Heat-integrated reactive distillation process for TAME synthesis

Separation and Purification Technology 132 (2014) 468–478

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Heat-integrated reactive distillation process for TAME synthesis Xin Gao a,b, Fangzhou Wang a, Hong Li a,c,⇑, Xingang Li a,b,c a

School of Chemical Engineering and Technology, Tianjin University, Tianjin 300072, China National Engineering Research Center of Distillation Technology, Tianjin 300072, China c Collaborative Innovation Center of Chemical Science and Engineering, Tianjin 300072, China b

a r t i c l e

i n f o

Article history: Received 15 March 2013 Received in revised form 1 June 2014 Accepted 4 June 2014 Available online 12 June 2014 Keywords: Reactive distillation tert-Amyl methyl ether (TAME) Heat integrated Process design Dynamic simulation

a b s t r a c t The integration of reaction and separation in one single process unit is generally more preferable than their application individually, which allows less energy consumption and environmental impacts. This work presents a novel practicable configuration of heat-integrated process based on reactive distillation that aims to reduce furthermore the energy requirements for tert-amyl methyl ether (TAME) synthesis, leading to competitive operating costs. The optimal design flowsheet of heat-integrated with reactive distillation was obtained by minimizing the total annual cost (TAC) of this process. Simulation study was conducted to compare the heat-integrated scheme and conventional reactive distillation (CRD). The proposed technology was demonstrated to decrease the TAC which was 6.56% less than that of the CRD. The control strategy of this heat-integrated design was also developed. Only tray temperature control loop was needed to reject feed disturbances to maintain the quality of the products. Overall results demonstrate that the heat-integrated reactive distillation is a promising approach for TAME synthesis. Ó 2014 Elsevier B.V. All rights reserved.

1. Introduction Process intensification presents one of the most important trends in the current chemical engineering and process technology [1]. Reactive Distillation (RD), one of the best-known examples of process intensification, has been used commercially. By combining chemical reaction and distillation separation into a single process unit, RD has economic advantages in some chemical reaction systems compared with conventional reactors followed by distillation columns, particularly for reversible reactions in which chemical equilibrium constraints limit conversion in a conventional reactor [2]. Majority of current studies has been focused on modeling [3,4], conceptual design [5,6], process development [7,8], process control [9], process scale-up and hardware design [10,11]. However, the information about heat integration and energy saving of RD process is scarce. In the classical reactive distillation, only conventional distillation columns are considered, which are well-known by their low thermodynamic efficiency [12]. Therefore, it is possible to further intensify reactive distillation and improve energy saving by applying heat-integrated technology. Hernandez and co-workers [13] simulated the steady state and dynamic process ⇑ Corresponding author at: School of Chemical Engineering and Technology, Tianjin University, Tianjin 300072, China. Tel.: +86 022 27404701; fax: +86 022 27404705. E-mail address: [email protected] (H. Li). http://dx.doi.org/10.1016/j.seppur.2014.06.003 1383-5866/Ó 2014 Elsevier B.V. All rights reserved.

of a reactive Petlyuk column through an equivalent reactive dividing wall distillation column (RDWDC), and they found that the implementation of RDWDC reduced energy consumption of reactive distillation process. Kiss [14] investigated a heat-integrated process performed by applying heuristic rules in the reactive distillation system. The hot bottom product of the column, a mixture of fatty ester, was used to pre-heat both reactants: the fatty acid and alcohol feed streams. Results showed that the energy requirements of this process were about 45% lower than previously reported CRD processes. Huang and co-workers [15] advocated that prudent arrangements of the reactive section and deliberate determination of feed location were two effective methods that could complement internal heat-integration between reaction and separation operations within a reactive distillation column. Combination of these two methods could frequently provide more benefits and flexibility than using them individually. These advantages have motivated the development of optimization strategies to get optimal designs of reactive distillation column, but it is not an easy task, because design issues for reactive distillation systems are significantly more complex than those involved in ordinary distillation. Therefore, in heat-integrated reactive distillation system, we expect additional energy savings and also an optimal design problem even more complex design of heat-integrated reaction distillation. The main purpose of this study was to investigate the feasibility and effectiveness of further heat-integration for a RD system with a high pressure reaction conditions by pressure-swing between

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Nomenclature

Latin letters xi molar fraction of component in liquid phase (mol/mol) Rm ideal gas constant (8.314 J/mol/K) T reaction temperature (K) R reflux ratio (kg/kg) Nrec number of trays for rectifying section NRD number of trays for reactive distillation section NS number of trays for stripping section DP pressure difference (bar) NC5 position of C5 feed stream Q heat transfer rate (kW) U overall heat-transfer coefficient (kW/m2/K) A heat exchange area (m2) DT temperature difference (K)

2M1B 2M2B PSTC r-PSTC

2-methyl-1-butene 2-methyl-2-butene pressure-swing thermal coupled pressure-swing thermal coupled with reactive distillation TAC total annual cost CRD conventional reactive distillation RD reactive distillation IA isoamylene RDWDC reactive dividing wall distillation column MTBE methyl tertiary butyl ether MeOH methanol LH Langmuir–Hinshelwood UNIFAC group-contribution UNIQUAC universal quasichemical functional group activity coefficients

Abbreviations TAME tert-amyl methyl ether

reactive distillation section and stripping section divided from one RD column. A novel pressure-swing thermal coupled (PSTC) technology process is proposed for the reinforcement of heat-integration within a RD system. The simulation of TAME etherification process was employed to evaluate the design philosophy. Pressure-swing thermal coupled (PSTC) with reactive distillation (r-PSTC) process for TAME synthesis was optimized by sensitivity analysis. The intrinsic characteristics of the supplementary r-PSTC were revealed through comparison with CRD column. It is also desirable to know if the r-PSTC process could be controlled properly for practical use. Thus, the control strategy of the most favorable design flowsheet will also be developed and studied. 2. Conventional reactive distillation process of TAME TAME is a blending ingredient used in oxygen-containing gasoline as a replacement for lead. The interest in TAME increased after recent bans on the use of methyl tertiary butyl ether (MTBE) because of environmental concerns [16]. In the TAME reaction process, the light gasoline fraction from the fluid catalytic cracking unit is used as a source of isoamylene (IA). Luyben [17] summarized the TAME production by reactive distillation process, which was focused on the heat-integration of methanol recovery system followed by reactive distillation system. In this subsection, we will review the design of this process as a candidate for exploring the economic analysis of PSTC technology in a reactive distillation system. 2.1. Reaction kinetics and thermodynamics models Generally, TAME is formed by liquid phase, reversible, exothermic reaction from methanol (MeOH) with IA, consisting of the isomers 2-methyl-1-butene (2M1B) and 2-methyl-2-butene (2M2B), catalyzed by a sulfonic acid ion-exchange resin. In addition to the above two reactions, the two isoamylenes (2M1B and 2M2B) undergo isomerization. The networks scheme for the reactions is presented in Fig. 1. In the presence of an ion-exchange resin, the reaction kinetic is too complex to be described by a precise model. Heterogeneously catalyzed kinetics can be approximated by a pseudo-homogeneous model [18,19]. Several approximate kinetic models have been published for the synthesis of TAME, including Langmuir–Hinshelwood

Fig. 1. The reaction network scheme of the TAME system.

model (LH) [20], Eley–Rideal model [21], modified Eley–Rideal mechanism [22] and Lewis–Bronstedt model [23]. In the work of Kiviranta and Krause [24], several kinetic formulations were tested to explain the experimental data of the reaction rate of TAME synthesis, which showed the goodness-of-fit of LH mechanism based on the assumptions of stronger methanol adsorption. In this study, the kinetics for the forward and reverse of three reactions (Eqs. (1)–(3)), which quoted by Al-Arfaj and Luyben [25], are obtained from the Krause’s research group.

R1 ¼ 1:3263  108 e76:1037=Rm T x2M1B xMeOH  2:3535  1011 e110:541=Rm T xTAME

ð1Þ

R2 ¼ 1:3718  1011 e98:2302=Rm T x2M2B xMeOH  1:5414  1014 e124:994=Rm T xTAME

ð2Þ

R3 ¼ 2:7187  1010 e96:5226=Rm T x2M1B  4:2933  1010 e104:196=Rm T x2M2B

ð3Þ

In TAME etherification process, the vapor–liquid equilibrium is complex because of system non-ideality and the C5 feed stream contains 2M1B, 2M2B and largely inert components which consist of n-pentane, isopentane, 1-pentene and 2-pentene. The activity coefficient methods are often used for simulation of etherification

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Table 1 Wilson interaction parameters for activity coefficient calculation in the TAME reaction system. Components

Parameters

i

j

ðkij  kii Þ=R (K)

ðkji  kjj Þ=R (K)

MeOH MeOH MeOH MeOH MeOH MeOH 2M1B 2M1B 2M1B 2M2B 2M2B 2M2B TAME n-Pentane

2M1B 2M2B TAME n-Pentane Isopentane 1-Pentene 2M2B TAME n-Pentane n-Pentane TAME Isopentane n-Pentane 1-Pentene

1175.4029 1220.4715 275.8697 1077.7432 1336.1237 839.2257 12.0101 73.5807 8.3298 11.7977 46.4325 95.9996 137.5872 174.2314

165.5641 116.5275 309.8265 625.8090 552.1224 519.8504 21.9537 114.4251 35.0643 26.4362 85.6784 54.4444 53.8658 91.6395

reaction system, including group-contribution (UNIFAC) [16,17,25], UNIQUAC [26] and Wilson method [27,28]. The Wilson method can describe the compound activities very well in the mixture of methanol with the C5 hydrocarbons for the whole concentration range at boiling temperature [29]. In view of this, the liquid-phase activity coefficients are calculated by the Wilson method using the binary parameters, while the vapor pressure of pure components is used to account for association in the vapor phase. The parameters and coefficients in phase equilibrium calculation are listed in Table 1. 2.2. Base flowsheet of the CRD process The CRD process proposed by Subawalla and Fair [30] to produce TAME using a pre-reactor and a reactive distillation column is shown in Fig. 2. The RD column consists of a reactive section in the middle with non-reactive rectifying and stripping sections at the top and bottom, respectively. TAME is component of the highest boiling point, so it leaves in the bottoms stream from the reactive distillation column. Because of the existence of azeotropes between methanol and all the C5 components, the lighter C5 components leave in the distillate stream along with a significant amount of methanol. Herein, the feed conditions and column configuration proposed by Subawalla and Fair [30] are used to simulate the RD column for TAME synthesis. The simulation results agree well with the results from Subawalla and Fair. The results demonstrated that the model of RD process for TAME synthesis proposed by this paper is accurate enough to describe the actual process. The information such as process material balance, design and operating parameters for the pre-reactor and reactive distillation column are given in Fig. 2. Stages and segments are numbered from the top of the RD column. To the best of our knowledge, most of current research is focused on the methanol recovery system [16,17] with few studies reporting alternative design for reactive distillation column in TAME synthesis system. Hence, the present study will examine one alternative design of RDC with a novel PSTC process to improve the performance for TAME synthesis, as discussed in the next section. 3. Pressure-swing thermal coupled (PSTC) design

are not independent, which reduces the degree of freedom [31]. In addition, the pressures of reactive section and stripping section are almost the same except the difference caused by the pressure drop. In order to increase the reaction rate and ensure the reaction occur at liquid phase, the reactive distillation section needs to be operated at higher pressure than 4.7 bar, and the stripping section pressure should also be at the same value, much higher than the usual value of 0.55 bar. Generally, increasing column pressure reduces the relative volatility for distillation, resulting in the increase of energy consumption of reboiler at column bottom to provide a good recovery of 2M2B and the TAME purity in bottom product. Based on the above analysis, energy saving becomes very important for the pressure-swing thermal coupled technology with reaction distillation (r-PSTC), which is therefore the focus of this study. The novel design divides the RD column into two parts denoted as high-pressure and low-pressure columns, respectively. The former consists of reactive and rectifying sections while the latter is the stripping section. The vapor from the stripping section is pressurized before entering into the bottom of the reactive distillation section. Because the pressure of stripping section is descended, the temperature at the top of reactive distillation section is higher than the bottom of stripping section. If only the temperature difference is set to be greater than 10 °C for the heat transfer, the heat can be transferred into the bottom of low-pressure column. Consequently, the energy at reboiler of the low-pressure column that originally needs to be supplied by stream can be saved from this heat-integrated design. It should notice that the operating temperature of the reactive distillation section in the high-pressure column has to be less than 120 °C to be in compliance with the requirement of cation exchange resin catalyst. Establishing the optimum pressure in a reactive distillation column is complex because of the interplay between reaction and phase separation. For process of r-PSTC, the pressure of low-pressure column is fixed but pressure of the high-pressure column floats with changing conditions in order to keep the adequate temperature difference between the integrated reboiler and condenser. The boiling point ranking of the reaction system should not change due to operating the stripping section at lower pressure. This makes the reaction system like TAME, ETBE synthesis as a good candidate for this investigation. In the fully thermally coupled case, the condenser duty in the high-pressure column must equal to the reboiler duty in the low-pressure column. The two columns run in the ‘‘neat’’ mode without any auxiliary reboiler and auxiliary condenser, only using a compressor between the top of the low-pressure column and the bottom of the high-pressure column. In this paper, however, a partially thermal coupled case has been investigated, and the flowsheet is shown in Fig. 3. The design variables of the r-PSTC process (e.g. reflux ratio in the high-pressure column, number of stages in each section, pressure of low-pressure column) are also shown in Fig. 3. The temperature of the top of high-pressure column is 77 °C at 4.7 bar, and the base temperature of the low-pressure column is 66 °C at 0.55 bar, suggesting that heat integration could be attractive in terms of energy consumption. The heat input to the base of the low-pressure column is a portion (about 69%) of the heat removed in the condenser of the high-pressure column. The remaining heat is removed by an auxiliary condenser of the high-pressure column. The auxiliary condenser provided an additional control degree of freedom, so both a tray temperature in each column and a reflux ratio in high-pressure column could be independently manipulated.

3.1. Heat-Integrated via pressure-swing thermal coupled (PSTC) technology

3.2. Process optimization of r-PSTC

Both reaction and separation occur in a single vessel operation, therefore, the temperature and pressure of reaction and separation

The design variables in the flowsheet (Fig. 3) include the position of feed stream and the pressure difference (DP) between the

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Fig. 2. Flowsheet of the reactive distillation for the TAME process.

Fig. 3. Conceptual design of the r-PSTC system for TAME synthesis.

high-pressure and low-pressure columns. The objective function for the optimization search is the total annual cost (TAC). The TAC is defined by following Douglas [32] as:

TAC ¼ Operating cost þ

Capital cost þ Capital cost Payback year

 Annual interest rate

ð4Þ

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where the operating cost includes the cost of stream for the reboiler, cooling water for the condenser, and the operating cost for the compressor. The capital cost covers the cost of the column shell, internal trays, auxiliary condenser, heat exchanger (condenser/ reboiler) and the compressor. The calculation methods of operating and capital cost can be found from elsewhere [33,34], respectively. The compressor efficiency is assumed to be 0.72. A capital charge payback of 5 years and a 5% annual interest rate of the capital cost are assumed in this study. Similar to CRD column, there are two degrees of freedom for the r-PSTC system. The TAME composition at high-pressure column top is controlled to be less than 0.6 mol% by varying the reflux ratio. The TAME composition at low-pressure column bottom is set to be at more than 99.8 mol% by varying the heat exchange quantity of heat exchanger between the top of high-pressure column and bottom of low-pressure column. According to Luyben’s study [35], the overall heat-transfer coefficients of the heat exchanger can be assumed as 0.00306 GJ/h m2 °C. The heat duty and temperature difference in the condenser/reboiler are known from the result of the steady-state design. The heat transfer area of the condenser/reboiler is determined by the equation as follows:

Q ¼ U  A  DT

difference between the top of the high-pressure column and the bottom of the low-pressure column (DT). So the heat transfer area can be calculated and fixed as 852.52 m2. In the CRD process, the position of C5 feed stream is retained in the middle of the reaction section and the stripping section. In con-

ð5Þ

where Q is the heat transfer rate calculated from the fixed area (A), fixed overall heat-transfer coefficient (U), and temperature

Fig. 6. Temperature profile of r-PSTC process.

Fig. 4. Influence of DP on TAC at two different locations of C5 feed.

Fig. 7. Composition profile of liquid phase for r-PSTC process.

Fig. 5. Temperature profile of CRD process.

Fig. 8. Vapor and liquid flow rate profile of CRD process.

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trast, it could be either on the bottom of the high-pressure column or the top of the low-pressure column in the r-PSTC process. Although both two positions of C5 feed stream in the r-PSTC are the same as the condition of CRD process, two locations may have an impact on the TAC program, because it will affect the operating costs and investment costs of compressor. So the position of C5 feed, in the r-PSTC process, was investigated to obtain the minimum TAC. The temperature difference of the heat exchanger is adjusted by the the pressure difference (DP) between the high-

pressure and low-pressure columns. The design of DP is based on the principle that the minimum TAC is obtained as a trade-off between the cost of heat exchanger and the cost of the compressor. Fig. 4 displays TAC at different pressure differences (DP) between the high-pressure and low-pressure columns at two positions (NC5) of C5 feed stream. The calculation results shown that the position of C5 feed stream at the top of low-pressure column resulted in slightly lower TAC than at the bottom of high-pressure column. The minimum TAC was observed at DP of 4.15 bar with the temperature of the compressor suction port and discharge port is 31.2 °C and 103.1 °C, respectively. This can be explained by increasing the compressor cost more than decreasing the compressor cost in the TAC calculations, when the DP was more than 4.15 bar; In contract, when the DP was less than 4.15 bar, decreasing this number would add the heat exchanger capital cost due to heat transfer temperature difference of heat exchanger less than 10 °C.

3.3. Comparison of r-PSTC and CRD

Fig. 9. Vapor and liquid flow rate profile of r-PSTC process.

The temperature profiles of r-PSTC and CRD processes are shown in Figs. 5 and 6, respectively. The composition profiles of liquid phase for r-PSTC process are shown in Fig. 7. It is clear in Figs. 5 and 6 that the stage temperatures at the stripping section were descended by decreasing the operating pressure so that heat integration is possible between the reactive distillation section and the stripping section. The minimum temperature difference is about 11 °C, which makes the heat exchange feasible. The liquid and vapor flow rates inside the column of CRD and r-PSTC

Table 2 Comparison of design and cost information of CRD and r-PSTC. Parameters

CRD

r-PSTC

Total no. of trays No. of trays in rectifying section (Nrec) No. of trays in reactive distillation section (NRD) No. of trays in stripping section (Nrec) C5 feed flow rate (kmol/h) Reflux ratio Molar ratio of C5 feed (2M1B/2M2B/MeOH/TAME) Component flow rate in feed (kmol/h) (MeOH/2M1B/2M2B/TAME/n-pentane/ isopentane/1-pentene/2-pentene) Feed position of C5 feed (NC5) MeOH feed flow rate (kmol/h) Feed position of MeOH feed (NMeOH) Top product flow rate (kmol/h) Component flow rate in top product (kmol/h) (MeOH/2M1B/2M2B/TAME/npentane/isopentane/1-pentene/2-pentene) Bottom product flow rate (kmol/h) Component flow rate in bottom product (kmol/h) (MeOH/2M1B/2M2B/TAME/n-pentane/isopentane/1-pentene/2-pentene)

33 4 19 10 1195.85 1.5 1/8.8/16.3/16.4 156.0/9.55/83.94/157.0/88.4/501.2/ 38.06/161.7 29 195 24 1054.0 257.97/0.358/0.0430/6.262/88.34/ 501.09/38.05/161.59 243.9

33 4 19 10 1195.85 1.5 1/8.8/16.3/16.4 156.0/9.55/83.94/157.0/88.4/501.2/ 38.06/161.7 29 195 24 1054.0 257.46/0.284/0.0298/6.458/88.40/ 501.19/38.06/161.58 243.9

Trace/0.0025/0.059/243.06/0.055/ 0.105/0.0096/0.103 3.36 570.78 626.11 0 -18.717 19.327 0 19.327 1858.63 0 0 2412.69 743.74 3156.43 0 5 5172.88 0

0.385/trace/trace/243.73/trace/trace/ trace/0.103 H:2.73/L:3.86 172.03 0 812.52 -5.846 0 4.967 4.967 1223.48 34.17 548.72 2363.51 525.97 3438.20 -8.93 5 4833.59 6.56

Column diameter (m) The heat transfer area of condenser (m2) The heat transfer area of reboiler (m2) The heat transfer area of condenser/reboiler (m2) Condenser duty (MW) Reboiler duty (MW) Compressor duty (MW) Total heat duty (MW) Operating cost ($1000/year) Operating cost saving (%) Compressor cost ($1000/year) Column cost ($1000/year) Heat exchanger cost ($1000/year) Capital cost ($1000/year) Capital cost saving (%) Annual interest rate of the capital cost (%) TAC ($1000/year) TAC saving (%)

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processes are shown in Figs. 8 and 9, respectively. Compared to the CRD process, the r-PSTC process could reduce the vapor flow rate at reboiler from 2500 to 1200 kmol/h and thereby reduce the reboiler duty of the column bottom. With reference to Fig. 3 where the detailed information of optimal flowsheet was given for the reactive distillation process using PSTC technology, the comparison results of r-PSTC to CRD design without heat integration are outlined in Table 2. It is noticed that the operating cost saving of the r-PSTC design is 34.17%. Although the annualized capital cost of the r-PSTC design more than CRD design due to the high compressor cost, the TAC is still lower than CRD design by 6.56%.

including the control of three liquid levels (the decanter of the high-pressure column, the bottom of the high-pressure column and the bottom of the low-pressure column), the top pressures of

4. Control strategy of r-PSTC process Although PSTC process is an efficient way of energy saving in the chemical process industry, it may pose great challenges to the process operation due to the high degree of integration between the high-pressure column and low-pressure column. In order to make the process with high operation resilience, efficient control systems are required. The objective of control system is to make sure we convert the desired amount of isoamylenes and to product the TAME with the desired purity. Only tray temperature control loops will be used in the overall control strategy for wider chemical industrial applications [9]. The control system configuration for the r-PSTC design flowsheet is shown in Fig. 10. Because a constant pressure is kept in the lowpressure column, it is rational to adopt a stage temperature control scheme for the bottom product purity. In the high-pressure column, the product concentrations in the top are controlled by manipulating the reflux ratio. Although both reflux ratio and reflux-to-feed ratio were found to control the product concentrations, reflux ratio was selected because it is desirable to take the ratio of the reactant in the reaction zone. In the RD case, it is necessary to control the temperature of reactive distillation zone at constant value to ensure the performance of separation, the reaction rate and conversion of IA. Therefore, control loop with only tray temperature is also adopted to control the reactive performance of high-pressure column. The controllers are tuned using the Tyreus–Luyben tuning method. The relay-feedback test is used to find the ultimate gain and ultimate period. Firstly, the inventory control loops are determined. In PSTC process for TAME synthesis, there are nine inventory control loops

Fig. 11. Open-loop sensitivity plot for high-pressure column.

Fig. 12. Open-loop sensitivity plot for low-pressure column.

Fig. 10. Control scheme of the r-PSTC system for TAME.

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the high-pressure column and low-pressure column, the flow rate of the C5 feed and the fresh methanol feed, the tray temperature at low-pressure column and high-pressure column. The inventory loops are arranged as follows: The reflux drum level of high-pressure column is controlled by manipulating the product flow at top of high-pressure column, the bottom level of high-pressure column is controlled by manipulating its bottom flow and the bottom outlet level of low-pressure column is controlled by manipulating the final TAME product flow. The top pressure of the high-pressure column is controlled by manipulating the auxiliary condenser duty. Because there is no condenser in the top of the low-pressure column, the top pressure of the low-pressure column is controlled by manipulating the vapor flow of the top. The C5 feed flow rate is under flow control and assigned as the throughput manipulator to handle production changes. The methanol feed flow rate is maintained at a floating ratio with the flow rate and the composition of C5 feed. The C5 feed composition measurement was used to feedforward composition information to trim up the ratio of the two feeds flow rate. This ratio is calculated by the stoichiometric ratio of the reaction and the azeotropes composition of methanol/C5. The remaining manipulated variable is the reflux ratio of the high-pressure column which can be used to control one tray

475

temperature at high-pressure column. The tray temperature of low-pressure column can be manipulated by the flow of the vapor from the top of the high-pressure column. The flow of vapor from the top of the high-pressure column to the integrated reboiler/condenser would change along with the heat duty of the bottom of the low-pressure column to supply the heat called for by the temperature controller. This is a temperature/flow cascade control structure. The deadtime of 1 min on the flowsheet between the column and the temperature controller is inserted. Open loop sensitivity analysis is used for determining these two tray temperature control points. Figs. 11 and 12 show the temperature changes in the high-pressure column (with ±0.5% changes of the reflux ratio) and low-pressure column (with ±0.5% changes of the vapor flow to the integrated reboiler/condenser), respectively. The largest temperature deviation was found at the 16th and 6th stage for the low-pressure and high-pressure columns, respectively. These results suggested that the 16th stage temperature in the high-pressure column and the 6th stage temperature in the low-pressure column are controlled to maintaining the performance of r-PSTC process for TAME synthesis. Three types of disturbances were used to test the proposed control strategy. The first one was the feed flow rate changes. To make

Fig. 13. Closed-loop responses with ±20% C5 feed flow rate changes.

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these changes, the set-point of the C5 feed flow rate was changed by ±20% at 0.5 h. With proper overall control strategy, all variables in the design flowsheet should be modified accordingly but still maintain the bottom product composition of low-pressure column and the top product composition of high-pressure column. Fig. 13 shows the two tray temperature control loops in the proposed control strategy. The solid lines are +20% step changes. The dashed lines are 20% changes. Stable regulatory control is achieved for these large disturbances. All temperature control loops performed well by forcing the temperatures to quickly return to their setpoints. The IA compositions at the two products flow are all maintained at very low values which indicate nearly complete conversion of IA. The two product flow rates are all increased or decreased as planed with the two product compositions maintained at desired values. The second disturbance is the C5 feed composition changes. The ±50 mol% changes of IA in the C5 feed compositions were introduced at 0.5 h. For these unmeasured disturbance changes, the control strategy should allow the methanol feed flow to increase or decrease accordingly to compensate the IA composition changes so that the stoichiometric ratio between IA and methanol and the

ratio of azeotropes between methanol and insert composition into the system is maintained. Fig. 14 shows that the methanol feed flow rate can be increased or decreased to desirable values. The IA compositions at the two product flows are all maintained at very low values which indicate the conversion of IA is achieved nearly complete. All products compositions are maintained at desired values. The third disturbance is the fresh methanol feed composition changes. The response of the system for the composition of the methanol feed changed from 100% to 95% as shown in Fig. 15. The 5 mol% change in the fresh methanol feed composition is introduced at 0.5 h. The 5 mol% change means that the feed composition is changed from pure methanol to 95 mol% methanol and 5 mol% inert C5. For this unmeasured disturbance change, the control strategy should allow the fresh methanol to increase accordingly to compensate the amount of the methanol decreased. From the first row and third column of Fig. 15, it is shown that the fresh methanol flow rate increased to desirable values. The top product flow rate of the high-pressure column is increased because of the extra inert C5 impurity coming into the system through the methanol feed stream. Two products compositions

Fig. 14. Closed-loop responses with ±50 mol% IA composition in C5 feed changes.

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Fig. 15. Closed-loop responses with 5 mol% methanol composition in methanol feed changes.

are all maintained at desired values. These results demonstrate that the r-PSTC process would provide good dynamic controllability by using simply tray temperature control strategy. 5. Conclusions The process of reactive distillation using r-PSTC technology for TAME synthesis from 2M1B and 2M2B and methanol was described. Design configuration and control strategy of this technology were investigated. The optimal design flowsheet of r-PSTC is obtained by minimizing the total annual cost (TAC) of the processes. Two variables on r-PSTC system are studied, including the position of feed stream and the pressure difference (DP) between high-pressure column and low-pressure column. Results showed that TAC is lowest at DP of 4.15 bar. The position of C5 feed stream is set to be the top of low-pressure column. By combining the reboiler of the low-pressure column and the condenser of the high-pressure column into one heat-exchanger, the energy consumption was reduced. Due to expansive compressor is necessary in r-PSTC process, the capital cost was 8.93% higher than the conventional RD design. Furthermore, the operating cost and TAC of the r-PSTC design were reduced by 34.17% and 6.56%, respectively.

The overall control strategy of the r-PSTC system is proposed with only one tray temperature control loop in each column. Large variations in the feed composition and also throughput changes can be handled by this proposed control strategy. Results also indicated that all products compositions could be maintained at appointed purity despite feed disturbances.

Acknowledgements The authors are grateful for the financial support from the Program for Changjiang Scholars and Innovative Research Team in University (No. IRT0936) and National Natural Science Foundation of China (Nos. 21336007 and 21176172).

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