MCP and HCl ) the 2nd and 15th stages; distillation rate ) 1.7 kmol/h; residence time ) 6000 s).

Figure 5. Simulated yield of DCP and separation efficiency vs material mole ratio of the RD column. (Simulated conditions: stage number ) 16; operational pressure ) 1 atm; reflux ratio ) 1.5; feed-stage positions of MCP and HCl ) the 2nd and 15th stages; distillation rate ) 1.7 kmol/h; residence time ) 6000 s).

(the whole preparation process is listed in sections 2 and 6). However, the separation efficiency is lower as shown in Figure 5. On the other hand, some researchers suggested the optimal molar ratio of HCl with GLY is about 1.15,40 Although the optimal material ratio is about 1.24 from the viewpoint of the yield of DCP according to Figure 6, the suggested material molar ratio is an eclectic solution between these situations practically. So, the value of 1.02 is selected in this work, which represents the point of intersection between the curves of the yield and the separation efficiency. 3.2.4. Effect of the Feed Stage Position. It is well-known that the product quality depends on the performance of the reaction zone in the RD column as well as the feed point position. In addition, the feed point position affects the performance of the reaction section. In our study, the same situation is observed. Corresponding simulations are obtained via our model and the simulated results are shown in Figure 6. When the feed positions of the reactants are located inside the reactive region, a relatively low yield of DCP is obtained. The yield of DCP increases to a maximum when the feed points are located directly below and above the reactive section of the RD column.

Figure 6. Simulated yield of DCP vs feed-stage position of the RD column. (Simulated conditions: stage number ) 16; operational pressure ) 1 atm; reflux ratio ) 1.5; nHCl0/nGLY0 ) 1.04; distillation rate ) 1.7 kmol/h; residence time ) 6000 s).

Figure 7. Simulated yield of DCP and separation efficiency vs distillation rate of the RD column. (Simulated conditions: stage number ) 16; operational pressure ) 1 atm; reflux ratio ) 1.5; nHCl0/nGLY0 ) 1.04; feedstage positions of MCP and HCl ) the 3rd and 15th stages; residence time ) 6000 s).

Figure 6 shows that the yield of DCP depends less on the feed stage of MCP than on the feed stage of HCl. And it shows that the optimal feed stages for MCP and HCl are the third and fifteenth stages, respectively. 3.2.5. Effect of the Distillation Rate. The relationship of the distillation rate with the yield of DCP and the separation efficiency is illustrated in Figure 7. As shown in Figure 7, with the increase of the distillation rate, the yield of DCP decreases and the separation efficiency increases. Accordingly, the optimal distillation rate is an eclectic solution between these situations. Furthermore, in this work, it is considered that the value of the distillation rate is 1.85, which represents the point of intersection between the curves of the yield and the separation efficiency. 3.2.6. Effect of the Residence Time. Figure 8 describes the influence of the residence time in the RD column on the operational performance of the RD column. The yield of DCP and the reboiler duty both increase with the increase of the residence time. In practice, the maximum of the yield of DCP and the minimum of the reboiler duty fit our expectations. Accordingly, the optimal residence time is an eclectic solution between these situations. Therefore, although it has been shown

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Figure 8. Simulated yield of DCP and reboiler duty vs residence time in the RD. (Simulated conditions: stage number ) 16; operational pressure ) 1 atm; reflux ratio ) 1.5; nHCl0/nGLY0 ) 1.04; feed-stage positions of MCP and HCl ) the 3rd and 15th stages; distillation rate ) 1.85 kmol/h).

Figure 9. Simulated mole fraction of DCP vs reflux ratio of the purified distillation column. (Simulated conditions: stage number ) 10; operational pressure ) 1 atm; feed-stage position ) the fifth stage).

that the optimal residence time is higher than 15 000 s, it is considered that 7500 s is better according to Figure 8.

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Figure 10. Simulated mole fraction of DCP vs feed-stage position of the purified distillation column. (Simulated conditions: stage number ) 10; operational pressure ) 1 atm; reflux ratio ) 5).

Figure 11. Simulated mole fraction of DCP vs stage number of the purified distillation column. (Simulated conditions: operational pressure ) 1 atm; reflux ratio ) 5; feed-stage position ) the 3th stage).

Figure 11, the simulated result of which shows that the optimal stage number is 8.

4. Design of the Purified Distillation Column (E-6) ASPEN PLUS software is also used to simulate the purified distillation. The selected distillation model is based on the assumption of vapor-liquid equilibrium on every stage, which is the same one described in section 3.1. Physicochemical properties such as liquid- and vapor-phase enthalpies as well as phase equilibrium are calculated using the software database. 4.1. Effect of the Reflux Ratio. The dependent on simulated effect of the reflux ratio on the separation efficiency of the distillation column is illustrated in Figure 9. The simulated result shows that the separation efficiency of the column is maximal when the reflux ratio is 4. 4.2. Effect of the Feed Stage Position. Figure 10 describes the effect of the feed stage position in the distillation column on the operational performance of the column. According to Figure 10, one can know that the separation efficiency of the column is maximal when the feed stage position is 3. 4.3. Effect of the Stage Number. The effect of the stage number on the separation efficiency of the column is shown in

5. Direct Preparation Process Simulation of DCP from GLY On the basis of the above study, the main units are optimized. Corresponding optimal units are used to construct the new technology shown in Figure 1. ASPEN PLUS software is also used to simulate the constructed direct preparation process of DCP from GLY. Simultaneously, all physicochemical properties as well as phase equilibrium contained in the process are calculated using the software database. 6. Pilot Plant Research 6.1. Pilot Plant Experiments. The simulated results are verified by pilot plant experiments. The DCP pilot plant scale process equipment is given in Figure 1. The central unit, i.e. the RD column, is given in Figure 12. The RD column has an inner diameter of 0.06 m, a height of 2.25 m, a length of 0.75 m, and a volume of 20 L. The column consists of three distinct parts, namely, (1) nonreactive stripping section with a height of 0.45 m, (2) reactive section with a height

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Figure 12. Reactive distillation setup used in the pilot plant experiments. (Operational conditions: stage number ) 16; operational pressure ) 1 atm; reflux ratio ) 1.5; feed-stage positions of MCP and HCl ) the 3th and 15th stages; distillation rate ) 1.85 kmol/h).

of 0.9 m, and (3) nonreactive enriching section with a height of 0.6 m. The column packing with ceramic is equipped with a total condenser and a forced circulation reboiler. The column is supplied with protective heating to ensure adiabatic operation. Temperature is measured at the bottom, the reactive zone, and the top, respectively. Samples taken from the bottom and the distillate are analyzed using gas chromatography to obtain the concentrations. The operational conditions of the experiments are selected based on the above optimal simulated results and shown in Figure 12 and section 2. 6.2. Comparison of Simulation and Experiment. 6.2.1. Comparison of Simulation and Experiment for the RD Column. In order to verify the above RD model, the liquidphase composition data along the RD column are obtained via experiment. Since Figure 2 shows the simulated composition profiles at certain conditions, corresponding experimental data at the same conditions as those are obtained. Accordingly, here, the selected experimental conditions are listed in the caption of Figure 2. The comparison of experimental and simulated liquidphase composition profiles along the RD column is given in Figure 2, which shows a good agreement between experimental data and the simulated result. Furthermore, temperature in the RD column is a key parameter of the RD process. The adjustment and control of the parameter influence the RD process directly.25-28 In this study, the temperature profile along the RD column are also predicted by the model. Corresponding

Figure 13. Comparison of the simulated and experimental temperature profile.

comparison of temperature profile along the RD column is given in Figure 13, which also describes a good agreement between experimental data and the simulated result. In addition, according to Figure 13, one can find that in the reaction zone of the RD column, i.e. between 2nd and 15th stage, the temperature decreases in the column decreases continuously. Generally, it

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Table 2. Comparison of Simulation and Experiment for the Whole Process shown in Figure 1

temperature/K pressure/Pa flow rate/(kmol/h) GLY/(kmol/h) 3-MCP/(kmol/h) 1,3-DCP/(kmol/h) 2,3-DCP/(kmol/h) HCl/(kmol/h) H2O/(kmol/h)

simulation experiment simulation experiment simulation experiment simulation experiment simulation experiment simulation experiment simulation experiment simulation experiment simulation experiment

mixture

cool

373.0 371.0 1.01 × 105 1.01 × 105 2.39

343.0 342.0 1.01 × 105 1.01 × 105 6.82

1.61 × 10-1

1.26 × 10-2

6.15 × 10-1

4.88

1.29 × 10-1

1.33

1.48 × 10-3

3.96 × 10-2

4.98 × 10-1

6.19 × 10-7

348.0 346.7 1.01 × 1.01 × 1.20 1.23 2.45 × 0 1.85 × 1.79 × 2.11 × 2.01 × 2.90 × 2.75 × 5.26 ×

9.24 × 10-1

5.12 × 10-1

6.67 × 10-1

is almost impossible to obtain a decreasing temperature in the RD column with increasing stage, especially when it is inside the reaction zone.25-27 However, in our work, an equilibrium stage model considers the exothermal reaction, which leads to the decreasing temperature. In practice, Thotla et al.28 also obtained a similar result in a RD column used to produce diacetone alcohol and mesityl oxide from acetone based on a similar equilibrium stage model. As a whole, the above comparisons prove the validity of the RD model suggested in this work. In practice, at our experimental condition, namely, the residence time of 6000 s and the distillation rate of 1.7 kmol/h, the reactive distillation entirely reaches the state of the equilibrium. Therefore, the equilibrium model can be used to predict corresponding parameters, design and optimize the pilot plant scale RD column. 6.2.2. Comparison of Simulation and Experiment for the Whole Process. On the basis of above process simulation, some data, i.e. streamflow rates, temperatures etc., can be obtained. In addition, certain data can also be obtained via the pilot plant experiments. The simulated and experimental data are shown in Table 2, which illustrates a good agreement between the simulated and experimental data in the whole process. Furthermore, according to Table 2, one can calculate the conversion of GLY, the yield of DCP, and the selectivity of the process. Corresponding equations and results based the simulated data are given as Conversion of GLY 0.917 - 0.0126 × 100% ) 98.62% X) 0.917

Selectivity of reaction 95.09 Y × 100% ) 96.42% S) ) X 98.62

product 389.5 389.0 1.01 × 1.01 × 1.45 1.42 5.00 × 0 4.88 × 0.17 × 8.44 × 8.36 × 1.09 × 1.15 × 6.19 ×

105 105 10-13 10-6 10-6 10-4 10-4 10-6 10-6 10-1

recycle 471.1 471.5 1.01 × 1.01 × 5.36 5.30 1.26 × 1.17 × 4.84 0.14 4.86 × 4.61 × 2.88 × 2.75 × 5.81 ×

105 105 10-7 10-2 10-2 10-1 10-1 10-2 10-2 10-7

5.12 × 10-1

105 105 10-2 10-2 10-1 10-1 10-2 10-2 10-23

5.86 × 10-7

optimize the RD column, a simple equilibrium mathematical model is used to describe the RD column based on the advanced software tool, namely ASPEN PLUS. In addition, ASPEN PLUS software is also used to simulate the purified distillation unit and the whole preparation process. The pilot plant scale process equipment is put forward and corresponding pilot plant experiments are performed to verify the theoretical prediction. The results show that it is possible to obtain high-yield and high-selectivity of DCP from the process based on the RD column designed. The obtained conversion of GLY is 98.62 mol %, and the yield of DCP is 93.23 mol %. Acknowledgment The authors thank Yantai Wanhua Polyurethanes Co., Ltd. for financial support. Authors also thank the anonymous referees for comments on this manuscript. The simulation work are implemented by advanced software tools (ASPEN PLUS) provided by China Lanzhou Petrochemical Company Research Institute and Dalian Nationalities University. Appendix: MESH Equations The main MESH equations are listed as follows:25,41 Mass balance of component i Lj-1,i - RLj Lj,i - Rvj Vj,i + rj,i +

(5) (

Yield of DCP

0.844 + 0.011 × 100% ) 93.23%, Y) 0.917

gas

∑

∑ (x V ) + V j j

j+1,i

+ fj,i ) 0 (8)

Heat balance

∑ L )H ( ∑ V )h

Lj-1,i)Hj-1 - RLj (

(6)

j,i

j+1,i

j

∑ V )h + + ( ∑ f )H + Q

- Rvj (

j+1

j,i

j

F j

j,i

j

)0

(9)

Vapor-liquid equilibrium Kpj,iLj,i /

(7)

∑L

j,i

- Vj,i /

∑V

j,i

)0

Chemical reaction equilibrium

Accordingly, it is possible to obtain high yield and high selectivity of DCP from the process introducing the RD column designed in this work.

Krj -

∏a

ni j,i

)0

Nomenclature 7. Conclusions It has been confirmed both by simulation and pilot plant scale experimental results that it is feasible to produce DCP via a new process using a RD column as the central unit. In order to

(10)

aj,i ) activity of component i on the jth stage, mol/L CC ) concentration of catalyst, 103 mol/L CH ) concentration of HCl, 103 mol/L CGLY ) concentration of GLY, 103 mol/L

(11)

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CR ) concentration of 3-MCP, 103 mol/L CRγ ) concentration of 1,3-DCP, 103 mol/L CW ) concentration of H2O, 103 mol/L fj,i ) feed rate of component i on the jth stage, mol/L F ) component feed rate, mol/L hj ) vapor mixture molar enthalpy on the jth stage, kJ/mol Hj ) liquid mixture molar enthalpy on the jth stage, kJ/mol i ) component index j ) stage index KA ) kinetic constant of the Ath reaction (A ) 1, 2, 3, 4), 1/(mol2 h) KE1 ) equilibrium constant of the 1st reaction KE3 ) equilibrium constant of the 3rd reaction Kjr ) reaction equilibrium constant on the jth stage p Kj,i ) phase equilibrium constant on the jth stage Lj,i ) liquid flow rate of component i on the jth stage, mol/h ni ) stoichiometric coefficient of component i nDCP ) mole flow of DCP in reaction system, mol/h nGLY0 ) initial mole flow of GLY in reaction system, mol/h nGLY ) mole flow of GLY in reaction system, mol/h nHCl ) mole flow of HCl in reaction system, mol/h nHCl0 ) initial mole flow of HCl in reaction system, mol/h Qj ) heat duty on the jth stage, J/h rj,i ) -dni/dt corresponding to these reactions (reactions 1-4 in Scheme 1), which can lead to the change of component i, mol/h ri ) reaction rate of the ith reaction (i ) 1, 2, 3, 4), mol/h RjL ) residual function for mass transfer rate from the interface to the liquid phase Rjv ) residual function for mass transfer rate from the vapor phase to the interface S ) selectivity of reaction T ) temperature, K Vj,i ) vapor flow rate of component i on the jth stage, mol/h xj ) rate extent for equilibrium reaction on the jth stage X ) conversion of GLY Y ) yield of DCP

Literature Cited (1) Horsley, H. L. Encyclopedia of Chemical Technology; John Wiley and Sons: New York, 1965. (2) Ma, L.; Zhu, J. W.; Yuan, X. Q.; Yue, Q. Synthesis of Epichlorohydrin from Dichloropropanols Kinetic Aspects of the Process. Chem. Eng. Res. Des. 2007, 85, 1580. (3) Carra`, S.; Santacesarla, E.; Morbldlll, M. Synthesis of Epichlorohydrin by Elimination of Hydrogen Chloride from Chlorohydrins: 1. Kinetic Aspects of the Process. Ind. Eng. Chem. Process Des. DeV. 1979, 18, 424. (4) Wang, L. L.; Liu, Y. M.; Xie, W. Highly Efficient and Selective Production of Epichlorohydrin through Epoxidation of Allyl Chloride with Hydrogen Peroxide over Ti-MWW Catalyst. J. Catal. 2007, 246, 205. (5) Lee, S. H.; Park, D. R.; Kim, H.; Lee, J.; Jung, J. C.; Woo, S. Y.; Song, W. S.; Kwon, M. S.; Song, I. K. Direct Preparation of Dichloropropanol (DCP) from Glycerol using Heteropolyacid (HPA) Catalysts: A Catalyst Screen Study. Catal. Commun. 2008, 9, 1920. (6) Rashid, U.; Anwar, F. Production of Biodiesel through Optimized Alkaline-Catalyzed Transesterification of Rapeseed Oil. Fuel 2008, 87, 265. (7) Xie, W.; Peng, H.; Chen, L. Transesterification of Soybean Oil Catalyzed by Potassium Loaded on Alumina as a Solid Base Catalyst. Appl. Catal. A-Gen. 2006, 300, 67. (8) Macleod, C. S.; Harvey, A. P.; Lee, A. F.; Wilson, K. Evaluation of the Activity and Stability of Alkali-Doped Metal Oxide Catalysts for Application to an Intensified Method of Biodiesel Production. Chem. Eng. J. 2008, 135, 63. (9) Gerpen, J. V. Biodiesel Processing and Production. Fuel Process. Technol. 2005, 86, 1097. (10) Seraphim, P.; Stylianos, F.; Michel, F.; Isabelle, C.; Maria, C. P.; Michael, K.; Ivan, M.; George, A. Biotechnological Valorization of Raw Glycerol Discharged after Bio-Diesel Manufacturing Process: Production of 1,3-propanediol, Citric Acid and Single Cell Oil. Biomass Bioenerg. 2008, 32, 60.

(11) Jiang, J. X.; Zhang, P. P.; Yao, C. The Development of the Production of Epichlorohydrin from Glycerol. Modern Chem. Ind. 2006, 26, 71; in Chinese. (12) Krafft, P.; Gilbeau, P.; Balthasart, D. Crude glycerol-based product, process for its purification and its use in the manufacture of dichloroproanol. W.O. Patent 1,44,335, 2007. (13) Zhao, X. J.; Bai, Z. L. Study on the Process of Producing Dichloropropanol from Glycerol Catalyzed by Organic Acid. Chem. Interm. 2008, 15, 19; in Chinese. (14) Gilbeau, P.; Krafft, P. Producing chlorinated organic compounds e.g. dichloropropanol inVolVes using glycerol obtained from renewable raw materials, as a starting product. F.R. Patent 2,868,419, 2005. (15) Krafft, P.; Gilbeau, P.; Gosselin, B.; Claessens, S. Process for producing dichloropropanol from glycerol, the glycerol coming eVentually from the conVersion of animal fats in the manufacture of biodiesel. W.O. Patent 0,54,167, 2005. (16) Kubicek, P.; Sladek, P.; Buricova, I. Method of preparing dichloroproanols from glycerine. W.O. Patent 0,21,476, 2005. (17) Schreck, D. J.; Kruper, W. J.; Varjian, R. D.; Jones, M. E.; Campbell, R. M.; Kearns, K.; Hook, B. D.; Briggs, J. R.; Hippler, J. G. ConVersion of a multihydroxylated-aliphatic hydrocarbon or ester thereof to a chlorohydrin. W.O. Patent 0,20,234, 2006. (18) Krafft, P.; Franck, C.; Andolenko, I. D.; Veyrac, R. Process for the manufacture of dichloropropanol by chlorination of glycerol. W.O. Patent 0,54,505, 2007. (19) Tesser, R.; Santacesaria, E.; Di Serio, M.; Di Nuzzi, G.; Fiandra, V. Kinetics of Glycerol Chlorination with Hydrochloric Acid: a New Route to R,γ-Dichlorohydrin. Ind. Eng. Chem. Res. 2007, 46, 6456. (20) Luo, Z. H.; You, X. Z.; Li, H. R. Direct Preparation Kinetics of 1,3-Dichloro-2-Propanol from Glycerol using Acetic Acid Catalyst. Ind. Eng. Chem. Res. 2009, 48, 446. (21) Smejkal, Q.; Hanika, J.; Kolena, J. 2-Methylpropylacetate Synthesis in a System of Equilibrium Reactor and Reactive Distillation Column. Chem. Eng. Sci. 2001, 56, 365. (22) Sundmacher, K.; Hoffmann, U. Multicomponent Mass and Energy Transport on Different Length Scale in a Packed Reactive Distillation Column for Heterogenously Catalyzed Fuel Ether Production. Chem. Eng. Sci. 1994, 49, 4443. (23) Luo, Z. H.; You, X. Z.; Hua, W. Q.; Luo, W. X.; Hu, B. B.; Zhang, H. K.; Liao, Z. T.; Ding, J. S.; Yang, W. H. Process for producing dichloropropanol from glycerol. C.N. 101,429,099, 2008. (24) Gao, J.; Zhao, X. M.; Zhou, L. Y.; Huang, Z. H. Investigation of Ethyllactate Reactive Distillation Process. Trans. Inst. Chem. Eng. Part A Chem. Eng. Res. Des. 2007, 85, 525. (25) Hanika, J.; Kolena, J.; Smejkal, Q. Butylacetate via Reactive Distillation-Modelling and Experiment. Chem. Eng. Sci. 1999, 54, 5205. (26) Harmsen, G. J. Reactive Distillation: the Front-Runner of Industrial Process Intensification a Full Review of Commercial Application, Research, Scale-up, Design and Operation. Chem. Eng. Process. 2007, 46, 774. (27) Taylor, R.; Krishna, R. Modelling Reactive Distillation. Chem. Eng. Sci. 2000, 55, 5183. (28) Thotla, S.; Agarwal, V.; Mahajani, S. M. Simultaneous Production of Diacetone Alcohol and Mesityloxide from Acetone Using Reactive Distillation. Chem. Eng. Sci. 2007, 62, 5567. (29) Smejkal, Q. Simulation of Reactive Distillation of Butylacetate. Ph.D. Dissertation, Institute of Chemical Technology Prague, Prague, 1999. (30) Venkataraman, S.; Chan, W. K.; Boston, J. F. Reactive Distillation using ASPEN PLUS. Chem. Eng. Prog. 1990, 87, 45. (31) Orbey, H.; Bokis, C. P.; Chen, C. C. Polymer-Solvent Vapor-Liquid Equilibrium: Equation of State versus Activity Coefficient Models. Ind. Eng. Chem. Res. 1998, 37, 1567. (32) Sadeghi, R. Modification of the NRTL and Wilson Models for the Representation of Phase Equilibrium Behavior of Aqueous Amino AcidElectrolyte Solutions. Can. J. Chem. 2008, 86, 1126. (33) Athes, V.; Paricaud, P.; Ellaite, M.; Souchon, I.; Furst, W. VapourLiquid Equilibria of Aroma Compounds in Hydroalcoholic Solutions: Measurements with a Recirculation Method and Modelling with the NRTL and COSMO-SAC Approaches. Fluid Phase Equilib. 2008, 265, 139. (34) Simoni, L. D.; Lin, Y.; Brennecke, J. F.; Stadtherr, M. A. Modeling Liquid-Liquid Equilibrium of Ionic Liquid Systems with NRTL, ElectrolyteNRTL, and UNIQUAC. Ind. Eng. Chem. Res. 2008, 47, 256. (35) Novak, L. T. Entity-based Eyring-NRTL Viscosity Model for Mixtures Containing Oils and Bitumens. Ind. Eng. Chem. Res. 2006, 45, 7392. (36) Aspen Physical Property System: Physical Property Methods and Models 11.0; Aspen Technology Corporation: Burlington, MA, 2001. (37) Hung, S. B.; Lai, I. K.; Huang, H. P.; Lee, M. J.; Yu, C. C. Reactive Distillation for Two-Stage Reaction Systems: Adipic Acid and Glutaric Acid Esterifications. Ind. Eng. Chem. Res. 2008, 47, 3076.

Ind. Eng. Chem. Res., Vol. 48, No. 24, 2009 (38) Florin, O.; Alexandre, C. D.; Alfred, B. Fatty Acid Esterification by Reactive Distillation. Part1: Equilibrium-based Design. Chem. Eng. Sci. 2003, 58, 3159. (39) Laureano, J.; Jose´, C. L. The Production of Butyl Acetate and Methanol via Reactive and Extractive Distillation: Process Modeling, Dynamic Simulation, and Control Strategy. Ind. Eng. Chem. Res. 2002, 41, 6735. (40) Shan, Y. H.; Han, L. L.; Xu, Z. H.; Li, M. S.; Zhu, J. J. A method for producing dichlorohydrin from glycerol. C.N. Patent 101,007,751, 2007.

10787

(41) Lee, J. H.; Dudukovic, M. P. A Comparison of the Equilibrium and Nonequilibrium Models for a Multicomponent Reactive Distillation Column. Comput. Chem. Eng. 1998, 23, 159.

ReceiVed for reView June 8, 2009 ReVised manuscript receiVed October 3, 2009 Accepted October 28, 2009 IE900933B