Dynamic Simulation and Control of a Complete Industrial Acetic Acid

Article pubs.acs.org/IECR

Dynamic Simulation and Control of a Complete Industrial Acetic Acid Solvent Dehydration System in Purified Terephthalic Acid Production Qianlong Li, Wei Guan, Lei Wang, Hui Wan,* and Guofeng Guan* College of Chemical Engineering and State Key Laboratory of Materials-Oriented Chemical Engineering, Nanjing Tech University, Nanjing 210009, P. R. China ABSTRACT: In the production of purified terephthalic acid (PTA), the acetic acid solvent dehydration system is a typical heterogeneous azeotropic distillation system. In this work, a rigorous process simulation of the complete industrial acetic acid solvent dehydration process is conducted. To achieve both the bottom- and top-product specifications, remove the accumulated p-xylene (PX), and recycle the methyl acetate (MA) despite feed flow rate disturbances, a suitable control strategy of this complete system is proposed. The control strategy requires only a side draw and organic reflux to control the tray temperatures of the dehydration column, the bottom heating steam to control the tray temperature of the PX purification column, and topproduct reflux to control the tray temperature of the entrainer recovery column; this strategy can be easily implemented in industry for wider applications. From dynamic simulation and analysis, the control strategy can maintain the product purities, overcome the drawback of the PX imbalance, achieve the transition to the desired operation more quickly, and better maintain the stability of the overall solvent dehydration system under a disturbance in the upper-zone feed streams. Tyreus23 first offered an azeotropic distillation column with a decanter for academic researchers pursuing control studies. Afterward, a nonlinear control system for the acetic acid dehydration column with NBA as the entrainer was proposed. In that study, a complicated exact input−output linearization controller was used. 15 Chien et al.18 investigated the dehydration column (including only three components). The proposed overall control strategy was very simple, requiring only one tray-temperature control loop inside the solvent dehydration column. Then, the optimum side-stream location and flow rate were designed (containing an impurity for dehydration), and an automatic purging strategy was proposed to prevent the accumulation of the impurity inside the column.22 For impurity disturbances, the side-stream flow rate needed to be manipulated accordingly. A novel side-stream operating strategy to maintain the impurity concentration inside the column was proposed for energy-efficient operation.24 However, very few researchers have discussed overall process control of this complex system (including a dehydration column, a PX purification column, a wastewater separation column, and an entrainer recovery column). The dynamic simulation and analysis of the entire process are more important for the control and optimization of the solvent dehydration system, thanks to the existence of serious interactions among the operating units. In this work, the design and control of an industrial column for a solvent dehydration system with a feed impurity were investigated. The entrainer used for this industrial column system was NPA. Multiple column feed streams from various parts of the upstream process were fed into this system.

1. INTRODUCTION Heterogeneous azeotropic distillation is used in industry to break azeotropes and separate mixtures with close relative volatilities, such as those encountered in ethanol or isopropyl alcohol dehydration and acetic acid dehydration (HAD).1−4 In the commercial production of purified terephthalic acid (PTA), p-xylene (PX) is catalytically oxidized with air in a reactor in the presence of acetic acid (HAc) solvent. The outcoming HAc solvent containing mainly three components [methyl acetate (MA), PX, and H2O] is fed to a solvent dehydration system to recover high-purity HAc and remove the oxidation byproducts. This system is a typical heterogeneous azeotropic distillation system.5−8 Some researchers have found that parametric sensitivity, multiple steady states, long transients, and nonlinear dynamics can exist in this system,9−11 which make this system difficult to design and control.12,13 The dehydration of acetic acid solvent by heterogeneous azeotropic distillation has been studied for several years. A variety of entrainers have been selected for the system, such as n-propyl acetate (NPA),14 n-butyl acetate (NBA),15,16 isobutyl acetate (IBA),13,17,18 ethyl acetate (EA),19 and p-xylene (PX).20 Wasylkiewicz et al.21 used residue curve maps and geometric methods to design an HAD column with NBA as the entrainer. Chien et al.22 investigated the influence of PX feed impurity on the design and operation of an industrial HAD column. Wang and Huang20 presented an economical design of an HAD column using PX as the entrainer with a side stream to purge the PX accumulated in the column. Recently, considering PX and methyl acetate (MA) as feed impurities existing in the solvent of acetic acid, an industrial-scale complete heterogeneous azeotropic distillation column system using n-propyl acetate (NPA) as the entrainer was developed.5 Most of the preceding works on the HAc dehydration system involved process synthesis steady-state design, but a relatively small amount of work has been reported on the dynamics and control of the solvent dehydration system. Luyben and © 2015 American Chemical Society

Received: Revised: Accepted: Published: 11330

January 29, 2015 October 11, 2015 October 28, 2015 October 28, 2015 DOI: 10.1021/acs.iecr.5b00390 Ind. Eng. Chem. Res. 2015, 54, 11330−11343

Article

Industrial & Engineering Chemistry Research Table 1. Parameter Values for the UNIQUAC Model component i

component j

aij

aji

bij

bji

HAc HAc HAc HAc H2O MA NPA PX MA H2O

PX H2O MA NPA MA NPA PX H2O PX NPA

3.957 0.745 0 −6.011 2.538 −19.485 0 0 −5.884 0

−9.981 0.004 0 12.541 −18.277 −0.320 0 0 1.900 0

−1519.200 −752.800 19.723 2439.448 −688.447 5921.619 120.006 −4327.188 1861.191 −96.727

3608.611 263.590 −67.295 −5103.140 5464.175 433.949 −169.884 80.714 −676.366 −396.795

recovery column (C-104). Five streams are fed into C-101 for use of NPA as the entrainer so that the top distillate stream 9 would contain 0.01 wt % HAc and the bottom stream 8 would contain 93 wt % HAc. Feed streams 3, 4, and 6 and the bottom reboiler provide the heat needed in C-101. After condensation to a temperature of 83 °C, the top stream 9 of C101 enters D-101. NPA makeup stream 14 and NPA recycle stream 16 are also fed to D-101, and then the mixed solution is separated into two liquid phases. Top organic phase stream 10, mostly consisting of NPA, is refluxed back to C-101, whereas aqueous phase stream 17 is sent to C-103 for further treatment. Noncondensable gas stream 15 is sent to C-104. PX accumulated in C-101 weakens the column performance and the separation efficiency, so it is necessary to draw side-draw stream 12 from the upper zone of C-101,29 which is sent to C102 to remove the accumulated PX in C-101. Stream 7 from stream 5 provides the heat needed for C-102. At the same time, wastewater steam 18 is introduced to break the azeotrope and recover NPA in C-102. The top distillate steam 11 (recover NPA) returns to C-101, and the bottom stream 13 containing mainly PX and HAc goes back to the upstream process. Aqueous phase stream 17 enters the top of C-103, which has no reboiler but uses live steam. The top vapor phase goes to C104, and bottom stream 13 (close to pure water) is sent to a sewage treatment plant. The top vapor phase from C-103 and noncondensable gas stream 15 from D-101 go to C-104. The overhead vapor phase of C-104 goes into partial condenser HE-103. Top liquid-phase stream 23 of C-104, containing mainly MA, is sent to the mother liquor tank and go back to the reactor. The noncondensable gas stream 22 is sent to atmospheric washing column. Bottom liquid-phase stream 16 of C104 returns to D101 for NPA recycling.

Optimum column designs and operating conditions were obtained for this system by rigorous process simulation. A suitable control strategy of this complete system was proposed to achieve both bottom- and top-product specifications, remove the accumulated p-xylene (PX), and recycle the methyl acetate (MA) despite feed flow rate disturbances. The control strategy used only a side draw and organic reflux to control the tray temperatures of the dehydration column, which was different from the strategy proposed by Wang et al.,13 which used the reboiler duty to control the tray temperature for this system. The bottom heating steam to control the tray temperature of the PX purification column and top-product reflux to control the tray temperature of the entrainer recovery column were considered so that the results of this control strategy can be easily used directly in industry.

2. THERMODYNAMIC MODEL The feed system consisted of HAc, H2O, PX, MA, and NPA. The universal quasichemical (UNIQUAC) activity coefficient model was used for the vapor−liquid−liquid equilibrium (VLLE). Table 1 lists the UNIQUAC model parameter set used in this work. The sets of UNIQUAC parameters (HAc + H2O, HAc + NPA) were obtained from our previous work.25 The UNIQUAC parameters for H2O + NPA and H2O + PX originated from the Aspen Plus built-in literature parameters. The other UNIQUAC parameters were obtained from Perelygin and Volkov26 and Huang et al.5 The Hayden and O’Connell27 second virial coefficient model with association parameters was used to account for the dimerization of HAc in the vapor phase. Table 2 lists the Hayden and O’Connell model parameters obtained from the Aspen Plus built-in binary parameters.28 3. PROCESS DESCRIPTION Figure 1 shows the layout of an industrial solvent dehydration system process, which contains a dehydrating column (C-101), a decanter (D-101), a PX purification column (C-102), a wastewater separation column (C-103), and an entrainer

4. STEADY-STATE DESIGN Table 3 lists information on the feed streams coming from the upstream process. It can be observed that there are four components in the feed streams (HAc, H2O, PX, MA). The commercial process simulation software Aspen Plus28 was used to simulate the steady-state flowsheet. The high nonideality of the phase equilibrium system makes the convergence of the recycle streams very difficult. Dehydrating column C-101 is very sensitive because of the existence of multiple steady states in this system.30−32 The simulation results for columns C-101, C-102, C-103, and C-104 are summarized in Tables 4−7, respectively.

Table 2. Parameter Values for the Hayden−O’Connell Model

HAc H2O PX MA NPA

HAc

H2O

PX

MA

NPA

4.5 2.5 0.4 2 2

2.5 1.7 0 1.3 1.3

0.4 0 0 0.6 0.6

2 1.3 0.6 0.85 0.53

2 1.3 0.6 0.53 0.53 11331

DOI: 10.1021/acs.iecr.5b00390 Ind. Eng. Chem. Res. 2015, 54, 11330−11343

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Industrial & Engineering Chemistry Research

Figure 1. Process flowsheet of an industrial HAD process.

Table 3. Feed-Stream Information for the Heterogeneous Azeotropic Distillation Column

Table 4. Design Values: Dehydrating Column C-101 total stages feed stages

stream no. PX (kg/h) HAc (kg/h) H2O (kg/h) MA (kg/h) flow rate (kg/h) temperature (°C) phase

1

2

3

4

5

0 1255 2913 18 4186 40 liquid

42 6469 4848 189 11548 54 liquid

31 12166 4352 186 16735 120 vapor

0 15812 4572 91 20475 130 vapor

0 12049 2685 61 14794 156 vapor

25 1/2/11 at sixth stage 3/4/6 at 18th stage 10 at first stage

temperature (°C) top bottom pressure (atm) top bottom reboiler duty (kW) condenser duty (kW) column diameter (m) diameter/length (m) base flow rates (kg/h) 1/2/3/4/6 8/9/10/11 14

5. DYNAMICS AND CONTROL The solvent dehydration system process is very complex, with serious interactions among the operating units. For example, when the feed stream flow rate has load disturbances, the C-102 bottom PX flow rate changes, and the C-103 amount of live steam also changes. Meanwhile, the C-104 reflux flow rate is affected. To obtain good dynamic performance, the commercial process simulator Aspen Dynamics was used for the dynamic simulation, and overall control strategies for this system were developed despite changes in the feed flow rate. Before the Aspen Plus steady-state simulation in the last section was exported to the dynamic simulation of Aspen Dynamics, the sizing of the equipment was needed to convert the steady-state simulation to a dynamic one. The pack sizing option in Aspen Plus was utilized to calculate the diameters of columns C-101, C-103, and C-104 as 4, 0.9, and 1.4 m. Information about the packing is listed in Table 8. The tray sizing option in Aspen Plus was utilized to calculate the diameter of column C-102 as 0.6 m. The weir height of each tray in every column was assumed to be 0.05 m. Other equipment sizing values recommended by Luyben33 were used in this work. The volumes of the sump and reflux drum were sized to give 10-min holdup with 50% liquid level. The decanter

stream(s)

PX

1/2/3/4/5 11 8 9 10

a 0.186 0 0.003 0.003

86.3 118.7 1.11 1.3 14670 21404 4 3.2/6.4 4188/11548/16736/20475/12784 49000/120525/104165/2057 4 Composition (wt %) HAc

H2O

MA

NPA

a 0.032 0.935 0 0 Decanter

a 0.218 0.065 0.158 0.030

a 0.005 0 0.035 0.035

a 0.559 0 0.804 0.932

temperature (°C) diameter/length (m) a

83 2.2/4.4

See Table 3.

was sized to be larger to allow for the two liquid phases to separate, so a holdup time of 20 min was used in the dynamic simulation. Most valves were specified to have pressure drops of 0.02 atm. Then, the Aspen Plus file was exported into Aspen 11332


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Industrial & Engineering Chemistry Research Table 5. Design Values: PX Purification Column C-102

Table 7. Design Values: Entrainer Recovery Column C-104

total stages feed stages

total stages feed stages

16 18 at first stage 12 at first stage 7 at 16th stage

temperature (°C) top bottom pressure (atm) top bottom reboiler duty (kW) column diameter (m) diameter/length (m) base flow rates (kg/h) 7/12/18/13

reflux ratio condenser duty (kw) temperature (°C) top bottom pressure (atm) top bottom reboiler duty (kW) column diameter (m) diameter/length (m) base reflux drum flow rates (kg/h) 15/16/22/23

93.4 112.2 1.2 1.4 0 0.6 0.9/1.8 2010/2300/507/2755 Composition (wt %)

stream

PX

HAc

H2O

MA

NPA

7 12 18 13

0 0.199 0 0.026

0.816 0.156 0 0.702

0.180 0.143 1 0.272

0.004 0.002 0 0

0 0.500 0 0

Table 6. Design Values: Wastewater Separation Column C103 total stages feed stages

a

45 83.3 1.04 1.09 0 1.4 1.7/3.4 1.6/3.2 3366/6280/287/229 Composition (wt %)

stream

PX

HAc

H2O

MA

NPA

15 16 22 23

0.001 0 0 0

0 0 0 0

0.149 0.565 0.015 0.027

0.105 0.015 0.98 0.957

0.745 0.420 0.005 0.016

Table 8. Characteristic Data of Mellapak Packing 250Y

6 17 at first stage 19 at sixth stage

temperature (°C) top bottom pressure (atm) top bottom reboiler duty (kW) column diameter (m) diameter/length (m) base flow rates (kg/h) 17/19/24/V

12 15 at eighth stage 16 at 13th stage 28 2640

100.4 102.2 1.1 1.2 0 0.9

19273/3500/18839/3430 Composition (wt %)

PX

HAc

H2O

MA

NPA

17 19 24 Va

0 0 0 0

0 0 0 0

0.980 1 1 0.887

0.012 0 0 0.067

0.008 0 0 0.046

value 250 45 0.5 12 97

reboiler duty was fixed and maintained at a constant ratio to the fresh feed flow rate. To achieve the top- and bottom-product specifications, an organic reflux flow rate was used to control the tray temperature. D-101: The entrainer makeup flow rate controlled the organic phase level in the decanter, as is done in most industrial applications.34 The aqueous-product flow rate controlled the aqueous-phase level in the decanter. The decanter temperature was controlled at 83 °C by manipulating the condenser duty. The decanter pressure was controlled at 1.1 atm by manipulating the vapor stream flow rate. C-102: The column top pressure was controlled at 1.2 atm by manipulating the top vapor flow rate. The bottom feed flow rate was used to control the tray temperature. The bottom product flow rate (containing mainly PX and HAc) controlled the column bottom level. C-103: The bottom wastewater flow rate controlled the column bottom level. The water vapor flow rate was kept at a constant ratio to the aqueous-phase flow rate. C-104: The column top pressure was controlled at 1.04 atm by manipulating the top vapor flow rate. The reflux drum temperature was controlled at 45 °C by manipulating the HE103 duty. The top-product flow rate controlled the reflux drum level.The bottom product flow rate controlled the column bottom level. The column top-product reflux flow rate controlled the tray temperature.

1.7/3.4

stream

parameter surface area (m2/m3) ripple tilt angle (deg) thickness (mm) peak height (mm) porosity (%)

Top vapor stream from C-103 to C-104.

Dynamics. The default control scheme was modified as discussed in the following section. Two inventory control strategies were used for this system. The first inventory control strategy (CS1) used a constant sidedraw flow rate of 2300 kg/h. The second inventory control strategy (CS2) maintained the tray temperature of the side draw by manipulating the side-draw flow rate. Other inventory control loops that used the same pairings for either of these strategies were as follows: C-101: The column top pressure was controlled at 1.11 atm by manipulating the top vapor flow rate. The bottom HAc product flow rate controlled the column bottom level. The 11333


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Figure 2. Proposed overall control strategy (CS2).

The overall control strategy for the base case of CS2 is shown in Figure 2. 5.1. Control at Base-Case CS1. The sensitivity analysis with small perturbations of the manipulated variables was subsequently performed to determine the tray-temperature control points. As the dynamic simulation analysis was more in line with the actual situation, the dynamic simulation analysis was more suitable than the steady-state simulation analysis for determining the control point. In this work, the controlled stage temperatures were also selected by an open-loop sensitivity analysis. Figure 3 shows the open-loop sensitivity between organic reflux and stage temperature changes (C-101). One discovery was seen from this figure: the temperature at the eighth stage, which was closer to the top of the column (C101), was selected as the controlled variable because it exhibited the greatest sensitivity to the organic reflux. Figure 4 shows the open-loop sensitivity between feed stream 7 and the stage temperature changes (C-102). The temperature at the fourth stage, which was closer to the bottom of the column (C102), was selected. Figure 5 shows the open-loop sensitivity between changes in the reflux flow rate and stage temperature (C-104). There was a sharp temperature break at about the fourth stage, which implied that this stage temperature control should be effective, so the fourth stage temperature was selected as the controlled variable, which was closer to the top of the column (C-104). In this control system, most of the loops discussed above were standard distillation control strategies. All level loops were proportional-only with KC = 2, as suggested in Luyben.32 All flow-rate and pressure controllers were proportional−integral

Figure 3. Open-loop sensitivity analysis between organic reflux and stage temperature changes (C-101).

(PI) and used the Aspen Dynamics default values. A 1-min dead time was inserted in each temperature loop. The loops were tuned individually by running a relay−feedback35 test with the other two controllers on manual control and using the Luyben−Tyreus settings. The resulting PI tuning constants for the tray 8 temperature loop of C-101 were KC = 4.9 and τI = 39.6 min. The tuning constants for the tray 4 temperature loop 11334


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returned to specification, whereas the bottom PX flow rate took a long time to cross its final operating value. Figure 6i,j shows the dynamic responses of the bottom (C-103) water concentration and the stream 19 flow rate (providing energy for C-103 and C-104). Because the stream 19 flow rate was fixed at a constant ratio to the feed stream 17 flow rate, the increase/decrease of the stream flow rate depended on the changes in the aqueous-phase stream flow rate. The stream flow rate crossed its final operating value after 100 h. The plot of the bottom (C-103) water concentration indicates that the water removed from this system met the specification and was almost pure water. From Figure 6k,l, one can see the dynamic responses of fourth stage temperature and reflux flow rate (C104). The stage temperature and reflux flow rate (C-104) could return to relatively stable specification after a period of time. The closed-loop dynamic responses of control structure CS1 to ±5% changes in the feed 2 flow rate are shown in Figure 7. The disturbances were also introduced at a time of 2 h. Figure 7a,b shows that the eighth stage temperature quickly returned to specification. However, the organic reflux was not steady at the final simulation time of 100 h under a +5% change in the feed 2 flow rate. Figure 7c,d shows that the bottom and top HAc purities were both maintained at the designed values. Figure 7e,f shows that the increase/decrease of the reboiler heat duty depended on the changes in the feed stream flow rate, because the reboiler duty was fixed at a constant ratio to the fresh feed flow rate. The side-draw flow rate was fixed at a constant value. Figure 7g,h shows that the 12th stage temperature quickly returned to specification, whereas the bottom PX flow rate took a long time to cross its final operating value after 80 h. From Figure 7i,j, one can see that the stream flow rate took a long time to cross its final operating value under the +5% change in feed flow rate. However, with the −5% change in feed flow rate, the stream flow rate was not steady at the final simulation time of 100 h, but the bottom water concentration met the specification and was almost pure water. It can be seen from Figure 7k,l that the fourth stage temperature and reflux flow rate could return to relatively stable specification after a period of time. The closed-loop dynamic responses of control structure CS1 to ±5% changes in the flow rate of feed 3 are shown in Figure 8. The disturbances were also introduced at a time of 2 h. Figure 8a,b shows that the eighth stage temperature and the organic reflux could returned to specification after a while. Figure 8c,d shows that the bottom and top HAc concentrations were not steady at the final simulation time of 100 h. Figure 8e,f shows that the reboiler duty and side-draw flow rate were fixed at constant values. Figure 8 g,h shows that the 12th stage temperature quickly returned to specification, whereas the bottom PX flow rate was not steady at the final simulation time of 100 h under a +5% change in the feed 3 flow rate. Figure 8i,j shows that the stream flow rate was not steady at the final simulation time of 100 h, whereas the bottom water concentration met the specification and was almost pure water. Figure 8 k,l shows that the fourth stage temperature could return relatively stable specification after a period of time, but the reflux flow rate did not reach a new steady-state operating value at the final simulation time. 5.2. Control at Base-Case CS2. The closed-loop dynamic responses of control structure CS1 to ±5% changes in the flow rate of feed 1 are shown in Figure 9. The disturbances were introduced at a time of 2 h. The eighth stage temperature quickly returned to specification, and the organic reflux also

Figure 4. Open-loop sensitivity analysis between feed 7 and stage temperature changes (C-102).

Figure 5. Open-loop sensitivity analysis between changes in the reflux flow rate and stage temperature (C-104).

of C-102 were KC = 1.0 and τI = 20 min, and those for the tray 4 temperature loop of C-104 were KC = 21.2 and τI = 39.6 min. The closed-loop dynamic responses of control structure CS1 to ±5% changes in the flow rate of feed 1 are shown in Figure 6. The disturbances were introduced at a time of 2 h. The controlled and manipulated variables of the important control loop [eighth-stage temperature and organic reflux loop (C101)] are shown in Figure 6a,b. The stage temperature returned to its final operating value. However, the organic reflux was not steady at the final simulation time of 100 h. Figure 6c,d shows the dynamic responses of the bottom (C-101) and top HAc product concentrations. The bottom and top HAc purities were maintained at the designed value. In addition, Figure 6c,d indicates that the HAc concentration at the bottom decreased with increasing feed flow rate and increased with decreasing feed flow rate, whereas the top HAc purity exhibited the opposite trends. Figure 6e,f shows that the reboiler duty and side-draw flow rate were fixed at constant flow rate (C-101). Figure 6g,h shows the dynamic responses of the 12th stage temperature and bottom PX flow rate (C-102). From these plots, one can see that the stage temperature (C-102) quickly 11335


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Figure 6. Dynamic responses using control structure CS1 to ±5% changes in the feed 1 flow rate.

reboiler duty was fixed at a constant value, and the side-draw flow rate was different from that in the front control structure CS1, because the side-draw flow rate manipulated the sixth tray

crosses its final operating value in Figure 9a,b. Figure 9c,d shows that the bottom and top HAc purities were both maintained at the designed values. Figure 9e,f shows that the 11336


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temperature of side draw. Figure 9g,h shows that the 12th stage temperature quickly returned to specification after 8 h, and the bottom PX flow rate was steady at a simulation time of 50 h.

Figure 9i,j indicates that the water removed from this system met the specification, the stream flow rate could cross its final operating value under ±5% changes in the feed flow rate, and 11337


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the increase/decrease of the stream flow rate depended on the changes in the aqueous-phase stream 17 flow rate. Figure 9k,l

shows that the fourth stage temperature and reflux flow rate could return to stable specification after a period of time. 11338


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to specification after 15 h. The organic reflux crossed its final operating value after 12 h. Figure 10c,d shows that the bottom and top HAc purities were both maintained at the designed

The dynamic responses of control structure CS2 to ±5% changes in the flow rate of feed 2 are shown in Figure 10. Figure 10a,b shows that the eighth stage temperature returned 11339


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values. Figure 10e,f shows that the increase/decrease of the reboiler heat duty depended on the changes in the feed-stream flow rate and the side-draw flow rate crossed its final operating

value after 25 h. Figure 10g,h shows that the dynamic responses of 12th stage temperature returned to specification after 8 h and the bottom PX flow rate returned to specification after 40 h. 11340


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flow rate could return to stable specification after a period of time. The dynamic responses of control structure CS2 to ±5% changes in the flow rate of feed 3 are shown in Figure 11.

Figure 10i,j indicates that the water removed from this system met the specification, and the stream flow rate could cross its final operating value under ±5% changes in the feed flow rate. Figure 10k,l shows that the fourth stage temperature and reflux 11341


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and 11, we found that CS2 removes the accumulated PX was much faster than CS1. However, the bottom product and the column top aqueous outlet streams did not meet their specifications, the stream flow rate could not cross its final operating value after 100 h, and the reflux flow rate did not reach a new steady-state operating value under a −5% change in feed 3 flow rate under the two control structures (CS1, CS2). Considering the whole solvent dehydration system, the dynamic performances of CS2 were much more stable under disturbances of upper-zone feed streams.

Figure 11a,b shows that the eighth stage temperature returned to specification after 40 h. The organic reflux crossed its final operating value after 20 h. Figure 11c,d shows that the bottom and top HAc concentrations were not steady at the final simulation time of 100 h and that the bottom and top HAc purities were not all maintained at the designed values. Figure 11e,f shows that the reboiler duty was fixed at a constant value and the increase/decrease of the side-draw flow rate depended on the sixth tray temperature of the side draw. Figure 11g,h shows that the dynamic responses of the 12th stage temperature returned to specification after 15 h and the bottom PX flow rate returned to specification after 45 h. Figure 11i,j indicates that the water removed from this system met the specification and the stream flow rate could not cross its final operating value after 100 h. Figure 11k,l shows that the fourth stage temperature could return relatively stable specification after a period of time, but the reflux flow rate did not reach a new steady-state operating value at the final simulation time. 5.3. Summary of Dynamic Simulation Responses. The closed-loop dynamic responses for CS1 and CS2 under ±5% changes in the flow rates of feeds 1, 2, and 3 are shown in Figures 6−8 and Figures 9−11, respectively. As can be seen from the dynamic responses, CS1 and CS2 could prevent ±5% disturbances in the flow rates of feeds 1 and 2. However, they had limited ability to prevent ±5% disturbances in the flow rate of feed 3. The dynamics results indicated that the control structure (organic reflux flow rate controlling the sensitive tray temperature of the upper zone of C-101) could control the disturbance of upper-zone feed streams effectively, and CS2 was a suitable control strategy for this system. However, the control structure had limited ability to prevent unmeasured disturbances (disturbances of lower feed streams). A possible reason for this deficiency is that the sensitive eighth stage might have been too far from the lower feed tray. The organic reflux control scheme might be more suitable for the disturbances of the upper feed streams but might not achieve the desired effect for the disturbances of lower feed streams. By comparing the results of Figures 6 and 9, we found that the C-101 bottom product was maintained quite close to its specifications of 93 wt % HAc, the column top aqueous outlet stream was kept at 0.1 wt % acetic acid loss, and the water removed from this system met the specification under the two control structures (CS1, CS2). However, the organic reflux and removal of the accumulated PX (CS1) needed to take much more time to cross its final operating value. By comparing the results of Figure 7 with Figure 10, we found that the bottom product and column top aqueous outlet stream were both maintained quite close to their specifications under the two control structures (CS1, CS2). However, the organic reflux was not steady at the final simulation time of 100 h under a +5% change in the feed 2 flow rate (CS1). On the contrary, the organic reflux crossed its final operating value after 12 h for control structure CS2. CS2 removed the accumulated PX much faster, whereas CS1was not steady at the final simulation time of 100 h under a +5% change in the feed 2 flow rate and needed much more time under a −5% change in feed 2 flow rate to cross its final operating value. The stream flow rate could cross its final operating value under CS2: Control structure CS2 gave much more stable performance than CS1. The fourth stage temperature returned to relatively stable specification more regularly than with CS1. At the same time, the reflux flow rate was faster than CS1 in terms of the arrival of the new steadystate operating values. By comparing the results of Figures 8

6. CONCLUSIONS In the present study, multiple column feed streams, coming from various parts of the upstream process, were fed into this system. Apart from the main components HAc and H2O, small amounts of PX and MA were also include in the feed components. Optimum column designs and operating conditions were obtained for this system by rigorous process simulation. The overall control strategy CS1 and CS2 were proposed for this column system to achieve both bottom- and top-product specifications, remove accumulated PX, and recycle MA despite ±5% disturbances in the flow rates of feeds 1, 2, and 3. The dynamic results indicated that the control structure (organic reflux flow rate controlling the sensitive tray temperature of upper zone of C-101) could control disturbances of the upper-zone feed streams (feed 1 and 2) effectively. In addition, for the disturbances of the upper-zone feed streams, CS2 provided better control performance than CS1, which not only overcame the drawback of PX imbalance, but also kept overall solvent dehydration system much more stable. Moreover, CS2 can be easily implemented in industry for wider applications.

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AUTHOR INFORMATION

Corresponding Authors

*E-mail: [email protected] (H.W.). Tel.: +86-25-8358 7198. *E-mail: [email protected] (G.G.). Tel.: +86-25-8358 7198. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS

The authors gratefully acknowledge the National Key Technology R&D Program of China (No. 2011BAE05B03), a Project Funded by the Priority Academic Program Development of Jiangsu Higher Education Institutions (PAPD), and the Sinopec Yangzi Petrochemical Company Ltd.

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NOMENCLATURE C = column D = decanter HE = heat exchanger KC = proportional gain P = pump τI = integral time

Subscripts

i, j=component i, j 11342


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Dynamic Simulation and Control of a Complete Industrial Acetic Acid

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