Methanol

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Design and Control of Distillation System for Methylal/Methanol Separation. Part 2: Pressure Swing Distillation with Full Heat Integration Baoru Yu, Qiaoyi Wang, and Chunjian Xu* State Key Laboratory of Chemical Engineering, Tianjin University, Tianjin 300072, China ABSTRACT: A new method for methylal/methanol separation is presented by using fully heat-integrated pressure swing distillation. Rigorous steady state and dynamic simulations for this neat operation are implemented on commercial simulators (Aspen Plus and Aspen Dynamics). On the basis of the proposed partial optimization and global economical optimization, an optimized configuration for this fully heat-integrated pressure swing distillation is developed. From the simulation results, it is found that this process is more competitive than the one via extractive distillation from the economical view. Several control structures for this system are presented. The dynamic simulation results reveal that the proposed basic control structure is unable to maintain the two bottom products at their quality specification. This problem can be resolved by using a pressure-compensated temperature control scheme. The dynamic responses of this pressure-compensated temperature control show that this control structure works pretty well for this fully heat-integrated pressure swing distillation, even for large feed flow rate and composition disturbances. Contrast between the dynamic controllabilities of extractive distillation process and pressure swing distillation process is also made. Results show that the dynamic performances of the two alternative processes are somewhat similar.

1. INTRODUCTION As stated in our previous study, because of its exceptional properties, methylal has been found to be an important organic that has wide applications in many industries, such as cosmetics, medicine, polymer material, fuels, etc., so it is desirable to get high purity methylal. Traditionally, methylal is synthesized by reaction of methanol with formaldehyde or paraformaldehyde in the presence of a catalyst.1 Because of the high nonideality in the liquid phase, a minimum-boiling azeotrope is formed between methylal and methanol. The methylal/methanol mixture cannot be separated completely through a simple distillation process. Azeotrope separation has been extensively studied in the open literature. The two most common methods for separating such a mixture are azeotropic distillation and extractive distillation (ED). In our previous study, extractive distillation using dimethylformamide (DMF) as entrainer was put forward for methylal/methanol separation. Despite the fact that the pressure sensitivity of azeotropes has been long known, pressure swing distillation (PSD) has not been broadly used in the chemical industry.2 M€uller3 suggested separation of methanol/methylal mixtures by pressure swing distillation. High-purity methylal can be withdrawn from the bottom of the high-pressure column, which is operated under a pressure of 930 bar. A number of mixtures can form an azeotrope whose position changes substantially with the system pressure (at some pressures, the azeotrope may disappear). Lewis4 appears to be the first to suggest distilling the azeotropic mixtures by pressure swing distillation. PSD uses the pressure sensitivity of the azeotrope point to break the restriction of the distillation boundary; thus, effective separation is achieved. Provided the azeotrope is pressure-sensitive or one end of the distillation boundary is a pressure-sensitive azeotrope when an entrainer is added to the system,2 pressure swing distillation can be used. r 2011 American Chemical Society

Wasylkiewic5 developed a new algorithm for pressure sensitivity analysis of the azeotrope using bifurcation theory and an arc length continuation. The azeotrope compositions can be tracked as they change with pressure, and all new azeotropes that appear within a specified pressure range can be found. Consider the case of a minimum-boiling homogeneous binary azeotrope for the mixture AB, with Txy phase diagram as shown in Figure 1A. As the pressure is increased from P1 to P2, the azeotropic composition moves toward a smaller fraction of A. Figure 1B gives the conventional PSD process flowsheet for this mixture AB. The fresh feed, F, and the recycled stream, D2 (distillate of second column), are fed to the first column, which operates at lower pressure, P1. High-purity A is removed as the bottom product, B1, while a distillate, D1, whose composition is near azeotropic composition at pressure P1, is removed from the top of the column. Stream D1 is changed to pressure P2 and fed to the second column, which operates at pressure P2. High-purity B is removed as the bottom product B2, while a mixture near the azeotropic composition at pressure P2 is the distillate, D2. Stream D2 is then recycled to the first column through a pressure release valve. Since PSD operates at different pressures, it is possible to reduce the energy consumption substantially by heat integration. In addition, there is no addition of entrainer to the system; thus, air pollution caused by the entrainer can be avoided.6 With rising fuel prices, more attention is turning to this energy-saving process, and this technology is enjoying growing popularity in the chemical industry. In recent years, extensive studies712 have Received: August 30, 2011 Accepted: December 20, 2011 Revised: November 20, 2011 Published: December 20, 2011 1293

dx.doi.org/10.1021/ie201949q | Ind. Eng. Chem. Res. 2012, 51, 1293–1310

Industrial & Engineering Chemistry Research

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Figure 1. PSD process (for minimum-boiling homogeneous binary azeotrope AB) (A) Txy curve at pressure P1 and P2. (B) Distillation sequence.

been made on the pressure swing distillation process, including heat integration. In the past three decades, great efforts have made to improve the heat efficiency of the distillation process. Heat integration may be the most widely studied. Two kinds of heat integration exist for a PSD process: one is between the condenser in the highpressure distillation column and the reboiler in the low-pressure distillation column (termed the condenser/reboiler type heat integration), and the other is between the rectifying section in the high-pressure distillation column and the stripping section in the low-pressure distillation column (termed as the rectifying/ stripping section type heat integration).11 Both have the potential of large energy savings for the separation of close-boiling mixtures.13 For condenser/reboiler type heat integration, a partial heatintegrated distillation column (heat integration with auxiliary reboiler or condenser) has been extensively studied in the open literature, both on steady state design and dynamic performance.1417 However, there is limited research on full heat integration. Generally speaking, despite the fact that full heat integration can lead to a significant cost savings, it introduces a control problem. One control degree of freedom is lost, since the heat input to the reboiler of the low-pressure column cannot be independently set. Luyben18 developed a control system for fully heat-integrated pressure swing azeotropic distillation for separating THF/water. The results demonstrate that the fully heatintegrated system can handle only a fairly small disturbance without having purity specifications violated. Because of the economical advantage and dynamic controllability penalty associated with this process, it is desirable to solve its control problem to make the most of its economical advantage. This is a continued work on the methylal/methanol separation process. In this work, on the basis of the vaporliquid equilibrium information, we present an optimized configuration for the separation of methanol/methylal by using fully heat-integrated PSD. Both steady state design and dynamic control of the system are considered here. Comparisons between the extractive distillation process and pressure swing distillation process are also made for this system. The comparisons will be in terms of both steady state design and dynamic controllability.

2. STEADY STATE DESIGN The PSD process considered in this study is simulated with the same basic data as the previous study using extractive distillation. The feed is a mixture made up of 85.8 wt % methylal, 13.9 wt % methanol, and 0.3 wt % water, with a mass flow rate of 3000 kg/h. The steady state and dynamic simulations are implemented with the help of commercial software (Aspen Plus and Aspen Dynamics). The NRTL activity model was chosen as the property package in the simulation using the built-in binary interaction parameters in the simulator. Useful skills to use the two simulators are covered in detail in Lyuben’s book.19 The product specifications are set to be exactly the same as in the previous work: the methylal impurity in the methanol product is not more than 0.2 wt %, and the methylal product should have a purity of 99.9 wt %. 2.1. VaporLiquid Equilibrium. The normal boiling points of methylal and methanol are 41.94 and 64.53 °C, respectively. Owing to the high nonideality of the liquid phase, a minimumboiling azeotrope is formed. At atmospheric pressure, an azeotrope with composition of 94.06 wt % methylal is formed at 41.37 °C. Since the two columns operate at different pressures, it is a convention to set the operating pressure of the low-pressure column within such a range that the recycled cooling water with temperature around 32 °C can be used as the cooling media of the condenser. Thus, the operating pressure of the low-pressure column (LPC) is set at atmospheric pressure (101.3 kPa). For the high-pressure column (HPC), the operating pressure should ensure 5% or more change in the azeotrope composition over a moderate range of pressure.20 The effect of pressure on the azeotropic composition and temperature for the methyl/methanol binary azeotrope is shown in Figure 2. The shifts in the azeotropic point with pressure are clearly seen from Figure 2. It indicates that the high-pressure column with an operating pressure within 3001500 kPa seems to be reasonable, and a further increase in the system pressure results in quite a small change in the azeotropic composition. The operating pressure of the high-pressure column will be optimized from an economical view in a later section of this paper. Figure 3 gives the Txy curves for a methyl/methanol binary mixture at 101.3 and 1000 kPa. 1294

dx.doi.org/10.1021/ie201949q |Ind. Eng. Chem. Res. 2012, 51, 1293–1310

Industrial & Engineering Chemistry Research 2.2. Preliminary Simulation. To begin the steady-state simulation, the theoretical trays of the low- and high-pressure columns are selected with 20 trays and 30 trays (including the condensers and reboilers), respectively. The operating pressure of the HPC is 1000 kPa. Initial estimate values for the reflux ratio of the columns are 0.4 for the LPC and 1.5 for the HPC, feed stages of both columns are exactly at the middle of the column, and the recycled stream is fed to stage 15 of the LPC. Bottom rate and reflux ratio are manipulated to meet the top and bottom product specification by using the “Design spec/Vary” function in Aspen Plus. All the pumps with efficiency of 0.7 are assumed. We use the Aspen notation of numbering stages from the top, with stage 1 being the reflux drum and the last stage being the reboiler. After running the simulation, the reboiler duties are 711.05 kW for the LPC and 908.32 kW for the HPC, respectively, and the condenser duties are 830.17 kW for the LPC and 643.35 kW for the HPC, respectively. Note here that the pump power is negligible, compared with the heat duty of the reboilers. This is useful information for partial optimization, which will be discussed below. The large temperature difference between the LPC reboiler and HPC condenser indicates the possibility of heat integration between the two columns, which could reduce fixed capital investment and operating costs by a large margin.

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2.3. Optimization. 2.3.1. Partial Optimization Based on the HPC Reboiler Heat Duty As a Reference Variable. To select the best

conditions for this process, we start with a partial optimization. In the partial optimization stage, only the operating costs are taken into consideration with both column tray numbers and the HPC operating pressure fixed. Since heat integration is possible between the two columns, full heat integration is considered here so that neat operation can be achieved without an auxiliary Table 1. Case Studies for Pressure Swing Distillation Process with HPC at 1000 kPa parameters

case 1 case 2 case 3 case 4 case 5 case 6

NT1

20

20

20

18

16

14

NT2

32

30

28

30

30

30

NF1 and NR (optimum)

6/15

6/15

6/15

6/13

5/12

5/10

NF2 (optimum) RR1

19 0.301

17 0.304

15 0.308

17 0.306

17 0.304

17 0.312

RR2 (optimum)

1.46

1.46

1.46

1.46

1.44

1.45

ID1 (m)

0.86

0.86

0.87

0.86

0.87

0.87

ID2 (m)

0.79

0.79

0.79

0.79

0.79

0.80

QR1 (kW)

670.08 673.38 677.55 674.78 677.13 684.07

QR2 (kW)

942.39 946.33 951.34 947.98 951.94 959.59

TAC ($1000 per annum) 572.84 572.04 572.43 570.35 569.91 571.57

Table 2. Case Studies for Pressure Swing Distillation Process with HPC at 1200 kPa parameters

Figure 2. Effect of pressure on azeotropic composition and temperature.


NT1

20

20

20

18

16

14

NT2

30

28

26

28

28

28


5/16

5/16

5/15

5/14

5/12

4/10

NF2 (optimum) RR1

17 0.219

15 0.223

14 0.234

15 0.224

15 0.226

15 0.225

RR2 (optimum)

1.34

1.34

1.37

1.34

1.34

1.32

ID1 (m)

0.83

0.83

0.83

0.83

0.83

0.83

ID2 (m)

0.78

0.78

0.78

0.78

0.78

0.78

QR1 (kW)

598.90 602.37 608.18 603.22 605.38 608.83

QR2 (kW)

896.19 900.53 905.88 901.57 904.21 909.73

TAC ($1000 per annum) 557.17 557.03 557.66 555.18 554.16 554.77

Figure 3. Txy diagrams for a methylal/methanol mixture at 101.3 and 1000 kPa. 1295



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reboiler or condenser. In view of the price of the cooling water being much cheaper than the steam as well as the pump power being negligible compared with the heat duty of the reboilers, Table 3. Case Studies for Pressure Swing Distillation Process with HPC at 1400 kPa parameters


NT1

20

20

20

18

16

14

NT2

28

26

24

26

26

26


5/15

5/15

5/16

4/14

4/12

4/10

NF2 (optimum)

16

14

12

14

14

14

RR1 RR2 (optimum)

0.173 1.31

0.174 1.30

0.174 1.28

0.168 1.27

0.172 1.28

0.178 1.29

ID1 (m)

0.80

0.80

0.81

0.80

0.8

0.81

ID2 (m)

0.78

0.78

0.78

0.78

0.78

0.78

QR1 (kW)

550.11 552.99 556.43 552.04 553.96 557.98

QR2 (kW)

866.96 871.39 877.28 872.24 873.95 878.31

TAC ($1000 per annum) 566.10 565.90 566.56 563.98 562.59 562.79

here, partial optimization using the HPC reboiler duty QR2 as a reference variable is carried out. Design variables and optimization variables must be specified in partial optimization. In this work, we specify the temperature, pressure, flow rate, and composition of feed. In addition, quality requirements of the two bottom products are also selected as the design variables. In each optimization, the number of theoretical plates and the operating pressure are fixed. In contrast, the reflux ratio of columns, feed streams, and recycled stream feeding locations of both columns are optimization variables. These optimization variables are optimized to minimize the HPC reboiler duty QR2. Before implement of rigorous distillation column simulation, we have determined the minimum number of stages (Nmin) and minimum reflux ratio (RRmin), assuming the recycle stream with a composition exactly its azeotropic composition at P2. The corresponding recycle stream flow rate is determined by simple material balance calculation. A heuristic that set the reflux ratio (RR) to 1.2 times minimum reflux ratio (RRmin) is used to determine the number of stages. This shortcut distillation design is implemented using the “DSTWU” block in Aspen Plus.

Figure 4. Flowsheet for the optimized pressure swing distillation process with full heat integration.

Figure 5. (A) Low-pressure column temperature profile. (B) High-pressure column temperature profile. 1296



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Figure 6. (A) Low-pressure column liquid composition profile. (B) High-pressure column liquid composition profile.

Figure 7. Aspen Dynamics flowsheet equations for full heat integration.

Then the “RadFrac” block is used to conduct rigorous distillation simulation. The “Design Spec” and “Vary” feature in each column block in Aspen Plus is used to adjust the bottom flow rate to meet the bottom product quality requirements. To achieve full heat integration between the two columns (that is, heat remove rate in the HPC condenser QC2 is exactly equal to the heat input rate in the LPC reboiler QR1), the “flowsheet design spec” function is used to make QR1 equal to QC2, so the reflux ratio of the two column cannot be set independently. Here, we adjust the LPC reflux ratio RR1 to make QR1 equal to QC2, while the feeding locations of both column and the HPC reflux ratio (RR2) are varied to minimized the HPC reboiler duty QR2. A sequential iterative optimization search is used to find the optimal design, with the HPC reflux ratio RR2 as the outer iterative loop and feed streams and recycled stream feeding locations as the inner iterative loop. The partial optimization is guided by the following optimization procedures: (1) Fix the number of theoretical plates (the value calculated from the shortcut distillation design as initial estimate) and the operating pressure for the both columns; purity requirements of the two bottom products are also specified. (2) Set the reflux ratio of the two columns using the value calculated from the shortcut distillation design as initial estimate. (3) Use the “Design spec/vary” function to adjust the bottom product flow rate until the two bottom product quality specifications can be met. (4) Use the “flowsheet design spec” function to adjustthe LPC reflux ratio RR1 to make QR1 equal to QC2.

(5) Vary the reflux ratio of the HPC RR2 until the HPC reboiler duty, QR2, is minimized. (6) Change the feed streams and recycled stream feeding locations of both columns until the HPC reboiler duty QR2 is minimized. (7) If the HPC reboiler duty, QR2, is minimized, partial optimization stops. If not, go back to step 5 and continue optimization. The partial optimization is carried out for several cases. In each case, the operating pressure and the number of theoretical plates are fixed. The aim of partial optimization is to facilitate the global economical optimization, by which a process with optimized operating conditions is gained. 2.3.2. Global Economic Optimization. Using the same objective function as our previous study, a global economic optimization is carried out on the basis of the total annual costs (TAC). TAC ð103 $=yearÞ ¼ Cv þ 0:3 3 FCI

ð1Þ

The major pieces of equipment in this process are the two distillation column vessels (including column internals), reboilers, and condensers. Small items such as reflux drums, pumps, valves, and pipes are usually not considered because their costs are much lower compared with the costs of the column vessels and heat exchangers. The “tray sizing” function employed to sizing the column vessel and sieve plate is selected. Since the two columns have small diameters, we specify the tray spacing is 0.4 m instead of the default value (0.61 m) in the simulator. Here, we assumed that all the column vessels (including the sieve plate) and heat 1297



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Figure 8. Basic control structure for pressure-swing distillation system with full heat integration.

Figure 9. Relayfeedback test results for the temperature controllers in the basic control structure: (A) relayfeedback test result for TC1 and (B) relayfeedback test result for TC2.

exchangers are carefully made of stainless steel. The heat transfer areas for the condensers and reboilers are determined by using the overall heat transfer coefficient and a differential temperature driving force. Here, the assumed overall heat transfer coefficients are 0.852 kW/(K 3 m2) and 0.568 kW/ (K 3 m2) for the condenser and reboiler, respectively, which are

taken from Luyben’s book.12 All the major equipment costs are estimated using the cost estimation program CAPCOST of Turton.21 As for the utility costs, they can be calculated from the heat duties of the reboilers and condensers as well as the power of all the pumps. Suppose that 5 barg low-pressure steam (159 °C) and 1298



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30 °C cooling water are available in the plant. The utility prices are also taken from CAPCOST. It is known that for the pressure swing distillation process, the larger the difference in the operating pressures of the two columns, the lower the recycle flow rate and energy consumption that are required; however, the HPC operating pressure is limited by the available utilities. The higher the operating pressure of the HPC, the lower the temperature difference between the steam and the HPC reboiler; thus, a larger heat transfer area is required, which results in an increase in the fixed capital investment. Strictly speaking, to determine the operating pressure for the Table 4. Temperature Controllers Tuning Parameters for the Basic Control Structure parameters

TC1

TC2

controlled variable

T1,14

T2,23

manipulated variable

RR1

Q R2

transmitter range

280380 K

350450 K

controller output range

00.452

06.510 GJ/h

ultimate gain ultimate period

7.983 12.45 min

6.769 4.56 min

gain Kc

2.495

2.115

integral time τI

27.39 min

10.03 min

HPC, several alternatives with different operating pressures should be compared from an economical point of view. Here, HPC operating pressures at 1000, 1200, and 1400 kPa are studied to make comparisons. To find the best operating condition for this process, five cases for each pressure were studied. Differences between each of the cases are the number of theoretical plates and the operating pressure of the HPC. In each case, the abovementioned partial optimization was carried out to minimize the HPC reboiler heat duty QR2 to find the optimal feeding locations and HPC reflux ratio RR2. Tables 1, 2, and 3 give the details of each case and the TACs estimated from CAPCOST are also listed. As can be seen from these tables, the optimum for this pressure swing distillation process corresponds to case 5, with HPC at 1200 kPa. An increase or decrease in the HPC operating pressure may result in an increase in the TAC. Although increasing HPC operating pressure will reduce the energy consumption and column diameter of the HPC, a bulkier reboiler for the HPC is expected because of the reduction of the temperature difference between the HPC base and the heating steam. This illustrates the trade-off between the fixed capital costs and the operating costs. On the basis of global economic optimization, we have good reasons to believe case 5 with HPC operated at 1200 kPa might work quite well in practical industrial production.

Figure 10. Dynamic responses for the basic control structure: 20% feed flow rate disturbances.

Figure 11. Dynamic responses for the basic control structure: feed composition disturbances. 1299



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Figure 12. Improved control structure with QR2/F ratio.

Figure 13. Dynamic responses for the improved control structure: 20% feed flow rate disturbances.

It is noticed that the effect of the number of theoretical plates is surprisingly smaller compared with HPC operating pressure. This reveals that selecting an appropriate operating pressure is of vital importance for this process. Figure 4 gives the fully heat-integrated flowsheet with detailed steam information, heat duties, equipment sizes, and operating conditions at the steady state design conditions. This corresponds to case 5 with HPC at 1200 kPa. Figures 5 and 6 give temperature and liquid composition profiles in the

two columns. As can be seen from Figures 5 and 6, there is a rapid rise in the temperature near the bottom of both columns as methanol builds up near the LPC bottom and methylal builds up near the HPC bottom. It is obvious that stage 14 displays a fairly steep slope for LPC, as does stage 23 for HPC. These indicate the proper temperature control point for the two columns. Comparing the pressure swing distillation process with the extractive distillation process which has been extensively studied 1300



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Figure 14. Dynamic responses for the improved control structure: feed composition disturbances.

in part 1 of this paper,24 it can be found that the extractive distillation column is much taller and the solvent recovery column is much thinner. Detailed comparison between these two techniques should be made from an economical view using the TAC as the objective function. The total annual costs of the overall process via pressure swing distillation is $5.542 105 per annum (see Table 2), which can be directly compared with the TAC via extractive distillation ($6.154 105 per annum). This shows that separation via fully heat-integrated pressure swing distillation cuts the TAC by as much as 9.94%. Although much heat is integrated (605 kW) for this fully heatintegrated pressure swing distillation, the TAC does not reduce substantially, as expected. This is because in the extractive distillation process, the column is operated at atmospheric pressure, and entrainer is introduced to the system, which facilitates the separation. In addition, in the pressure swing distillation process, one column is operating at high pressure, which means that a thicker vessel wall is required for the HPC vessel and heat exchangers. If heat integration is not introduced to the process, with separate reboilers and condensers for the columns, the total annual cost (TAC) is $7.951 105 per annum.

3. CONTROL SYSTEM DESIGN 3.1. Basic Control Structure for Pressure Swing Distillation System with Full Heat Integration. The above-

mentioned fully heat-integrated pressure swing distillation system is highly intercoupled and interacted through the two distillate steams connecting them and heat transfer through the combined condenser/reboiler. A disturbance occurring in one column can be directly transmitted to the other column, and this transmission will inevitably bring about a control problem. Because of the economical advantage of the steady state design of case 5 with HPC operated at 1200 kPa, this case will be selected for the control system design study. The plumbling system and major equipment sizes are specified according to the criteria mentioned in part 124 of this paper. Then the steady state flowsheet is pressure checked, and the Aspen Plus file is exported to Aspen Dynamics. Since the two columns are fully heat-integrated, the combined condenser/reboiler heat duty is the product of the overall heat

Figure 15. Effect of system pressure on bubble point temperature for liquid mixture for stage 23 of the HPC.

transfer coefficient, heat transfer area, and the temperature difference between the reflux drum of the HPC and the base of the LPC. The results of the steady state design and assumed overall heat transfer coefficient gives a heat transfer area of 18.03 m2. To achieve full heat integration in Aspen Dynamics, the “flowsheet equations” function is employed. As shown in Figure 7, the appropriate equations are entered in the text editor window. The heat removal rate of the HPC is calculated by the first equation, and the equality of the heat input rate of the LPC reboiler and heat removal rate of the HPC are set by the second equation. After compiling the flowsheet equations, the simulation is overspecified by two variables, which makes the simulation unable to run. To cope with this problem, the heat duty in the condenser of the HPC and the heat duty in the reboiler of the LPC must be changed from “fixed” to “free”. The following control structure is proposed for this fully heatintegrated pressure swing distillation system, and its effectiveness will be evaluated later. Figure 8 gives the basic control structure for this pressure swing distillation system with full heat integration. (1) Feed is flow-controlled (reverse acting). (2) Reflux drum levels in both columns are held by manipulating distillates flow (direct acting). 1301



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Figure 16. Flowsheet equations for fully heat-integrated pressure swing distillation by using pressure-compensated temperature control.

Figure 17. Control structure for fully heat-integrated pressure swing distillation by using pressure-compensated temperature control.

(3) Base levels in both columns are held by manipulating flow bottom (direct acting). (4) The pressure in the low-pressure column is controlled by manipulating the heat removal rate in the condenser of the low-pressure column (reverse acting). (5) The reflux ratio in the high-pressure column is fixed. (6) The temperature for stage 23 in the high-pressure column is controlled by manipulating the heat input rate in the reboiler of the high-pressure column (reverse acting). (7) The temperature for stage 14 in the low-pressure column is controlled by manipulating the reflux ratio in the lowpressure column (direct acting). A noteworthy feature revealed here is that there is no pressure controller in the high-pressure column, which is also described by

Luyben.17 It is not controlled, but floats with the operating conditions. This is because if the feed flow rate increases, more heat is needed to the condenser/reboiler to maintain the purity of the bottom product of the LPC, and this is implemented by increasing the operating pressure of the HPC to get a larger temperature difference, then the HPC temperatures drop as a result of more cold feed being charged to this column, which increases the reboiler duty; thus, more vapor rises and condenses in the condenser/reboiler to ensure the LPC reboiler heat input. On the other hand, if the feed flow rate decreases, the operating pressure of the HPC should be lowered to decrease the heat transfer rate. In this dynamic simulation, a “PC2” block is added in the flowsheet and put on manual as a HPC top pressure indicator. 1302



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Figure 18. Dynamic responses for the pressure-compensated temperature control structure: 20% feed flow rate disturbances.

Figure 19. Dynamic responses for the pressure-compensated temperature control structure: feed composition disturbances.

Here, it is worth noting that there are some differences between the above-mentioned control structure and the control structure developed by Luyben.18 In Luyben’s control system, the temperature of the low-pressure column is controlled by manipulating the reboiler heat input of the high-pressure column, whereas the temperature of high-pressure column is controlled by manipulating the reflux ratio of the high-pressure column. Although the HPC reboiler heat input has a rapid response in the presence of the feed flow rate and composition disturbances for Luyben’s control structure, using a control structure as we have proposed may reflect the fluctuations in the HPC more effectively. Conventional PI controllers are used for all controllers except the four liquid level controllers. Proportional controllers are used for all liquid levels with Kc = 2. Two deadtime elements are inserted in the two temperature control loops with deadtime of 1 min. Relay-feedback tests are run on the two temperature controllers to determine ultimate gains and periods, and Tyreus Luyben tuning is used in both controllers. Obviously, the two-temperature control loop is highly interactive. On one hand, an increase in the HPC temperature controller output will reduce the heat input to the HPC and, in turn, reduce the HPC condenser heat duty or heat input to the reboiler of LPC; this will reduce the LPC temperature.

Table 5. Temperature Controllers Tuning Parameters for the Control Structure with RR2 Adjusted by TPC parameters

TC1

TC2

TC3

controlled variable

T1,14

TPC

TPC

manipulated variable

RR1

QR2

RR2

transmitter range

280380 K

350450 K

350450 K

controller output range ultimate gain

00.452 7.983

06.510 GJ/h 6.769

02.68 19.794

ultimate period

12.45 min

4.56 min

7.44 min

gain Kc

2.495

2.115

6.186

integral time τI

27.39 min

10.03 min

16.37 min

In contrast, reduction in the heat input to the HPC will produce less distillate recycled to the LPC, which causes the LPC temperatures to rise. The LPC temperature fluctuates under the combination of these two opposite effects. On the other hand, a rise in the LPC temperature controller output will increase the reflux ratio in the LPC, and thus, the LPC distillate flow rate reduces, which will reduce the HPC temperatures. A good way to take this interaction into account is to first tune TC2 with TC1 on manual. Then TC2 is placed on automatic, and a relayfeedback test is run on TC1. This sequential tuning 1303



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Figure 20. Improved control structure with RR2 adjusted by the pressure-compensated temperature.

method is also recommended by Luyben.19 Figure 9 shows the relayfeedback test results for the basic control structure using the tuning method mentioned above. Table 4 gives controller parameters for the two temperature controllers. Now the dynamic performance of this basic control structure for fully heat-integrated pressure swing distillation is evaluated by the identical feed flow rate and composition disturbances as used in the first part of this paper. Figure 10 shows the dynamic responses of this control structure to positive and negative 20% step changes in the feed flow rate at t = 0.2 h. Figure 11 gives the dynamic responses for three feed composition disturbances at t = 0.2 h. As can be seen from Figure 11, there is a large transient deviation in the composition of the methanol product when a +20% step change in the feed flow rate takes place at t = 0.2 h, and the mass impurity of methylal rises to 1.29% at t = 0.33 h. This is because an increase in the feed flow rate causes a temperature drop in the column stripping section, and then a reduction of the reflux ratio. This results in more methylal escape from the bottom. An improved control scheme is using a feedforward ratio control structure, which will be discussed in the subsequent section. 3.2. Improved Control Structure with QR2/F Ratio. Adding a ratio control can improve the dynamic performance. The heat input rate to the high-pressure column reboiler QR2 is proportional to the feed flow rate. A multiplier block is added to the basic control structure. From the results of steady state simulation, the calculated ratio (second input to the multiplier) is

(3.25514 GJ/h)/(47.3397 kmol/h) = 0.06876 GJ/kmol (in metric units). The ratio is not fixed but adjusted by the HPC temperature controller TC2. Figure 12 shows the improved control structure with the QR2/F ratio. Relay-feedback tests are run on the two temperature controllers to determine ultimate gains and periods, and TyreusLuyben tuning is used. It can be found that the two temperature controllers' tuning parameters are almost not changed for this improved control structure with the QR2/F ratio. Figures 13 and 14 demonstrate the effectiveness of this improved control structure. As can be seen from Figure 13, the largest transient deviation reduced from 1.29 to 0.97% for the +20% feed flow rate disturbance when a QR2/F ratio control scheme is used. This is because an increase in the feed flow rate causes a rapid increase in the QR2; thus, more vapor rises from the bottom, which increases the heat removal rate in the HPC condenser (that is to say, more heat input to the LPC reboiler) and then prevents the more volatile component, methylal, escape from the bottom. The effect of this improvement is not significant owing to the hydraulics lag in the HPC. A certain time is required for the vapor to rise through trays from the bottom to the top in the HPC. Although the methylal impurity level in the methanol product is brought back almost to its set point (0.2 wt %) for a rather large feed flow rate and composition disturbances, the methylal product cannot be maintained at the 99.9 wt % specification at the new steady state conditions, especially for feed flow rate 1304



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Figure 21. Dynamic responses for the improved control structure with RR2 adjusted by TPC: 20% feed flow rate disturbances. 1305



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Figure 22. Dynamic responses for the improved control strcture with RR2 adjusted by TPC: feed composition disturbances. 1306



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Figure 23. Comparison of dynamic performance between PSD and ED for feed flow rate disturbances.

disturbances (a +20% step change in the feed flow rate brings the methylal purity to 99.45 wt % at the new steady state, and 99.995 wt % for 20% step change in the feed flow rate). This is because the pressure of the high-pressure column is not controlled but floats with the operating conditions. If the feed flow rate increases, more heat is needed in the condenser/reboiler, and this will produce a higher operating pressure of the HPC so that a larger temperature difference is achieved to ensure the LPC reboiler heat input. If the set point of the HPC temperature controller is fixed, a lower methylal concentration for stage 23 is expected, and then more methanol escapes from the bottom, so fixing the sensitive plate temperature is not feasible for this fully heat-integrated pressure swing distillation process. A pressurecompensated temperature control structure is required to eliminate this problem. 3.3. Pressure-Compensated Temperature Control. A pressure-compensated temperature control scheme should be able to bring the methylal product to its quality specification under the new steady state conditions. Pressure-compensated temperature has been briefly described in Buckley’s book,22 and implementation of pressure-compensated temperature control has been illustrated in detail in Luyben’s papers,18,23 but only a pressurecompensated temperature control structure for the partial heatintegrated pressure swing distillation process is covered without a pressure-compensated temperature control structure for the fully heat-integrated pressure swing distillation process. Luyben18 developed a control system for fully heat-integrated pressure swing azeotropic distillation for separating THF/water, and the results demonstrate that the fully heat-integrated system can handle only a fairly small disturbance without having purity

specifications violated. Here, a pressure-compensated temperature control scheme for pressure swing distillation with full heat integration is set up. 3.3.1. Setting up Pressure-Compensated Temperature Control Scheme. Before setting up the pressure-compensated temperature control scheme, vaporliquid equilibrium calculations are required. From the steady state simulation results, we find the liquid phase composition for stage 23 in the high-pressure column is 97.77/2.09/0.14 wt % methylal/methanol/water. Since the HPC is operated at 1200 kPa at the design conditions and the top pressure is not controlled to ensure the full heat integration, the HPC top pressure was assumed to fluctuate within the range of 11001300 kPa for various disturbances. Bubble point temperatures for this liquid are calculated within this pressure range. Figure 15 plots the bubble point temperature as a function of pressure for this mixture. It is observed that there is an almost linear relationship between the bubble point temperature and pressure within this range, with a slope of 4.026. Since the temperature controller of the high-pressure column is made as reverse-acting, the following equation is used to calculate the pressure-compensated temperature. TPC ¼ T2, 23 4:026 ðP 12Þ

ð2Þ

The temperature for stage 23 and the top pressure in the highpressure column are measured. Using the above equation, the pressure-compensated temperature is acquired, which is fed to the corresponding deadtime element in the flowsheet, and then sent to the temperature controller, TC2 (see Figure 17). 1307



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Figure 24. Comparison of dynamic performance between PSD and ED for feed composition disturbances.

Implementation of this pressure-compensated temperature control is via the “Flowsheet Equations” function in Aspen Dynamics.14,18 As is shown in Figure 16, the third equation is used to provide this pressure-compensated temperature control. 3.3.2. Preliminary Pressurecompensated Temperature Control Structure for Fully Heat-Integrated Pressure Swing Distillation. Figure 17 shows the control structure for fully heat-integrated pressure swing distillation by using pressure compensated temperature control. The effectiveness of this control structure is demonstrated in Figures 18 and 19. The same feed flow rate and composition disturbances are made to verify its effectiveness. From Figures 18 and 19, it can be found that this pressurecompensated temperature control scheme can handle both feed flow rate and composition disturbances well. The purity of the

methylal product returns to a point that is in close proximity to its quality specification (99.9 wt % methylal). 3.3.3. Improved Control Structure with RR2 Adjusted by the Pressure-Compensated Temperature. Notice that it takes much more time to come to a new steady state for the 20% feed flow rate disturbance. This can be improved with an improved control structure, in which the reflux ratio of the high-pressure column RR2 is not fixed, but varied with the pressure-compensated temperature. Figure 20 shows this improved control structure with RR2 adjusted by the pressure-compensated temperature. A “TC3” block is added in the flowsheet to implement this. A relayfeedback test is run on temperature controller TC3 to determine the ultimate gain and period, and TyreusLuyben tuning is used to obtain the controller gain and integral time. Table 5 gives the three temperature controller tuning parameters for this improved 1308


Industrial & Engineering Chemistry Research control configure. Figures 21 and 22 demonstrate the dynamic controllability of this improved control structure. The same feed flow rate and composition disturbances are made to verify its effectiveness. As can be seen from Figures 21 and 22, this improved control structure with RR2 adjusted by the pressure-compensated temperature (TPC) can handle large feed flow rate and composition disturbances quite well. The two bottom products return to their desired values. Notice that this control structure provides a good dynamic control for a 20% feed flow rate disturbance; it takes much less time to come to a new steady state. This is because the HPC reflux ratio, RR2, is not held constant but varies with TPC. With a decrease in the feed flow rate, less heat is input to the HPC reboiler immediately, which causes a rapid decrease in the stage 23 temperature, and the reflux ratio in the HPC decreases, which causes the HPC condenser duty to decrease faster compared with the control system with fixed RR2. As the HPC condenser duty decreases, less heat is input to the LPC reboiler. This provides a faster signal transmission scheme between the two columns, so much less time is required to come to a new steady state. The temperature for stage 23 in the HPC is not held at its set point (406.2 K), but increases to 412.2 K for a +20% step change in the feed flow rate and decreases to 398.8 K for a 20% step change in the feed flow rate. It should be noticed that the maximum deviation reduced to 0.66 wt % instead of 0.97 wt %. This means that the pressure-compensated temperature control scheme also can improve the methanol product quality to some extent. Since the pressure of the HPC is not controlled but floats with operating conditions to provide an appropriate driving force to meet the heat duty requirement for the combined condenser/ reboiler, if a +20% feed flow rate disturbance is imposed on the system, the top pressure in the HPC increases to ∼1347 kPa at the new steady state and decreases to 1015 kPa for a 20% feed flow rate disturbance. Large pressure changes are also observed for feed composition disturbances. It is interesting to find that there is a larger pressure change for a feed composition change to 90.8/8.9/0.3 wt % methylal/methanol/water. This phenomenon can be explained as when a larger deviation of the LPC bottom flow rate occurs for this feed composition disturbance, a larger pressure change is required to meet the heat transfer requirement. A noteworthy feature revealed here is that nonmonotonicities are presented for all the investigated variables. It is difficult to explain this phenomenon because this fully heat-integrated system is highly intercoupled and interacted through the two distillate streams connecting them and the condenser/reboiler heat integration. A disturbance occurring in one column can be easily transmitted to the other column. 3.3.4. Comparison of Dynamic Performance of Fully Heat Integrated PSD Process and ED Process. Now the dynamic performance of the two processes for methylal/methanol separation is compared. Figures 23 and 24 give a direct comparison of the two alternatives for the same feed flow rate and feed composition disturbances. The solid lines are for the pressure swing distillation process with the improved pressure-compensated temperature control structure, and the dashed lines are for the extractive distillation process with an R/F ratio control structure. As is shown in Figure 23, larger transient deviations occur for the extractive distillation process when a +20% feed flow rate disturbance is applied. In contrast, there are larger transient deviations for the pressure swing distillation process for a 20% feed flow rate disturbance. Despite this fact, both processes return to

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a new steady state in 3 h, with both products at acceptable levels of quality. Figure 24 depicts the closed-loop responses for the two alternatives for feed composition disturbances. Although the pressure swing distillation exhibits large transient deviations in the methanol product, the pressure swing distillation process shows better controllability in maintaining the methylal product purity. Note that the extractive distillation has a worse dynamic controllability for 2 wt % water disturbance, compared with its counterpart, pressure swing distillation. For the extractive distillation process, it takes more time to reach a new steady state and the methylal product with a composition of 99.7 wt % at the new steady state. A recommended method presented in our previous study is a pretreatment to remove some water in the raw material. It can be seen from Figures 23 and 24 that both processes performed nicely in maintaining the two products' quality, and the dynamic controllability of the two alternative processes is somewhat similar. From the standpoint of dynamic control stability, both processes can be used for methylal/methanol separation.

4. CONCLUSIONS A new method for separating methylal/methanol is developed by using PSD. Full heat integration between the columns is considered here to save costs. Steady state simulation is carried out with the help of the commercial simulator Aspen Plus. Using the proposed optimization method, an optimized configuration is presented. From the steady state simulation results, it was found that a HPC working at 1200 kPa is reasonable. Lowering or raising the pressure may result in an increase in total annual costs. This optimal design flowsheet is compared directly with an earlier study via extractive distillation. It is found that the fully heat-integrated pressure swing distillation process has lower costs for its intrinsic energy saving characteristic. Several control structures for this fully heat-integrated pressure swing distillation system are explored, and dynamic simulations are realized using commercial software Aspen Dynamics. The proposed basic control structure with a fixed sensitive plate temperature cannot maintain the quality specification for large disturbances. A pressure-compensated temperature control structure is discussed to maintain methylal product quality specification. Results show that a pressure-compensated temperature control structure provides good dynamic control for this fully heat-integrated pressure swing distillation. A dynamic controllabilities comparison between pressure swing distillation and extractive distillation is also demonstrated. Somewhat similar dynamic performances are observed for the two alternatives. This paper updates the works of Luyben,18 in which a control system for fully heat-integrated pressure swing distillation for separating THF/water is developed, and results show that the control system can handle only fairly small disturbances without having purity specifications violated. In the present paper, all these results reveal that for this fully heat-integrated pressure swing distillation process, robust control may be achieved, even for large feed flow rate and feed composition disturbances. This energy-saving technology is worth considering in the conceptual design stage of azeotropic distillation processes. ’ AUTHOR INFORMATION Corresponding Author

*Tel.: +86 022-27404440. Fax: +86 0222-27404440. E-mail: [email protected]. 1309



’ ACKNOWLEDGMENT This work was supported by the Programme of Introducing Talents of Discipline to Universities (No. B06006). We are grateful to the staffs in the State Key Laboratories of Chemical Engineering (Tianjin University) for assistance. ’ NOMENCLATURE Bn = bottom flow rate from column n (kg/h) Dn = distillate flow rate from column n (kg/h) Fn = feed flow rate to column n (kg/h) HPC = high-pressure column IDn = internal diameter for column n (m) LPC = low-pressure column NFn = feeding location for feed stream Fn NR = feeding location for recycle stream NTn = number of theoretical plates for column n Pn = top pressure of column n QCn = condenser heat removal for column n (KW) QRn = reboiler heat input for column n (KW) RRn = reflux ratio for column n TAC = total annual costs ( $1000 per annum) TC = temperature controller TPC = pressure-compensated temperature (K) Tn,m = temperature for tray m in column n (K)

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(15) Dingt, S. S.; Luyben, W. L. Control of a Heat-Integrated Complex Distillation Configuration. Ind. Eng. Chem. Res. 1990, 29, 1240–1249. (16) Knapp, J. P.; Doherty, M. F. Thermal Integration of Homogeneous Azeotropic Distillation Sequences. AIChE J. 1990, 31, 969–984. (17) Repke, J. U.; Klein, A.; Forner, F. Homogeneous Azeotropic Distillation in an Energy- and Mass-Integrated Pressure Swing Column System. Comput.-Aided Chem. Eng. 2004, 18, 757–762. (18) Luyben, W. L. Design and Control of a Fully Heat-Integrated Pressure-Swing Azeotropic Distillation System. Ind. Eng. Chem. Res. 2008, 47, 2681–2695. (19) Luyben, W. L. Distillation Design and Control Using AspenTM Simulation; John Wiley & Sons, Inc.: Hoboken, NJ, 2006. (20) Seada, J. D.; Henley, E. J. Separation process principles; John Wiley & Sons, Inc.: Hoboken, NJ, 1998; pp 612. (21) Turton, R.; Bailie, R. C.; Whiting, W. B.; Shaeiwitz, J. A. Analysis, Synthesis and Design of Chemical Processes; Prentice Hall: Upper Saddle River, NJ, 2009. (22) Buckley, P. S, Luyben, W. L.; Shunta, J. P. Design of Distillation Column Control Systems; Instrument Society of America: Research Triangle Park, NC, 1985; pp 234. (23) Luyben, W. L. Comparison of Extractive Distillation and Pressure-Swing Distillation for Acetone-Methanol Separation. Ind. Eng. Chem. Res. 2008, 47, 2696–2707. (24) Wang, Q.; Yu, B.; Xu, C. Design and Control of Distillation System for Methylal/Methanol Separation: Part 1: Extractive Distillation Using DMF As an Entrainer. Ind. Eng. Chem. Res. 2012, 51, 10.0121/ie201949q.

’ REFERENCES (1) Satoh, S.; Yukio T. Process for Producing Methylal. U.S. Patent 6,379,507, April 30, 2002. (2) Knapp, J. P.; Doherty, M. F. A New Pressure-Swing-Distillation Process for Separating Homogeneous Azeotropic Mixtures. Ind. Eng. Chem. Res. 1992, 31, 346–357. (3) M€uller, W. H. E.; Kaufhold, M. Process for the Recovery of Pure Methylal from Methanol-Methylal Mixtures. U.S. Patent 4,385,965, May 31, 1983. (4) Lewis W. K. Dehydrating Alcohol and the Like. U.S. Patent 1,676,700, July 10, 1928. (5) Wasylkiewicz, S. K.; Kobylka, L. C.; Castillo, F. J. L. Pressure Sensitivity Analysis of Azeotropes. Ind. Eng. Chem. Res. 2003, 42, 207–213. (6) Repke, J. U.; Forner, F.; Klein, A. Separation of Homogenous Azeotropic Mixtures by Pressure Swing Distillation Analysis of the Operation Performance. Chem. Eng. Technol. 2005, 28, 1151–1157. (7) Phimister, J. R.; Seider, W. D. Semicontinuous, Pressure-Swing Distillation. Ind. Eng. Chem. Res. 2000, 39, 122–130. (8) Hamad, A.; Dunn, R. F. Energy Optimization of Pressure-Swing Azeotropic Distillation Systems. Ind. Eng. Chem. Res. 2002, 41, 6082–6093. (9) Luyben, W. L. Comparison of Pressure-Swing and ExtractiveDistillation Methods for Methanol-Recovery System in the TAME Reactive-Distillation Process. Ind. Eng. Chem. Res. 2005, 44, 5715–5725. (10) Varbanov, P.; Klein, A.; Repke, J. U.; Wozny, G. Minimising the Startup Duration for Mass- And Heat-Integrated Two-Column Distillation Systems: A Conceptual Approach. Chem. Eng. Process. 2008, 47, 1456–1469. (11) Huang, K.; Shan, L.; Zhu, Q.; Qian, J. Adding Rectifying/ Stripping Section Type Heat Integration to a Pressure-Swing Distillation (PSD) Process. Appl. Therm. Eng. 2008, 28, 923–932. (12) Modla, G.; Lang, P. Feasibility of New Pressure Swing Batch Distillation Methods. Chem. Eng. Sci. 2008, 63, 2856–2874. (13) Jana, A. K. Heat Integrated Distillation Operation. Appl. Energ. 2010, 87, 1477–1494. (14) Abu-Eishah, S. I.; Luyben, W. L. Design and Control of a TwoColumn Azeotropic Distillation System. Ind. Eng. Chem. Proc. Des. Dev. 1985, 24, 132–140. 1310


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