Design and Control of Different Pressure Thermally Coupled Reactive

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Design and control of different pressure thermally coupled reactive distillation for methyl acetate hydrolysis Lumin Li, Lanyi Sun, Jun Wang, Jian Zhai, Yuliang Liu, Wang Zhong, and Yuan-yu Tian Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.5b03041 • Publication Date (Web): 24 Nov 2015 Downloaded from http://pubs.acs.org on November 28, 2015

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Industrial & Engineering Chemistry Research

Design and control of different pressure thermally coupled reactive distillation for methyl acetate hydrolysis Lumin Li,a Lanyi Sun,*,a Jun Wang,b Jian Zhai,a Yuliang, Liua , Wang Zhonga and Yuanyu Tiana a

State Key Laboratory of Heavy Oil Processing, China University of Petroleum, Qingdao 266580, China

b

China Petroleum Engineering Construction Corporation Dalian Design Company, Dalian 116011, China

ABSTRACT Thermally coupled distillation is designed to reduce energy requirements and achieve environmental benefits in chemical engineering development process. In this study an attractive different pressure thermally coupled reactive distillation (DPTCRD) process was proposed for the hydrolysis of methyl acetate，and the DPTCRD process was optimized to evaluate its economic feasibility. The calculation results clearly demonstrated that the total annual cost of this thermally coupled sequence can be saved by 7.49% and its exergy consumption is reduced by 40.07% compared with the conventional reactive distillation. In addition, the controllability of this promising sequence was evaluated, and the dynamic results showed that the system with improved control structure can maintain stable operation and handle large feed upsets effectively. 1. INTRODUCTION Distillation process is the most common unit operation used in chemical and petrochemical industry, while it also requires a great deal of energy to obtain high-purity products. Therefore a variety of enhanced distillation techniques have been extensively studied to reduce the energy requirements. Reactive distillation integrates the reaction and separation into a single column, and it offers several potential benefits, such as improving reaction conversion, increasing product selectivity and reducing energy requirements.1 On the other hand, thermally coupled distillation technologies have received great attention, which can achieve both energy and capital cost savings 1

ACS Paragon Plus Environment


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through eliminating the condensers/reboilers of the classical distillation columns.2-5 Heat integrated distillation column (HIDiC) is proposed as a promising separation technology, which consists of a rectifying column and a stripping column. A lot of internal heat exchangers are used to transfer heat between the overhead vapor of the rectifier and the bottom liquid of the stripper, through which reboiler duty can be reduced significantly.6 Lee et al.7 investigated the HIDiC process for hydrolysis of methyl acetate (MeAc). Although this HIDiC process can achieve 8.05% operating cost savings, its total annual cost (TAC) is much higher than that of the basic process for the high compressor cost required in HIDiC system. Hydrolysis reaction of methyl acetate is a typical system used in industrial polyvinyl alcohol (PVA) plants, while the conversion of MeAc reactant is very low because of the small equilibrium constant. Moreover, MeAc will form azeotropic mixtures with methanol (MeOH) and water, respectively. Thus the distillation process for MeAc hydrolysis is greatly energy-consuming and need relatively large-scale equipment, and it is an important target for researchers to improve the hydrolysis process.8 In conventional process of MeAc hydrolysis, the bottoms of reactive distillation column mainly contains MeOH, HAc and H2O, thus two separation columns are demanded to separate the three components. Gao et al.9 reported a new process for MeAc hydrolysis, which only included a reactive distillation column and a separation column, and in this new process, the reactive distillation column adopted a sidestream that was fed to the separation column. The simulation results showed that this new two-column process can achieve 28.6% energy savings compared with the conventional three-column process. Besides, a multieffect distillation process was developed by Lee et al.7 for the hydrolysis of MeAc, and 6.42% reduction in TAC can be obtained through this heat-integrated distillation sequence. 2


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Gao and his co-workers10 proposed an interesting structure of different pressure thermally coupled reactive distillation (DPTCRD) for tert-amyl methyl ether (TAME) synthesis. In that typical DPTCRD design, the reactive distillation column was divided into two parts that included a high-pressure (HP) column and a low-pressure (LP) column, and the top stream from HP column was used as the heating medium of LP column reboiler. It was demonstrated that the TAC of the DPTCRD process was 6.56% less than that of the conventional reactive distillation (CRD) process. The attractive DPTCRD design inspired us to explore its feasibility to MeAc hydrolysis system. In the work of Gao et al.,10 TAME product can be directly obtained in the bottoms of reactive distillation column, thus only one column was needed in the system. However, the MeAc hydrolysis is a reversible reaction with small equilibrium constant, hence excess water is demanded to improve the conversion of MeAc, resulting in the requirement of two subsequent distillation columns. Therefore, it is indeed necessary to explore the dynamic performance of this DPTCRD configuration along with the subsequent separation sequences. In the present work, this attractive DPTCRD design will be applied to the MeAc hydrolysis. Firstly the hydrolysis system was simulated in both the conventional and DPTCRD processes, and then the economical optimization and exergy analysis of the two processes were conducted. Furthermore, the control performance of this DPTCRD process was examined, and some conclusions will be given in the last section. 2. PROCESS DESCRIPTION This work considers the production of acetic acid (HAc) and MeOH by the liquild-phase reaction of MeAc hydrolysis in the presence of an acid catalyst Amberlyst 15. The reaction can be 3



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described by the following expression (1), MeAc + H2 O ← → HAc + MeOH

(1)

Pöpken et al.11 reported the kinetic model for this chemical reaction system which is described as follows,

r = mcat ×

k f a'MeAc a'H2 O - kr a'HAc a 'MeOH (a 'MeAc + a 'H 2 O + a 'HAc + a 'MeOH ) 2

(2)

where mcat (kg) is the mass of the catalyst; a 'i = K i ai / M i , with Mi denoting the molar mass of component i, ai is the activity coefficient of component i, Ki is the adsorption equilibrium constant of component i, KMeAc=4.15, KH2O=5.24, KHAc=3.15, KMeOH=5.64, whereby the rate constants kf and kr (kmol/(kgcat·s)) are temperature (T) dependent, as described, k f = 6.127 × 105 exp(-6.373 × 104 / RT )

(3)

k r = 8.497 × 106 exp(-6.047 × 104 / RT )

(4)

The unit of reaction rate r is kmol/s; R (=8.314 kJ/(kmol·K)) is the gas constant, and T is in K. The simulation of MeAc hydrolysis process is implemented using the rigorous equilibrium stage module RadFrac in Aspen Plus. Since HAc is associated in vapor phase, we adopt the Hayden-O’Conell model to account for vapor nonideality and UNIQUAC model to calculate the liquid activity coefficient.7 Liquid holdup was calculated based on the column diameter with the assumption of 10 cm weir height. Besides, the catalyst with density of 770 kg/m3 occupies half of the tray holdup volume. 2.1. Conventional Reactive Distillation Process Lin et al.8 summarized four types of possible structures for the hydrolysis of MeAc, while it was hard to determine which one was better. In this hydrolysis system the equilibrium constant is very low and consequently excessive water is used to increase the reaction conversion. Figure 1 gives a 4


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flowsheet of the conventional reactive distillation process for MeAc hydrolysis. In Figure 1, the catalyst is fixed in the middle of reactive distillation (RD) column, and the fresh feed mixture containing 65.0 mol% MeAc and 35.0 mol% MeOH enters the bottom of reactive section. The water is fed to the top of catalyst bed. The distillate of RD column mainly containing MeOH and unreacted MeAc is recycled to mix with the fresh feed, and the bottom product of RD column containing less than 0.1 mol% MeAc is fed to the HAc column. 99.0 mol% HAc product is obtained in the bottoms of HAc column, and its distillate is further separated in the third column to obtain MeOH product.

Figure 1. Conventional reactive distillation process for MeAc hydrolysis. 2.2. Different Pressure Thermally Coupled Reactive Distillation Process

Figure 2. Different pressure thermally coupled reactive distillation process for MeAc hydrolysis. The different pressure thermally coupled concept has been applied to the RD column, and the corresponding flowsheet of DPTCRD process is shown in Figure 2. The RD column is divided 5



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into two parts, namely a HP column and a LP column. Compared with the traditional RD column, the HP column is equal to the rectifying section and reactive section, and the LP column acts as the stripping section. The liquid outflow from the bottom of HP column flows into the top of LP column. The overhead vapor of LP column is compressed and then driven into the bottom of HP column. The vapor stream from HP column is used as the heat source of the LP column and then cooled by a trim-condenser. The outflow with specification of 0.1 mol% MeAc in LP column bottom is fed to subsequent separation sequences. It should be noted that the feed conditions and product specifications in DPTCRD process are consistent with those in CRD process to ensure a fair comparison. 3. ECONOMICAL OPTIMIZATION The economical optimization is carried out to evaluate properly the economic feasibility of DPTCRD design. In this DPTCRD process, a number of design variables need to be determined, which include the number of stages in each column (NHP, NLP, NHAc, NMeOH), the number of stages in reaction section (NR), the feed location of each column (NF1, NF2, NF3, NF4), and the reflux ratio of each column (RHP, RLP, RHAc, RMeOH). Here, Ni means the number of stages in column i; NF1 and NF2 are the two feed locations of HP column, NF3 and NF4 are the feed locations of HAc column and MeOH column respectively; Ri is the reflux ratio of column i. 3.1. Sensitivity Analysis Pressures in HP column and LP column are essential for the DPTCRD process, which have great influence on the product purities and energy requirements. For simplification purposes, the pressure of LP column (PLP) remains at 1.107 atm, and the pressure of HP column (PHP) will be varied from 3.2 atm to 7.4 atm to obtain the optimal value. Here, we chose 7.4 atm as the upper 6


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limit of pressure mainly because there is a limit on the operating pressure and temperature.7 Note that the minimum heat transfer temperature difference is set as 10 oC in this system, which is crucial to achieve a normal heat transfer. Figure 3 illustrates how the pressure of HP column affects the MeAc composition in the bottoms of LP column, heat transfer temperature difference, compressor duty, condenser duty of HP column and reboiler duty of LP column. It is evident that, as PHP increases from 3.2 to 7.4 atm, the MeAc molar composition increases slightly, and it is appropriate to meet the purity specification when PHP is less than 5 atm. When the pressure of HP column increases, the top temperature of HP column rises from 95.23 oC to 128.27 oC and the bottom temperature of LP column keeps at around 89 oC, resulting in the increase of heat transfer temperature difference. At the same time, the compressor duty rises with the increase of PHP. Therefore, under the premise of minimum heat transfer temperature difference and the temperature ceiling of 120 oC, PHP should be reduced as far as possible. Finally 3.6 atm has been selected for PHP. As shown in Figure 3(d), the reboiler duty of LP column is always less than the condenser duty of HP column, hence the required heat for LP column can be completely provided by the overhead steam of HP column, and an auxiliary condenser is demanded in HP column to remove the redundant heat. (b) 130

130

120

120

110

110

o

Top temperature of HP column ( C)

0.004

0.003

0.002

0.001

0.000 2.5 3.0 3.5 4.0 4.5 5.0 5.5 6.0 6.5 7.0 7.5 8.0

100

100 Temperature difference=10.42oC

90

90

80 80 2.5 3.0 3.5 4.0 4.5 5.0 5.5 6.0 6.5 7.0 7.5 8.0

PHP (atm)

PHP (atm)

7


Bottom temperature of LP column (oC)

(a)

MeAc composition (mol%)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60



(c)

(d)

Condenser duty of HP column (kW)

1400 1300 1200 1100 1000 900 800 2.5 3.0 3.5 4.0 4.5 5.0 5.5 6.0 6.5 7.0 7.5 8.0

4800 -4600 4600 -4700 4400 -4800 4200 -4900 4000 -5000 3800 -5100 3600 -5200 3400 -5300 2.5 3.0 3.5 4.0 4.5 5.0 5.5 6.0 6.5 7.0 7.5 8.0

Reboiler duty of LP column (kW)

-4500

1500 Compressor duty (kW)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

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PHP (atm)

PHP (atm)

Figure 3. Effects of pressure of HP column on (a) MeAc molar composition in bottoms of LP column, (b) top temperature of HP column and bottom temperature of LP column, (c) compressor duty, and (d) condenser duty of HP column and reboiler duty of LP column. 3.2. Optimization of the DPTCRD Process

Figure 4. Sequential iterative optimization procedure for DPTCRD design. It is important to simplify the optimization procedure since too many design variables should be determined to find the optimal values under the specified product purities. Therefore, the 8


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sequential iterative optimization procedure is used to find the optimal design with the Ni (i.e., NHP, NLP, NHAc or NMeOH) as the outer iterative loop and NFi (i.e., NF1, NF2, NF3 or NF4) as the inner iterative loop. The optimization procedure to minimize the TAC is presented in Figure 4. TAC is used as the objective function, which can be described as Eq. (5), TAC ($/year) = OC + CI/T

(5)

Here, OC is the operating cost, CI is the capital investment and T is the payback period with 3 years. The operating cost contains the steam cost for reboiler, the cost of cooling water, the catalyst cost and the electricity cost for compressor. The capital investment includes the cost for column trays, column shell, heat exchangers and compressor excluding the small equipment like pumps, valves for their low costs. A plant availability of 8000 hours/year is assumed. The diameter of the column is determined by Tray Sizing option in RadFrac block of Aspen Plus, and the materials of construction are carbon steel. Sieve tray is chosen as the tray type, and the tray spacing is specified as 0.6096 m. All equations for determining TAC are mentioned in Douglas’s book12 as shown in the Supporting Information. The overall heat transfer coefficients and utility prices shown in Table 1 are taken from the relevant papers.13-16 The Marshall & Swift index (M&S) is specified at 1468.2.17 Table 1 Utility prices. Utility

Price o

Low-pressure steam (147 C, 50 psig) Cooling water Catalyst Electricity

6.60 $/103 kg 0.03 $/103 gal 7.7162 $/kg 0.0605 $/(kW·h)

For a fixed NHP, the NF1, NF2 and NR are varied to find a lowest point that gives the minimum TAC in the HP column. Figure 5 shows that, for a fixed NHP, the CI, OC and TAC are greatly influenced by NR. Figure 6-8 demonstrate how the CI, OC and TAC are affected by NHP, NHAc, and 9



NMeOH, respectively. Finally, we choose that NHP =14, NR =14, NF1 =1, NF2 =1, NLP=10, NHAc=27, NF3=12, NMeOH =22, and NF4 =11 with the lowest TAC of 2746.25 × 103 $/year. The detailed operating parameters of DPTCRD process are summarized in Table 2. Table 2. Operating parameters of DPTCRD process Parameters

HP column

LP column

HAc column

MeOH column

Number of stages Reactive stages

14 1-14

10 —

27 —

22 —

Liquid holdup (m3)

0.14

—

—

—

MeAc/MeOH feed stage H2O feed stage

1 1

1

12

11

MeAc/MeOH feed flowrate (kmol/h)

76.92

975.88

326.92

304.04

1.24 1.107 1.62 4387.32

0.08 1 1.51 -3614.56 3588.14

1.79 1 0.99 -2092.94 2181.87

H2O feed flowrate (kmol/h)

250

Reflux ratio Operating pressure (atm) Column diameter (m) Condenser duty (kW) Reboiler duty (kW)

7.75 3.6 1.50 -4858.65 -

(a)

(b). 6530

6090

8560

6525

6075

8550

6520

6045

6515

6030

6510

6015

6505

6000

6500

5985

6495

8490

6490

8480

11

12

13

14

3

3

5970

TAC (10 $/year)

6060

OC (10 $/year)

3

CI (10 $)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

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8540 8530 8520 8510 8500

11

NR

12

13

NR

Figure 5. Effects of the NR on (a) capital investment and operating cost, (b) total annual cost.

10


14

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(b)

5970

6498

5965

6497

5960

6496

5955

6495

5950

6494

5945

6493

5940

6492

5935

6491

8490 8485

5930

14

15

16

NHP

17

8480

3

6499

TAC (10 $/year)

6500

5975

3

5980

OC (10 $/year)

3

CI (10 $)

(a)

6490

8475 8470 8465

14

15

16

NHP

17

Figure 6. Effects of the NHP on (a) capital investment and operating cost, (b) total annual cost. (a)

(b)

3800

3520

2200

3750 3700

3360

1900

3500

1800

3450

3

3550

TAC (10 $/year)

3600

3280

3

2000

OC (10 $/year)

3

CI (10 $)

3440

2100

3650

3200 3120 3040 2960 2880 2800

1700 3400

2720

1600

3350 18

20

22

24

26

28

30

32

34

2640 18

36

20

22

24

26

28

30

32

34

36

NHAc

NHAc

Figure 7. Effects of the NHAc on (a) capital investment and operating cost, (b) total annual cost. (a)

(b) 3428

2755

1614

2754 1612

3

1608

3

3

3426

2753

TAC (10 $/year)

1610

OC (10 $/year)

3427

CI (10 $)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


3425 1606 3424

1604

2752 2751 2750 2749 2748 2747 2746

3423

18

19

20

21

22

23

24

25

1602

2745

18

19

NMeOH

20

21

22

23

24

25

NMeOH

Figure 8. Effects of the NMeOH on (a) capital investment and operating cost, (b) total annual cost. 3.3. Optimization of the CRD Process The CRD process is also optimized by a sequential iterative optimization method to make a fair comparison between the CRD and DPTCRD processes. A number of parameters of the CRD 11



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process are optimized on the basis of the TAC, and Figure 9 illustrates the systematic global optimization sequence for the conventional system. Finally we choose that NRD =31, NR =18, NF1 =1, NF2 =1, NHAc=28, NF3=12, NMeOH =22, and NF4 =11 with the lowest TAC of 2968.89 × 103 $/year. The detailed operating parameters of CRD process are summarized in Table 3. Table 3. Operating parameters of CRD process Parameters

RD column

HAc column

MeOH column

Number of stages Reactive stages Liquid holdup (m3) MeAc/MeOH feed stage H2O feed stage MeAc/MeOH feed flowrate (kmol/h) H2O feed flowrate (kmol/h) Reflux ratio Operating pressure (atm) Column diameter (m) Condenser duty (kW) Reboiler duty (kW)

31 1-18 0.15 1 1 76.92 250 5.98 2.2 1.55 -8312.66 8939.97

28 — —

22 — —

12

11

326.92

304.04

0.12 0.99 1.49 -3750.78 3488.09

1.78 0.99 0.99 -2087.33 2176.51

Figure 9. Sequential iterative optimization procedure for CRD design. 12


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3.4. Comparison of the CRD and DPTCRD processes 3200 2800

CRD DPTCRD

-7.49%

3 Cost (10 $/year)

2400 2000 -21.34%

1600 1200 +24.74%

800 400 0

Operating Cost

Capital Cost

Total Annual Cost

Figure 10. Cost comparison between the CRD and DPTCRD processes. Figure 10 presents the comparison of the operating cost, annualized capital cost and TAC for the CRD process versus the proposed DPTCRD process. The results indicate that the operating cost saving of DPTCRD process is 21.34%, while the capital cost increases by 24.74% because of the high expenditure on compressor. Although the annualized capital cost of the DPTCRD process increases, its TAC is still lower than that of the CRD process by 7.49%. (a)

(b) 1.0 0.9 0.8 0.7 0.6 0.5 0.4 0.3 0.2 0.1 0.0

MeAc Water HAc MeOH

1

4

Liquid mole composition

Vapor mole composition

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


7 10 13 16 19 22 25 28 31 Stage

1.0 0.9 0.8 0.7 0.6 0.5 0.4 0.3 0.2 0.1 0.0

MeAc Water HAc MeOH

1

4

7 10 13 16 19 22 25 28 31 Stage

(c)

13



115

105

o

Temperature ( C)

110

100 95 90 85 80

1

4

7 10 13 16 19 22 25 28 31 Stage

Figure 11. (a) Vapor composition profile, (b) liquid composition profile, and (c) temperature profile of the reactive distillation column in CRD process. (b) 1.0 0.9 0.8 0.7 0.6 0.5 0.4 0.3 0.2 0.1 0.0

MeAc Water HAc MeOH

HP

2

4 6

Liquid mole composition

Vapor mole composition

(a)

LP

8 10 12 14 16 18 20 22 24 Stage

1.0 0.9 0.8 0.7 0.6 0.5 0.4 0.3 0.2 0.1 0.0

MeAc Water HAc MeOH

HP

2 4 6

LP

8 10 12 14 16 18 20 22 24 Stage

(c)

o

Temperature ( C)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

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105 100 95 90 85 80 75 70 65 60

HP

2

4

6

LP

8 10 12 14 16 18 20 22 24 Stage

Figure 12. (a) Vapor composition profiles, (b) liquid composition profiles, and (c) temperature profiles of the reactive distillation column in DPTCRD process. Both composition profiles and temperature profiles of the reactive distillation column in CRD and DPTCRD processes are respectively shown in Figure 11 and Figure 12. It reveals that the vapor 14


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and liquid composition profiles of the RD column in DPTCRD process are similar to those in CRD process. In Figure 12, although there is a small fluctuation of MeAc composition near the top of LP column, the purity of MeAc in the bottoms of LP column is low enough to satisfy the specification. Besides, the temperature of reactive section in HP column is higher and more stable than that in the basic column, which is in favor of the positive reaction. 4. EXERGY ANALYSIS The energy requirements of distillation process focus on the heat and power energy consumption. In order to account for the difference in the qualities of the heat and power energy, the exergy analysis is adopted in this study. Generally, the work required for separation is supplied by heat. However, the heat suffers from great losses when it converts into work. If the condenser and reboiler of distillation column have nearly equal heat duties, the work available from the heat energy can be described as the following relation:18 Wheat = ExR − ExC = QRTa (

1 1 − ) TC TR

(2)

where ExR and ExC are the exergy of the reboiler and condenser, respectively, QR is the reboiler duty, Ta is the ambient temperature, TC is the cooling media temperature and TR is the heating media temperature. To evaluate the exergy requirement of reboiler, a simplified Eq. (2) can be introduced as the form below:19   Ta ExR = QR 1 −  T + ∆ T  B 

(3)

where TB is the boiling point temperature of bottoms, ∆T is the temperature difference of the heating medium used in the reboiler. Here, 298.15K is adopted as the ambient temperature, and the low-pressure steam with 420.15K is chosen as the heating medium. The results of exergy analysis are shown in Table 4. It can be seen from Table 4, the DPTCRD process provides 42.70% 15



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reduction in the total energy requirements. In the case of DPTCRD process, the energy is also added into the system in a form of compressor work. As a result, the overall exergy of the DPTCRD process can be calculated by adding up the reboiler exergy and the compressor duty. The overall exergy consumption of the DPTCRD process and CRD process are 2541.63 kW and 4240.77 kW, respectively, and savings of about 40.07% in overall exergy is achieved by the DPTCRD process. This is mainly because the exergy of the condenser in HP column of the DPTCRD system is recovered rather than completely discharged as in that of the CRD process. Table 4. Results of exergy analysis CRD process

Reboiler duty (kW) Compressor duty (kW) Total energy requirements (kW) (% difference) Exergy of reboiler (kW) Overall exergy (kW) (% difference)

DPTCRD process

RD column

HAc column

MeOH column

HAc column

MeOH column

8939.97 —

3488.09

2176.51

3588.14 866.17

2181.87

8368.52 (-42.70%)

14604.57 (0%) 2595.92

1012.85

4240.77 (0%)

632.00

1041.90

633.56

2541.63 (-40.07%)

5. CONTROL STRUCTURE DESIGN Although the promising DPTCRD process can be implemented smoothly at steady state, the real operation of this system might not be an easy problem because of the strong interactions through the streams connecting the LP column and HP column. As a result, it’s of great importance to study the controllability of this heat integration system. In this section two plantwide control schemes were considered for the DPTCRD process and their dynamic performances were investigated for large disturbances in throughput and feed composition. The Aspen Dynamics simulator was used to conduct the dynamic simulations. The heat integration for this DPTCRD system can be achieved by using the function of “flowsheet equations” in Aspen 16


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Dynamics. As shown in Figure 13, two equations are entered in the text editor window. The reboiler heat duty of LP column is calculated by the first equation, where the energy supply is determined by the product of the heat transfer area, the overall heat transfer coefficient and the current temperature difference between the top of HP column and the bottom of LP column. The second equation shows that the pressure in HP column is changed based on the difference of the condenser/reboiler duties of the thermally coupled distillation columns. After compiling the flowsheet equations, the condenser duty of HP column and the reboiler duty of LP column are changed from “fixed” to “free” to eliminate the over specified variables.

Figure 13. Heat integration equations implemented in Aspen Dynamics. 5.1. Basic Control Structure (a)

(b)

1.5 1.0

0.15

- 0.5% QR of HAc column + 0.5% QR of HAc column

0.10

0.5 o

0.0 -0.5 -1.0 -1.5

- 0.5% RR of MeOH column +0.5% RR of MeOH column

0.05 ∆T ( C)

o

∆T ( C)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


0.00 -0.05 -0.10 -0.15

3 6 9 12 15 18 21 24 27 Number of stage

2 4 6 8 10 12 14 16 18 20 22 Number of stage

Figure 14. Open-loop responses for the (a) HAc column and (b) MeOH column. The control schemes for the HAc column and MeOH column are firstly conducted. Considering the widespread use of temperature controller in the chemical industry, the product composition 17



1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

will be inferred by some tray temperature control loops. In this DPTCRD system, the temperature of stage 24 in HAc column is controlled by the column’s reboiler duty, and in MeOH column the temperature in tray 10 is regulated by its reflux ratio. These controlled trays are chosen based on the open-loop sensitivity analysis20 as shown in Figure 14. Besides, other regulatory controls are implemented. The reflux drum levels in the separation columns are kept by manipulating their corresponding distillate flow rate, and the base levels are held by adjusting the bottom flow. In addition, the operation pressures in the two columns are maintained by changing their condenser heat duties. Next, the control strategy for the heat integration part of the DPTCRD process will be explored. The heat integration system can be considered as the combination of conventional pressure-swing distillation and reactive distillation. Thus a basic control structure is developed for this DPTCRD process as presented in Figure 15. The two feed flowrates are proportional and the multiplier receives its signal from the composition controller (CC), which supervises the MeAc composition in the bottoms of HP column. Herein we adopted the composition controller because we found that the DPTCRD system cannot run steadily with the use of temperature controller. The reflux drum level in HP column is held by its reflux rate, and the base levels in HP column and LP column are all controlled by their corresponding bottom flow. In addition, the pressure of HP column is held by manipulating the heat removal rate of auxiliary condenser, and the pressure of LP column is kept by altering the overhead vapor flowrate through the compressor. The control loops implemented in Aspen Dynamics and their control faceplates are shown in Figure 15.

18


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(a)

(b)

Figure 15. (a) Initial control structure and its related (b) controller faceplates. Level loops are proportional with the gain KC = 2, and proportional-integral controllers are used in all the other control loops. Flow controllers are set with KC = 0.5 and integral time constant τI = 0.3 min. The final KC and τI for the composition and temperature controllers are obtained using the Tyreus-Luyben tuning settings. After the system runs steadily, a ±20% disturbance in MeAc/MeOH feed mixture is introduced to explore the effectiveness of the initial control structure.

19



0.984 2

4

78 76 74 72 70 68 66 64 62

0.98

+20% F -20% F

0

2

4

6 8 10 12 14 Time (hr)

+20% F -20% F

0.43 0.42

0.97

0.41

0.96 0.95 0

6 8 10 12 14 Time (hr)

+20% F -20% F

2

4

0.40 0

6 8 10 12 14 Time (hr)

119.0

o

o

Tray 6 of LP column ( C)

0.980 0

0.44

0.99

o

0.988

Tray 24 of HAc column ( C)

HAc (mol%)

0.992

1.00

Tray 10 of MeOH column ( C)

+20% F -20% F

0.996

MeOH (mol%)

1.000

XBMeAc of HP column (mol%)

(a)

+20% F -20% F

118.5 118.0 117.5 117.0 116.5 116.0 0

2

4

6 8 10 12 14 Time (hr)

2

4

6 8 10 12 14 Time (hr)

79 +20% F -20% F

78 77 76 75 74 0

2

4

6 8 10 12 14 Time (hr)

0.984 0.980 0

4

78 76 74 72 70 68 66 64 62

0.98 0.97 0.96 0.95 0

6 8 10 12 14 Time (hr)

+20% MeAc -20% MeAc

2

4

6 8 10 12 14 Time (hr)

119.0

o

o

2

0.99

o

0.988


HAc (mol%)

0.992

MeOH (mol%)

+20% MeAc -20% MeAc

0.996

+20% MeAc -20% MeAc

0

2

4

6 8 10 12 14 Time (hr)

+20% MeAc -20% MeAc

118.5 118.0 117.5 117.0 116.5 116.0 0


1.00

1.000


(b)


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

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2

4

6 8 10 12 14 Time (hr)

0.44 +20% MeAc -20% MeAc

0.43 0.42 0.41 0.40 0

2

4

6 8 10 12 14 Time (hr)

79 +20% MeAc -20% MeAc

78 77 76 75 74

0

2

4

6 8 10 12 14 Time (hr)

Figure 16. Dynamic results of the basic control structure in ± 20% (a) feed flowrate and (b) MeAc composition changes. Figure 16 reveals the relevant dynamic results when the heat-integrated distillation system suffers the feed flowrate and composition changes at one hour. As shown in Figure 16, this DPTCRD system can finally settle down in the face of various feed disturbances, and the HAc product can arrive at a new stable value with desired purity. However, the purity of MeOH product has a large 20


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offset when the feed flowrate or MeAc composition increases. It has been revealed that the effective temperature control on stage 10 of MeOH column cannot guarantee the purity of MeOH product. On the other hand, the sensitivity tray (stage 6) temperature of LP column can reflect the MeOH product quality to an extent. This is mainly because the water content in LP column directly affects the product purities in the subsequent separation processes, especially in MeOH column, and at the same time the temperature of LP column is also influenced by the water content. 5.2. Improved Control Structure In order to overcome the drawbacks of the basic control structure, the control of stage 6 temperature in LP column should be taken into account. Since the LP column and HP column are heat-integrated distillation sequence, the reboiler duty of LP column cannot be adjusted alone as the manipulated variable, which is different from the conventional case. Thus, in the improved control structure, the temperature on stage 6 of LP column will be controlled by the overhead vapor flowrate with PC2 on cascade. Table 5 presents the final tuning parameters of the controllers21 in the improved control structure, and the related dynamic results are illustrated in Figure 17. Table 5. Tuning parameters of the composition and temperature controllers Controller gain KC (%)

Controller integral time τI (min)

CC TC6 TC24

2.94 0.85 5.40

71.28 9.24 7.92

TC10

4.18

11.88

21



MeOH (mol%)

0.992 0.988 0.984 2

4

o

78 76 74 72 70 68 66 64 62

+20% F -20% F

0

2

4

6 8 10 12 14 Time (hr)

0.990 0.985

+20% F -20% F

0.980 0.975 0.970 0

6 8 10 12 14 Time (hr) Tray 24 of HAc column ( C)

o


0.980 0

0.995

2

4

6 8 10 12 14 Time (hr)

119.0

o

HAc (mol%)

1.000

+20% F -20% F

0.996


1.000


(a)

+20% F -20% F

118.5 118.0 117.5 117.0 116.5 116.0 0

2

4

6 8 10 12 14 Time (hr)

0.44 +20% F -20% F

0.43 0.42 0.41 0.40 0

2

4

6 8 10 12 14 Time (hr)

79 +20% F -20% F

78 77 76 75 74

0

2

4

6 8 10 12 14 Time (hr)

0.984 0.980 0

4

78 76 74 72 70 68 66 64 62

0.990 0.985 0.975 0.970 0

6 8 10 12 14 Time (hr)

+20% MeAc -20% MeAc

0.980

2

4

6 8 10 12 14 Time (hr)

119.0

o

o

2

0.995

o

0.988


HAc (mol%)

0.992

MeOH (mol%)

+20% MeAc -20% MeAc

0.996

+20% MeAc -20% MeAc

0

2

4

6 8 10 12 14 Time (hr)

+20% MeAc -20% MeAc

118.5 118.0 117.5 117.0 116.5 116.0 0


1.000

1.000


(b)


1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 22 of 28

2

4

6 8 10 12 14 Time (hr)

0.44 +20% MeAc -20% MeAc

0.43 0.42 0.41 0.40 0

2

4

6 8 10 12 14 Time (hr)

79 +20% MeAc -20% MeAc

78 77 76 75 74

0

2

4

6 8 10 12 14 Time (hr)

Figure 17. Dynamic results of the improved control structure in ± 20% (a) feed flowrate and (b) MeAc composition changes. As shown in Figure17, the dynamic responses of HAc product in this improved control structure are similar with those in the basic control structure, and the HAc purity maintains close to its specification. It can be seen that the stage 6 temperature in LP column can always come back to its original value when the system undergoes different feed disturbances, which demonstrates that the 22


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control of stage 6 temperature in LP column is effective in this improved control structure. At the same time, the MeOH purity can return to a new steady state with negligible deviation as expected. Given the above, the results indicate that the improved control structure can handle large feed stream changes well and maintain the products at their desired purities. Therefore, this heat integration system can be well operated without deterioration on the controllability. 6. CONCLUSIONS In this work a novel different pressure thermally coupled technology was proposed for the hydrolysis of methyl acetate. The CRD and DPTCRD processes were simulated using Aspen Plus simulator, and their economical optimizations were carried out based on the calculation of minimum TAC. The results show that savings of 7.49% in TAC are achieved by the DPTCRD process. In addition, the exergy consumption of the DPTCRD design was reduced around 40.07% because of the heat integration between the HP column and LP column. The overall control strategies of the DPTCRD process were studied. For the initial control structure, the MeOH product composition could be stable but with large offsets after feed disturbances were introduced. However, after adding the temperature control of LP column, the improved control structure can handle large variation after introducing feed disturbances and keep the products at their desired purities, which shows that this attractive DPTCRD process could be operated normally with reasonable control.

23



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ASSOCIATED CONTENT Supporting Information Detailed calculation procedures of TAC. The Supporting Information is available free of charge on the ACS Publications website.

AUTHOR INFORMATION Corresponding Author Lanyi Sun. Tel.: +86 13854208340. Fax: +86 0532 86981787. E-mail address: [email protected].

ACKNOWLEDGEMENTS Financial supports of the National Natural Science Foundation of China (Grant: 21276279 and Grant: 21476261) and the Fundamental Research Funds for the Central Universities (No. 14CX05030A; No. 15CX06042A) are acknowledged with gratitude. Finally the authors are grateful to the editor and the anonymous reviewers.

NOMENCLATURE CC = composition controller CI = capital investment CRD = conventional reactive distillation DPTCRD = different pressure thermally coupled reactive distillation ExC = exergy of the condenser ExR = exergy of the reboiler HAc = acetic acid 24


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HIDiC = heat integrated distillation column HP column = high-pressure column KC = controller gain LP column = low-pressure column MeAc = methyl acetate MeOH = methanol M&S = Marshall & Swift index NF1 = water feed location of high-pressure column NF2 = MeAc/MeOH feed location of high-pressure column NF3 = feed location of HAc column NF4 = feed location of MeOH column Ni = the number of stages in column i NR = the number of stages in reaction section OC = operating cost PHP = pressure of high-pressure column PLP = pressure of low-pressure column PVA = polyvinyl alcohol QR = reboiler duty RD = reactive distillation Ri = reflux ratio of column i T = payback period TAC = total annual cost 25



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TAME = tert-amyl methyl ether Ta = ambient temperature TC= temperature controller TC = temperature of the cooling media TR = temperature of the heating media GREEK SYMBOL τI = controller integral time constant REFERENCES (1) Mali, S. V.; Jana, A. K. A partially heat integrated reactive distillation: feasibility and analysis. Sep. Purif. Technol. 2009, 70, 136-139. (2) Ramírez-Corona, N.; Mascote-Pérez, D.; Sánchez-Hijar, A.; Fernández-Pastrana, M. I.; Jiménez-Gutiérrez, A. Insights on the dynamic behavior of thermally coupled distillation columns implemented on processes with recycles. Chem. Eng. Res. Des. 2015, 93, 120-135. (3) Rong, B. G.; Kraslawski, A.; Nyström, L. Design and synthesis of multicomponent thermally coupled distillation flowsheets. Comput. Chem. Eng. 2001, 25, 807-820. (4) Errico, M.; Rong, B. G.; Tola, G.; Turunen, I. Process intensification for the retrofit of a multicomponent distillation plant an industrial case study. Ind. Eng. Chem. Res. 2008, 47, 1975-1980. (5) Rong, B. G. Synthesis of dividing-wall columns (DWC) for multicomponent distillations-A systematic approach. Chem. Eng. Res. Des. 2011, 89, 1281-1294. (6) Shenvi, A. A.; Herron, D. M.; Agrawal, R. Energy efficiency limitations of the conventional heat integrated distillation column (HIDiC) configuration for binary distillation. Ind. Eng. Chem. 26


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Res. 2010, 50, 119-130. (7) Lee, H. Y.; Lee, Y. C.; Chien, I. L.; Huang, H. P. Design and control of a heat-integrated reactive distillation system for the hydrolysis of methyl acetate. Ind. Eng. Chem. Res. 2010, 49, 7398-7411. (8) Lin, Y. D.; Chen, J. H.; Cheng, J. K.; Huang, H. P.; Yu, C. C. Process alternatives for methyl acetate conversion using reactive distillation. 1. Hydrolysis. Chem. Eng. Sci. 2008, 63, 1668-1682. (9) Gao, X.; Li, X.; Li, H. Hydrolysis of methyl acetate via catalytic distillation: simulation and design of new technological process. Chem. Eng. Process. 2010, 49, 1267-1276. (10) Gao, X.; Wang, F.; Li, H.; Li, X. Heat-integrated reactive distillation process for TAME synthesis. Sep. Purif. Technol. 2014, 132, 468-478. (11) Pöpken, T.; Steinigeweg, S.; Gmehling, J. Synthesis and hydrolysis of methyl acetate by reactive distillation using structured catalytic packings: experiments and simulation. Ind. Eng. Chem. Res. 2001, 40, 1566-1574. (12) Douglas, J. M. Conceptual Design of Chemical Processes; McGraw-Hill: New York, 1988. (13) Sun, L.; Wang, Q.; Li, L.; Zhai, J.; Liu, Y. Design and control of extractive dividing wall column for separating benzene/cyclohexane mixtures. Ind. Eng. Chem. Res. 2014, 53, 8120-8131. (14) Wu, Y. C.; Hsu, P. H. C.; Chien, I. L. Critical assessment of the energy-saving potential of an extractive dividing-wall column. Ind. Eng. Chem. Res. 2013, 52, 5384-5399. (15) Tang, Y. T.; Chen, Y. W.; Huang, H. P.; Yu, C. C.; Hung, S. B.; Lee, M. J. Design of reactive distillations for acetic acid esterification. AIChE J. 2005, 51, 1683-1699. (16) Turton, R.; Bailie, R. C.; Whiting, W. B.; Shaeiwitz, J. A. Analysis, Synthesis and Design of Chemical Processes, 4th ed.; Prentice Hall: Upper Saddle River, New Jersey, 2012.

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(17) Luo, H.; Bildea, C. S.; Kiss, A. A. Novel heat-pump-assisted extractive distillation for bioethanol purification. Ind. Eng. Chem. Res. 2015, 54, 2208-2213. (18) Ognisty, T. P. Analyze distillation columns with thermodynamics. Chem. Eng. Prog. 1995, 91, 40-46. (19) Olujic, Z.; Fakhri, F.; De Rijke, A.; De Graauw, J.; Jansens, P. J. Internal heat integration-the key to an energy-conserving distillation column. J. Chem. Technol. Biotechnol. 2003, 78, 241-248. (20) Luyben, W. L. Distillation Design and Control Using Aspen Simulation, 2nd ed.; Wiley: 2013. (21) Luyben, W. L. Plantwide Dynamic Simulators in Chemical Processing and Control; Marcel Dekker: New York, 2002.

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Design and Control of Different Pressure Thermally Coupled Reactive

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