Design of an Intensified Reactive Distillation Configuration for 2

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Design of an Intensified Reactive Distillation Configuration for 2-Methoxy-2-Methylheptane Arif Hussain, Le Quang Minh, Muhammad Abdul Qyyum, and Moonyong Lee Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.7b04624 • Publication Date (Web): 12 Dec 2017 Downloaded from http://pubs.acs.org on December 14, 2017

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Industrial & Engineering Chemistry Research

Design of an Intensified Reactive Distillation Configuration for 2-Methoxy-2Methylheptane Arif Hussaina, Le Quang Minha, Muhammad Abdul Qyyuma, Moonyong Leea* a

School of Chemical Engineering, Yeungnam University, Gyeongsan 712-749, Rep. of Korea.

Submitted to: Journal of Industrial & Engineering Chemistry Research

*

Correspondence concerning this article should be addressed to:

Prof. Moonyong Lee Email: [email protected] Telephone: +82-53-810-2512

ACS Paragon Plus Environment

Industrial & Engineering Chemistry Research 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

Abstract The design of an intensified reactive distillation (RD) configuration for the production of 2-methoxy-2methylheptane (MMH) was proposed. Conventionally, high yield of MMH can be achieved by a largesized reactor or by recycling 2-methyl-1-heptene (MH). Both strategies are associated with high capital and operating costs. This study proposed an innovative RD configuration that allows the production of MMH with significantly lower energy costs, and total annual costs (TACs). The results demonstrate that the proposed RD configuration can reduce energy demands, and total annual costs by up to 18.2, and 10.8 %, respectively, compared to those for a conventional reactor column sequence for the desired MMH yield. In addition, different heat integration sequences were investigated to utilize the sensible heat to save energy and the subsequent TAC. One heat integration sequence for the proposed configuration further realized savings in energy, and TAC by up to 36.7, and 21.8 %, respectively.

Keywords: Reactive distillation; 2-methoxy-2-methylheptane; reactive holdup; capital and energy savings.


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1. Introduction The development of the lead-containing compounds added to gasoline to minimize knocking has gained tremendous importance 1. Methyl tert-butyl ether (MTBE) has been considered an important “oxygenate” to raise the oxygen content of gasoline. However, restrictive energy and environmental policies have banned the use of MTBE as a fuel additive in gasoline due to its solubility in water and its widespread detection in groundwater 2. This led to increased interests in higher-molecular-weight ethers that exhibit significantly decreased solubility. One viable alternative is the chemical 2-methoxy-2-methylheptane (MMH), produced through an etherification reaction between 2-methyl-1-heptene (MH) and methanol (MeOH) 3. However, the main barrier to this approach is the high investment and operating costs associated with the conventional process that uses a reactor followed by three distillation columns in sequence. High yield of the desired product MMH could be achieved by using a large-sized reactor or an excess of MH. Both strategies are associated with high capital and operating costs. For a specified yield of MMH, Luyben 3 studied the optimum values of reactor vessel sizes and recycle flow rates to minimize the total annual cost. Furthermore, Griffin et al.

4

studied the effect of high and low values of an

equilibrium constant for the optimum operating conditions for the MMH plant. By designing innovative equipment and implementing more energy-efficient production processes, practical solutions can be found to overcome the constraints imposed on various raw materials and products. Process intensification is a novel design strategy that aims to improve existing processes by reducing equipment size, energy consumption, and overall plant footprint

5-6

. RD falls into the category of process intensification because

of its potential to integrate a conventional reactor column sequence into a single shell to improve the performance of reactive processes. Currently, more than 150 RD columns with annual capacities of 100– 3000 kt are in operation worldwide 7. A huge scholarly database dealing with the modelling, simulation, and industrial applications of RD can be found elsewhere

8-11

. Industrial RD columns for etherification

reactions, especially for the MTBE process, have seen tremendous growth across the world over the last few decades, with improved capital and operating savings. The MTBE process is, in fact, the largest application of RD in terms of the number of columns and production capacity


12

, but the future is


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uncertain due to the legislation that prohibits its use. The growing interest in the chemical MMH to replace MTBE necessitates an increase in the capacity in a cost-effective way to meet the increasingly growing market demand. Heat integration and intensification is another promising practical technique for minimizing the energy consumption of distillation processes. The latent and sensible heat of the process stream is circulated into the processes for exergy destruction. Several studies deal with the implementation of heat-integration in reactive distillation systems for enhancing energy efficiency 6, 13-16. This paper proposes an innovative RD configuration that allows the production of MMH with significantly lower capital and operating costs. During the design, the temperature limitation of an acid resin catalyst was overcome by selecting the appropriate reactive zone, reflux ratio, catalyst holdup, and the suitable column pressure. This study further emphasized the potential benefits of heat integration for the proposed RD column configuration to replace the conventional energy-intensive MMH production process.

2. Process Studied The proposed study is based on the conventional flowsheet for making MMH taken from the works of Luyben

3

and Griffin et al.

4

to verify the feasibility of RD configuration. The reaction mechanism

involves the desired reversible reaction (Equation 1) between methanol (MeOH, CH3OH) and 2-methyl1-heptene (MH, C8H16) to form 2-methoxy-2-methylheptane (MMH, C9H20O). The same reactants produce dimethyl ether (DME, C2H6O) and 2-methyl-2-heptanol (MHOH, C8H18O) as byproducts in an undesirable irreversible reaction (Equation 2). CH3OH + C8H16 ↔ C9H20O 2CH3OH + C8H16 → C2H6O + C8H18O

(1) (2)

The corresponding rate equations in terms of the mole fraction for the desired and undesired reactions are represented by Equations 3 and 4, respectively:


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RD = kDfXMeOHXMH - kDrXMMH

(3)

RU = KU(XMeOH)2

(4)

where kDf and kDr are the forward and backward rate constants for the desired reaction. KU is the rate constant for the undesired reaction. Table 1 list the values of reaction rate constants and the activation energies for both reactions 4. Table 1 Kinetic parameters. Adapted with permission from (Luyben, W. L. Design and Control of the Methoxy-Methyl-Heptane Process. Ind. Eng. Chem. Res. 2010, 49, 6164-6175). Copyright (2010), American Chemical Society.

kDf kDr kU

= = =

kmol s-1kgcat-1 6.7 x 107 2.1 x 10-6 1.3 x 109

EDf EDr EU

= = =

kJ kmol-1 90000 900 105900

2.1. Catalyst The industrial grade Amberlyst 35Wet (Rohm & Haas) catalyst, which is macroreticular, strongly acidic, cationic, polymeric catalyst was used to report the kinetic data for this etherification reaction 17. As per Dow chemical product data sheet 18, the maximum operating temperature and the density were reported as 150°C and 800 g/L, respectively. 2.2. Conventional Process The simplified flowsheet for the conventional process of producing MMH is illustrated in Figure 1. The reaction between MeOH and MH is carried out in a fixed bed catalytic reactor. The reactor effluent containing DME, MMH, MHOH, and unreacted MeOH and MH is fed to column C1. Column C1 separates the DME with mole purity of up to 99.9 %. The bottom stream originating from C1 column is fed to column C2. The unreacted reactants are separated from MMH in column C2 and are recycled back to the reactor from the distillate. Column C3 further purifies the desire product MMH and undesired MHOH up to a mole fraction of 99.9 % as the distillate and bottom product, respectively.



Figure 1 Conventional flowsheet for producing MMH. Adapted with permission from (Luyben, W. L. Design and Control of the Methoxy-Methyl-Heptane Process. Ind. Eng. Chem. Res. 2010, 49, 6164-6175). Copyright (2010), American Chemical Society. 2.3. Feasibility of RD for 2-methoxy-2-methylheptane (MMH) An article by Griffin et al.

4

suggested that the reaction mechanism for MMH production is a possible

candidate for a reactive distillation configuration, which can offer significant advantages. However, to assure the feasibility of reactive distillation, the methodology adopted is based on an article by Shah. et.al 19

. For preliminary design, they provide guidelines and recommended values for (e.g. reaction/separation

temperature difference, chemical equilibrium constant Keq, rate constant for main reaction, side reaction temperature, and the production capacity) to verify the economic feasibility of RD process. The step by step procedure to perform feasibility analysis is shown in Figure 2. The process chemistry of MMH involve reversible desired reaction that exhibit the problem of limiting chemical equilibrium conversion. In addition, the main reaction is associated with an undesired irreversible reaction that produces byproducts. Reactive distillation is a wise choice as it has ability to overcome the equilibrium limitations and enhance selectivity when there are byproducts formation. Another criterion for economic feasibility of reactive distillation is the temperature compatibility between the reaction [TR-main] and the distillation [TD] operation. This temperature difference [TD-TR-main] should not exceed 50°C. In the conventional process of MMH, the reaction occurs at 400 K due to catalyst activity, whereas the temperature of MMH distillation column is 426 K. As the temperature difference between the separation and reaction is lower than 50°C, conducting this reaction in a single unit can offer significant advantages. Chemical equilibrium


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constant (Keq) has significant effects in the design of RD column. For conventional RD processes, the recommended value of Keq should be greater than 1. The equilibrium constant for the desired reaction can be determined by kf/kr. Based on Table 1, the equilibrium constant (Keq, at 423 K) gives a value of 316, which confirms that the MMH system has the most favorable reaction kinetics to be implemented for RD column design. The specific reaction rate for the main reaction should be higher than 10-5 kmol/kg.sec to avoid large liquid holdups. For MMH reaction, the specific reaction rate for main reaction is significantly higher which confirms the economic feasibility of reactive distillation. For RD processes, the selectivity towards desired product could be reduced due to byproduct formation. Therefore, the operating temperature of RD column [TRD] should be lower than that of side reactions [TS]. For MMH process, the temperature of the side reaction is significantly higher than the RD operating temperature. Furthermore, the RD is also attractive in terms of gross profit as the MMH production rate is 62 kt/year which is higher than the range (0.5-1.0 kt/year) where RD is economical feasible.

Figure 2 Feasibility flow diagram of MMH process for RD application.



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3. Phase Equilibrium The base case by Luyben 3 and Griffin et al.

4

provided the sketchy information about the conventional

flowsheet of 2-methoxy-2-methylheptane. In addition, their flowsheet provides the reaction kinetics and the desired production rate. The intent of this paper is to present a preliminary results of a reasonable conceptual design based on reactive distillation configuration. However, the feasibility of RD requires valid vapor-liquid equilibrium data. This data is essential in order to obtain the Residual curve maps (RCM). Since, there is no VLE data reported in the literature, therefore, all the unit operations were simulated using Aspen Plus® built-in “UNIQUAC” physical property method, as suggested by Luyben

3

in the conventional process. The built-in UNIQUAC thermodynamic model only provide binary interaction parameters between MeOH and DME. For all other separation, this study assumes ideal vaporliquid equilibrium behavior to propose the feasibility of RD configuration. The components MMH and MHOH were generated using the molecular structures (Figure 3) because they were not available in the Aspen databank. The result for the base case was retrieved to confirm the phase equilibrium diagrams for separation. Figures 4a and 4b show the T-xy diagrams for separating DME/MeOH and MMH/MHOH used in this simulation, respectively.

Figure 3 Molecular structures for MMH and MHOH. Adapted with permission from (Luyben, W. L. Design and Control of the Methoxy-Methyl-Heptane Process. Ind. Eng. Chem. Res. 2010, 49, 6164-6175). Copyright (2010), American Chemical Society.


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550

410

A 500 450 400 350 T-x 10 atm T-y 10 atm

B

400

Temperature (K)

Temperature (K)

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


390 380 370 360 350

300

T-x 0.1 atm T-y 0.1 atm

340 0.0

0.2

0.4

0.6

0.8

1.0

Mole Fraction DME

0.0

0.2

0.4

0.6

0.8

1.0

Mole Fraction MMH

Figure 4 VLE diagram (A) DME/MeOH at 10 atm and (B) MMH/MHOH at 0.1 atm. 4. Proposed Configuration 4.1. RD Column The newly proposed configuration based on the RD column is shown in Figure 5. The RD column has 35 total stages. There is no rectifying section; the reactive zone is between stages 2 and 12, followed by the stripping section at the bottom. The stages are numbered according to the Aspen Plus stage notation with stage 1 and 35 being the condenser and reboiler, respectively. The fresh MeOH stream as a low-boilingpoint reactant is fed at stage 13, while the fresh MH stream as a heavier reactant combined with a recycle stream of a total of 128.98 kmol/h is introduced at stage 1. The column is operated at 1.77 atm. The high MH/MeOH ratio makes this process a suitable candidate for RD implementation. A total estimated load of 660 kg of acid resin catalyst per stage was used within the reactive zone. The column diameter from liquid holdup calculation was estimated as 3.90 m. 4.2. Column C2 The distillate from the RD column is fed to stage 5 of a 12-stage C2 column. Column C2 serves the purpose of delivering high-purity DME as the distillate product as a green energy source, while the unreacted MeOH and MH are recycled at the bottom. The normal boiling point of DME is 248.2 K. The operating pressure of column C2 was set at 10 atm, a value that permits the use of cooling water in the



condenser. The small distillate stream of 0.518 kmol/h with 99.9 mol % DME gave a condenser duty of 0.0045 MW. The bottom stream contained the unreacted MeOH and MH, which were recycled to the RD column. The reboiler duty was 0.72 MW with a base temperature of 501.9 K, which used high-pressure steam. The column diameter from Aspen tray sizing was 0.83 m. 4.3. Column C3 Column C3 received a feed stream from the bottom of the RD column. This column had a total of 15 stages with stage 8 being the feed location. The distillate product was 49.02 kmol/h with 99.9 mol % MMH. The bottom was 0.49 kmol/h with 99.9 % of undesired MHOH. This column operated under 0.1 atm pressure, giving 0.64 and 0.18 MW condenser and reboiler duties, respectively. The reflux drum temperature was 352 K and the temperature at the bottom was 401 K, accompanied by low-pressure steam. The column diameter was estimated as 1.42 m.

Figure 5 Proposed RD configuration for the MMH process.


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5. Sensitivity Analysis 5.1. Optimization of MeOH/Overall MH ratio The reaction chemistry for MMH process also involve an undesirable irreversible reaction that produces DME and MHOH. To achieve a higher selectivity towards the desired product MMH, the methanol concentration must be kept low in the RD column, which requires a large recycle of MH from the downstream separation section. The MeOH/overall MH feed ratio is optimized for different MH recycle by using flowsheet design specification tool in Aspen plus. The fresh methanol feed was fixed at 50 kmol/h. The results demonstrate that small total MH recycle requires more catalyst weight per tray to achieve design specification of 99 % MMH yield and vice versa. The corresponding energy costs also reduces as the recycle flow rate decreases. The catalyst cost and the corresponding column diameter however dominates the capital investment with small recycle flowrates. Table 2 summarizes the alternative design to obtained 99 % MMH yield for different combination of required catalyst weight and total MH recycle. The design that gives minimum TAC is associated with 7260 kg total catalyst with a total MH flow rate of 129.5 kmol/h. Table 2 Alternative designs to obtain 99 % MMH yield. Alternative design to obtained 99% yield Total MH Recycle (kmol/h) 135 6

129.9

125.5

Catalyst Cost (10 $) @ 10$/kg C1/RD Column H(m) ID (m) QR(kW) QC(kW)

0.0572

0.0726

0.132

24.16 3.48 1521 3144

24.16 3.90 1340 2917

24.16 5.28 1209 2752

Column Vessel Cost (106$)

0.857

0.968

1.337

0.200

0.187

0.177

0.429

0.379

0.343

7.32 0.85 770 4.29

7.32 0.83 725 4.51

7.32 0.81 691 4.51

6

Exchanger Capital Costs (10 $) 6

Energy Costs (10 $/year) C2 Column H(m) ID (m) QR(kW) QC(kW)



Column Vessel Cost (106$) 6


Energy Costs (10 $/year) C3 Column H(m) ID (m) QR(kW) QC(kW) Column Vessel Cost (106$) 6


Energy Costs (10 $/year)

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0.073

0.071

0.070

0.084

0.081

0.079

0.240

0.225

0.215

9.52 1.40 166 628

9.52 1.42 186 647

9.52 1.41 178 640

0.154

0.156

0.155

0.076

0.079

0.078

0.050

0.055

0.053

6

1.501

1.615

2.027

6

0.718

0.660

0.611

1.218

1.198

1.287

Total Capital Cost (10 $) Total Energy Cost (10 $/year) 6

TAC (10 $/year)

5.2. Effect of Catalyst Holdup Stage holdup has a pronounced effect in the design of a reactive distillation column. The extent of reaction at the reactive stages was examined by the Damkohler number (Da), which is the ratio of the characteristic residence time to the characteristic reaction time. The methanol feed temperature (393 K) was chosen as the reference temperature for calculating the value of the forward reaction rate constant (kf, ref),

leading to a value of 0.264 kmol h-1 kgcat-1. The catalyst weight was varied inside the reactive stages

M k , F to obtain different values of the Da using Equation 5. D =

(5)

where Mcat is the mass of catalyst and F is the total feed flow rate to the RD column. The total feed flow rate to the RD column was fixed at 178.98 kmol/h. A quick estimation of the Damkohler number could determine whether the reaction rate was slow or fast in the RD column. For low values of Da (Da ≤ 0.1), the system is dominated by phase equilibrium, whereas high values of Da (Da ≥ 10) correspond to a fast reaction rate where chemical equilibrium is approached at the reactive stages. If the Da does not lie between these values, then the process is in fact kinetically controlled 20.


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The methanol conversion and MMH yield as a function of the Damkohler number are illustrated in Figure 6a. The large value of Damkohler number corresponding to the large catalyst holdup showed better improvement in the reaction conversion and yield profile. Up to 660 kg of the catalyst, the temperature of the last reactive stage was within the permissible catalyst deactivation temperature (423 K). Addition of more catalyst, however, led to the catalyst deactivation temperature being exceeded, as shown in Figure 6b. Therefore, the optimum value of the Damkohler number was considered to be 0.97, associated with 7260 kg of total catalyst, for achieving 99 % MMH yield. The impact of the catalyst holdup on process intensification in the RD column is illustrated in Figure 6c, where the reboiler and condenser heat duties first decrease and then begin to increase with an increase in the Damkohler number.

0.994

A 1.000

Yield Conversion

0.992

0.999

Yield (MMH)

0.990

0.998

0.988 0.997 0.986 0.996

0.984

0.995

0.982

0.994

0.980 0.978 0.4

0.6

0.8

1.0

0.993 1.4

1.2

Damkohler number

425

B Stage 12 temperature (K)

424 423 422 421 420 0.5

0.6

0.7

0.8

0.9

1.0

1.1

Damkohler number


1.2

1.3

Conversion (MeOH)

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



2.965

1.375 1.370

Reboiler duty (MW)

C

Reboiler duty Condenser duty

2.960 2.955

1.365

2.950

1.360

2.945 2.940

1.355

2.935 1.350

2.930

1.345

2.925

1.340

2.920

Condenser duty (MW)

2.915

1.335 0.4

0.6

0.8

1.0

1.2

1.4

Damkohler number

Figure 6 (A) Conversion and yield profile, (B) last stage temperature, and (C) reboiler and condenser duties as a function of the Damkohler number. The philosophy behind understanding the impact of catalyst holdup in the design of the RD column was further analyzed through the net reaction rate profile, which has been plotted as a function of the Damkohler number in Figure 7. The high peaks showing a shift toward the bottom on increasing the Damkohler number illustrate that the reaction mainly took place at the bottom of the reactive zone, which is favorable in terms of internal heat integration due to heat of reaction between the reaction and separation processes 21. The fact is further clarified on considering the corresponding heat duty profiles for the reboiler and condenser of the RD column (Figure 6c).

22

Da = 0.53 Da = 0.71 Da = 0.89 Da = 0.97

20 18

Net reaction rate (kmol/h)

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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16 14 12 10 8 6 4 2 0 0

2

4

6

8

10

12

14

16

RD column stages

Figure 7 Net reaction rate profile at different Damkohler numbers.


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5.3. RD Column Sizing Due to the hydraulic limitations, the reactive holdup cannot be more than the system liquid holdup, otherwise, the catalyst will fill full in the stage. The following assumptions were made to size the column diameter: (1) A reasonable liquid depth of 0.152 meters (6 inches) was assumed, (2) the catalyst Amberlyst 35 occupies half of a reactive tray, (3) the downcomer occupies 10 % of the tray area 16. With these assumptions, the diameter of the column was then determined using equation 6 to accommodate this catalyst per tray. The column diameter is estimated as 3.90 meter using a bulk density of 800 kg/m3.

4

= 0.152

(6)

6. Effect of Number of Reactive Trays The effect of the number of reactive trays on the various design parameters was analyzed keeping in view the upper limit of the temperature of the catalyst. As seen in Figure 8a, a yield of 99 % towards the desired product (MMH) could be achieved at 11 reactive trays. These reactive trays also correspond to the lower reboiler and condenser duties, as illustrated in Figures 8b and 8c. Furthermore, for the 11 reactive stages, the temperature of the last stage was within the permissible limit where catalyst deactivation would be avoided, as shown in Figure 8d. Adding more trays will result in either decrease in the required yield or increase in the heat duties and the upper limit of the catalyst temperature.



1.347

A

99% MMH yield

48.96

B Reboiler duty (MW)

MMH produced (kmol/h)

48.98

48.94 48.92 48.90 48.88 48.86

1.346 1.345 1.344 1.343 1.342

48.84 2.926

C

428

Stage 12 temperature (K)

2.925

Condenser duty (MW)

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 16 of 35

2.924 2.923 2.922 2.921 2.920 2.919 8

9

10

11

12

13

14

D

426 424 422 420 418 8

9

Number of reactive trays

10

11

12

13

14

Number of reactive trays

Figure 8 (A) Net MMH produced, (B) RD reboiler duty, (C) RD condenser duty, and (D) last reactive stage temperature as a function of the number of reactive trays. 7. Optimization of Column Trays The total number of trays for the RD column, column C2, and column C3 were optimized for the minimum TAC. Adding more trays would reduce the column diameter and heat exchanger area; however, it would also increase the height of the column, which would ultimately increase the column capital investment. Table 3 lists the effects of the column trays on these competing variables. Moreover, for these columns, the optimum feed locations were chosen to give the minimum reboiler duties. The optimum numbers of trays for the RD column, column C2, and column C3 were found to be 35, 12, and 15, respectively. Table 3 Optimization of column trays. NT NFMHopt

30 1

RD-Column 35 40 1 1

10 -

C2-column 12 15 -


10 -

C3-column 15 22 -

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NFMeOHopt NFopt ID(m) QR (MW) QC (MW) Shell Heat Exchanger Tot. Capital Cost (106$) Tot. Energy Cost (106 $/y) TAC (106 $/y)

13

13

13

3.61 1.478 3.055 0.317 0.196 0.513 0.383 0.554

3.60 1.340 2.917 0.350 0.187 0.537 0.347 0.526

3.58 1.313 3.235 0.389 0.205 0.593 0.340 0.538

4 0.93 0.791 0.005 0.067 0.086 0.153 0.246 0.297

5 0.83 0.724 0.004 0.071 0.081 0.153 0.225 0.276

6 0.8 0.718 0.001 0.085 0.080 0.165 0.223 0.278

4 1.6 0.352 0.814 0.120 0.103 0.223 0.091 0.166

8 1.42 0.185 0.647 0.156 0.079 0.235 0.048 0.126

13 1.39 0.173 0.634 0.216 0.077 0.293 0.045 0.142

8. Impact of Operating Pressure 8.1. Pressure RD column According to the calculated values of the Damkohler number, the system was kinetically controlled. When a RD column involves kinetically controlled reactions, the operating pressure thus has a significant influence on the RD column performance parameters such as the conversion, yield, and net reaction rate profiles 21. A higher operating pressure results in better process intensification in terms of internal heat integration due to reaction heat involved. However, the operating pressure should be selected carefully to achieve the maximum benefits from process intensification and to avoid elevated temperatures in the reaction zone that could lead to catalyst deactivation. The net reaction rate profile with different operating pressures is shown in Figure 9a. The high peaks move further to the bottom on increasing the column operating pressure. This is because reaction mainly occurs at the bottom of the reactive zone, which favors process intensification between the reaction and separation involved due to exothermic heat of reaction. The corresponding heat duties for the reboiler and condenser decreased, as shown in Figure 9b. Figure 9c illustrates the effect of the column pressure on the conversion and yield of MMH. The high operating pressure resulted in a high temperature in the reactive zone, which favored high conversion because of the activation energy of forward reaction being larger than that of the backward reaction. The desired yield of MMH was obtained at the operating pressure of 1.77 atm. Moreover, the temperature profile was within the permissible limit (423 K) within which catalyst deactivation could be avoided. Further increase in the operating pressure results in a higher temperature not suitable for the MMH



process. Figure 9d shows the temperature of the last stage of the reactive zone as a function of the RD column operating pressure. The best option was to fix the operating pressure at 1.77 atm to ensure a suitable upper limit temperature for the reactive zone. However, the operating pressure of 2 atm was optimum in terms of internal heat integration.

Net reaction rate (kmol/h)

A

1.55-atm 1.77-atm 2.0-atm

24 21 18 15 12 9 6 3 0 2

4

6

8

10

12

14

16

18

RD column stages

6.0 Condenser Duty Reboiler Duty

5.5

B

5.0

Heat Duties (MW)

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

4.5 4.0 3.5 3.0 2.5 2.0 1.5 1.0 0.4

0.6

0.8

1.0

1.2

1.4

1.6

1.8

RD column pressure (atm)


2.0

2.2

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Page 19 of 35

1.0

Yield Conversion

Converision and Yield

0.9

C

0.8 0.7 0.6 0.5 0.4 0.3 0.2 0.4

0.6

0.8

1.0

1.2

1.4

1.6

1.8

2.0

2.2


440

D

Stage 12 Temp

430 420

Temperature (K)

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


410 400 390 380 370 360 0.4

0.6

0.8

1.0

1.2

1.4

1.6

1.8

2.0

2.2


Figure 9 RD column pressure as a function of (A) net reaction rate, (B) heat duties, (C) conversion and yield, and (D) stage 12 temperature. 8.2. Pressure C2 and C3 column. To assess the economic impact of pressure, different operating pressures for C2 and C3 column are chosen for minimum TAC. Table 4 gives detailed results for both columns over a range of pressures. For both columns energy costs increases as pressure is increased. In addition, columns diameter decreases as pressure increases. The operating pressure of column C2 is set at 10-atm, which gives a reflux-drum temperature of 318 K and permits the use of cooling water. The base temperature of C2 column is high (501 K), which uses high pressure steam. There is a possibility to operate C2 column at lower pressure that might be economical since the condenser duty is very small compared to the reboiler duty. This alternative design was briefly explored for lower pressures. For operating pressure of 6-atm, the reflux



Page 20 of 35

drum temperature was 299 K which require 267 K refrigerant at $7.89/GJ. The cost of 5.616 kW of refrigeration is only 1396 $/year. The bottom temperature decreases from 501 to 472 K, which still require high-pressure steam. The savings in energy cost is about 79415 $/year as reboiler duty decreases from 0.724 to 0.469 MW. These alternative designs were not considered since this paper is comparative study with the base case configuration where DME column operates at 10-atm as well. Furthermore, by operating the C2 column at 10-atm, the bottom stream at 501 K can be effectively utilized for possible heat integration sequences. The resulting savings from this stream in a heat integration sequence (developed in later section) gives savings of about 143009 $/year in terms of energy cost. For pressure in C3 column, the selection criteria is based on an article by Luyben

22

, where the minimum practical

operating pressure is assumed to be 0.1-atm. Since, the normal boiling points of MMH and MHOH are higher than the water-cooled condenser (318 K), this means that column could be operated under vacuum if it is economical attractive. Therefore, the operating pressure selected for C3 column is 0.1-atm which gives minimum TAC. Table 4 Pressure selection for C2 and C3 column. Pressure (atm) QR (kW) QC(kW) TR (K) TC (K) Steam Diameter (m) Capital cost (106$) Energy cost (106 $/y) TAC (106 $/y)

4 306.1 7.534 450.9 285.6 HP 1.21 0.152 0.097 0.147

Column C2 6 469.4 5.616 472 299.5 HP 1.12 0.120 0.147 0.187

10 724.8 4.507 501.9 318.1 HP 0.83 0.153 0.225 0.276

0.1 185.4 647.2 401.6 352.1 MP 1.42 0.235 0.055 0.134

Column C3 0.3 349.7 673.9 432.1 383.3 MP 1.07 0.197 0.098 0.164

0.6 576.8 792.9 453.8 406 HP 1 0.205 0.188 0.257

9. Selection of Reflux Ratio The influence of the reflux ratio on the performance of the RD column was investigated further. The relationship between the conversion of MeOH and MH with the reflux ratio is illustrated in Figure 10a. A slight decrease was observed for MeOH conversion, while the MH conversion increased, when the reflux ratio increased. The desired MMH yield of 99 % was achieved at a reflux ratio of 3.60, where the


Page 21 of 35

conversion of MeOH was 99.9 %, as illustrated in Figure 10b. The corresponding reboiler duty as a function of the molar reflux ratio is illustrated in Figure 10c. Therefore, a reflux ratio of 3.60 was selected for the RD column for obtaining 99 % MMH yield.

0.385

A

0.9996

0.380

MeOH conversion

0.9995 0.9994

0.375

0.9993 0.370

0.9992 0.9991

0.365

MH conversion

0.9990 X(MeOH) X(MH)

0.9989 1.8

2.1

2.4

2.7

3.0

3.3

3.6

3.9

0.360 4.2

Molar reflux ratio

1.00

B

Yield specification 99%

0.98

MMH Yield

0.96 0.94 0.92 0.90 0.88 0.86 1.8

2.1

2.4

2.7

3.0

3.3

3.6

3.9

4.2

Molar reflux ratio

1.8

C Reboiler duty RD column (MW)

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


1.6 Yield specification 99%

1.4 1.2 1.0 0.8 0.6 0.4 0.2 1.8

2.1

2.4

2.7

3.0

3.3

3.6

3.9

4.2

Molar reflux ratio

Figure 10 (A) Conversion, (B) yield, and (C) RD reboiler duty, as a function of the reflux ratio.



10. Heat Integration Sequences Column C2, which operated at 10 atm, had a base temperature of 501.9 K. Different heat integration sequences were investigated by circulating the hot stream to various sections of the proposed RD configuration to further reduce the energy consumption. First, this hot temperature stream was circulated to provide boil-up energy to the bottom section of the RD column and the C3 column for potential energy reduction. Furthermore, it can be used to boil up the bottom section of column C3 while the feed preheating of C2 and C3 columns can be exploited further. Because this process stream contains unreacted MeOH and MH that should be recycled after possible heat integration sequences to achieve the desired conversion and yield inside the RD column. 10.1. Heat Integration Sequence 1 In the first heat integration sequence, the bottom of the RD and C3 column were subjected to boiling using the sensible heat from the bottom of C2 column. In the simulation, a refluxed absorber was selected for the RD and C3 column. The bottom streams originating from C2 was firstly introduced to the hot side while the vapor boilup stream from RD column was introduced to the cold side of the shell and tube heat exchanger HEX-1. Note that the minimum approach was set at 10 K and the overall heat-transfer coefficient U was assumed by 0.0203 cal/sec-m2-K from Luyben 3 for all heat exchangers considered in this study. The heat exchanger area was required to be maintained at 16.63 m2 so 0.36 MW of heat can be transferred into the cold liquid stream. The reboiler duty could be reduced from 1.34 MW to 0.98 MW using an auxiliary trim reboiler to fully vaporize the cold vapor boilup stream. In sequence, the outlet of the hot stream originating from HEX-1 was passed through the hot side of another heat exchanger HEX-2 to provide boilup to the bottom of the C3 column. The heat exchanger area for HEX-2 was required at 5.12 m2 for providing 0.18 MW of heat to fully vaporized the cold vapor boilup stream through sensible heat. The diameter of column C3 was reduced to 1.4 m because the internal vapor flow decreased as a result of heat integration. The hot outlet stream at 433 K from HEX-2, which contained unreacted MeOH and MH, was recycled back to the RD column. The resulting heat integration sequence 1 gives savings of


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149050 $/year in terms of operating cost. Figure 11 illustrates the schematic of the proposed heat integration sequence 1.

Figure 11 Heat integration sequence (1) for the proposed configuration. 10.2. Heat Integration Sequence 2 In this heat integration, the bottom of the RD column was subjected to boiling using the sensible heat of the bottom stream from column C2. In the simulation, a refluxed absorber was selected for the RD column. The bottom stream originating from the RD column was split into the column C3 feed and vapor boilup streams. Meanwhile, the bottom stream originating from column C2 was introduced to the hot side while the vapor boilup stream originating from the RD column was introduced to the cold side of the shell and tube zones of heat exchanger HEX-1. The heat exchanger area was required to be maintained at 16.54 m2. The reboiler duty could be reduced from 1.34 MW to 0.98 MW. Because the heat duty of column C2 cannot replace that of the RD column, an auxiliary trim reboiler was needed to fully vaporize the cold vapor boilup stream. The boilup stream was then introduced to the bottom of the RD column. In sequence, the outlet of the hot stream originating from HEX-1 was passed through another heat exchanger HEX-2 to preheat the feed stream of column C2. There is a maximum temperature constraint of 413 K for



preheating the feed stream of C2 column. Increasing the temperature further will result in lower DME purity which needed a large reflux ratio to meet the design specification. Therefore, in all heat integration sequences the maximum temperature was fixed at 413 K where feed preheating of C2 column was done. As a result, this sensible heat required a heat transfer area of 0.82 m2 to increase the feed temperature from 407 K to 413 K. The reboiler duty was reduced from 0.72 MW to 0.692 MW. The diameter of column C2 and C3 was reduced to 0.78 m and 1.40 m respectively, because the internal vapor flow decreased as a result of heat integration. The hot outlet stream at 453 K from HEX-2, which contained unreacted MeOH and MH, was recycled back to the RD column. The resulting heat integration sequence 2 gives savings of 112182 $/year in terms of operating cost. Figure 12 illustrates the schematic of the proposed heat integration sequence 2.

Figure 12 Heat integration sequence (2) for proposed configuration. 10.3. Heat Integration Sequence 3 In this heat integration, the bottom stream of column C2 was circulated to provide the sensible heat for vapor boil-up of column C3. The bottom stream originating from column C3 was first split into the MHOH product and vapor boilup streams. As illustrated in Figure 13, the bottom stream originating from


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column C2 was introduced to the hot side while the vapor boilup stream from column C3 was introduced to the cold side of the shell and tube heat zones of exchanger HEX-1. The hot stream from column C2 fully vaporized the cold vapor boilup stream through sensible heat giving a heat exchanger area of 2.42 m2. The hot outlet stream was circulated through another heat exchanger HEX-2 to provide sensible heat to preheat the feed stream of column C2, which gave a heat transfer area of 0.49 m2. The diameter of column C2 and C3 was reduced to 0.79 m and 1.39 m respectively, because the internal vapor flow decreased as a result of heat integration. The hot outlet stream at 477 K from HEX-2, which contained unreacted MeOH and MH, was recycled back to the RD Column. The resulting heat integration sequence 3 gives savings of 59324 $/year in terms of operating cost. Figure 12 shows the schematic of the proposed heat integration sequence 3.

Figure 13 Heat integration sequence (3) for proposed configuration. 10.4. Heat Integration Sequence 4 This heat integration sequence proposed feed preheating of the columns C2 and C3 using the process stream originating from the bottom of column C2 as the heat source. The heat integration sequence 4 is shown in Figure 14. First, the hot stream from column C2 was used as a heat source in the heat exchanger HEX-1 to preheat the feed stream of column C3. As a result, the reboiler duty was reduced to 0.013 MW,



which gave a heat transfer area of 4.72 m2. The hot outlet stream from HEX-1 was further introduced to another heat exchanger HEX-2 to provide sensible heat to preheat the column C2 feed stream. The reboiler duty decreased to 0.69 MW, while the required heat exchanger area was 0.48 m2. The hot outlet stream at 478 K from HEX-2, which contained unreacted MeOH and MH, was recycled back to the RD Column. In the simulation, the minimum approach temperature was set at 10 K, while an overall heattransfer coefficient U = 0.0203 cal/sec-m2-K was used for heat exchanger HEX-1 and HEX-2 3. The diameter of column C2 and C3 was reduced to 0.79 m and 1.40 m respectively, because the internal vapor flow decreased as a result of heat integration. The resulting heat integration sequence 4 gives savings of 55912 $/year in terms of operating cost.

Figure 14 Heat integration sequence (4) for proposed configuration. 11. Design Economics For economic evaluation, the capital costs (pertaining to the column shell, heat exchangers, and catalyst) and operating costs (steam in the reboilers, and condenser utilities) were taken into consideration. An economic assessment of the proposed configuration for the desired yield of MMH was carried out in terms of the total annual cost (TAC), with the aid of Equation 7 taken from Kaymak and Luyben 23:


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TAC =

Capital cost + Operating Cost Payback period

(7)

The basis for economics and equipment sizing are summarized in Table 5, taken from the report by Douglas 24 and Turton et al. 25. The payback period of three years and energy costs for low pressure (LP), medium pressure (MP), and the high pressure (HP) steams were taken to be similar as those in the base case for fair comparison. Table 5 Basis for economic and equipment sizing. Column Shell Cost Column Diameter (m) Column Length (m) Heat Exchanger Capital costs Condenser Capital Cost ($) Heat Transfer Coefficient Differential Temperature (K) Reboiler Capital Cost ($) Heat Transfer Coefficient Differential Temperature (K) Energy Cost ($/year)

Cooling Water ($/year) Catalyst Cost ($)

17640 ∗ D7.899 ∗ L8.;8
? 8.9@ , A>? =

U>? = 0.852 kW/K. m<

Q >? U>? + ∆T>?

∆T>? = Reflux Drum Temp − 318 K 7296 ∗ A O 8.9@ , A O =

Q O U O + ∆T O

U O = 0.568 kW/K. m< ∆T O = 34.8 K LP steam = $7.78/GJ MP steam = $8.22/GJ HP steam = $9.88/GJ $0.354/GJ $10/kg

12. Results and Discussion (Proposed Configuration w/o HI) 12.1. RD Column The bottom product at stage 35 contained the desired MMH and a small quantity of the undesired product MHOH. The distillate product was a small quantity of DME, produced as a byproduct, together with unreacted MeOH and excess MH, to be recycled. The conversions of MeOH and MH were 99.9 % and



38.5 %, respectively. The last stage temperature of the reactive stage was 422.4 K, which satisfied the temperature constraint for the reactive zone. The reboiler duty was 1.34 MW with a bottom temperature of 447.8 K, which used medium-pressure steam. The column diameter according to liquid holdup calculation was 3.90 m. Figures 15a and 15b show the composition and temperature profiles of the RD column, respectively.

A

1.0

Mole composition

0.8 0.6 0.4

Reactive Stages (2-12)

DME MEOH MH MMH MHOH

0.2 0.0

0 2 4 6 8 10 12 14 16 18 20 22 24 26 28 30 32 34 36

RD column stages

B 448 440

Temperature (K)

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

432 424

Reactive Stages (2-12)

416 408 400 0 2 4 6 8 10 12 14 16 18 20 22 24 26 28 30 32 34 36

RD column stages

Figure 15 RD column (A) Mole composition profile, and (B) temperature profile. 12.2. Column C2 The distillate had a small flow rate of 0.518 kmol/h with 99.9 % DME purity. The condenser duty was 0.0045 MW with a reflux drum temperature of 318 K. The bottom stream was mostly excess MH with a small amount of unreacted MeOH that was recycled back to the RD column. The reboiler duty was 0.72


Page 28 of 35

Page 29 of 35

MW with a base temperature of 501.9 K. The column diameter according to Aspen tray sizing was 0.83 m. Figures 16a and 16b show the composition and temperature profiles of column C2, respectively.

A

1.0

Mole composition

0.8 0.6 0.4

DME MEOH MH

0.2 0.0 0

2

4

6

8

10

12

Column C2 stages

B

500

Temperature (K)

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


450

400

350

300 0

2

4

6

8

10

12

Column C2 stages

Figure 16 Column C2 (A) Mole composition profile, and (B) temperature profile. 12.3. Column C3 Column C3 separated the bottom stream originating from the RD column into MMH and MHOH. The distillate was 49.02 kmol/h with 99.9 % MMH. The condenser duty was 0.64 MW with a reflux drum temperature of 352 K. The bottom stream was 0.49 kmol/h with 99.9 % MHOH. The reboiler duty was 0.18 MW with a base temperature of 401 K that was accompanied by low-pressure steam. The column diameter was 1.42 m. The composition and temperature profiles of column C3 are shown in Figures 17a and 17b, respectively.



A

1.0

Mole composition

0.8 0.6 0.4

MH MMH MHOH

0.2 0.0 0

2

4

6

8

10

12

14

16

Column C3 stages

410

B 400

Temperature (K)

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

390 380 370 360 350 0

2

4

6

8

10

12

14

16

Column C3 stages

Figure 17 Column C3 (A) Mole composition profile, and (B) temperature profile. 13. Economic Analysis Economic analysis was carried out for the proposed RD configuration in terms of the TAC. For the desired yield of MMH, the proposed RD configuration can reduce the energy consumption, and TAC by up to 18.2, and 10.8 %, respectively. The results of the heat integration sequence demonstrated superior performance for HI-sequence 1 that could further reduce the energy consumption, and TAC by up to 36.7, and 21.8 %, respectively. Table 6 summarizes the economic analysis of the proposed RD configuration and the different heat integration sequences applied.


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Table 6 Economic comparison of RD configuration with the base case.

Economic Analysis C1/RD Column Reactor cost (106$) Catalyst Cost (106$) @ 10$/kg Column Vessel Cost (106$) Exchanger Capital Costs (106$) Energy Costs (106$/year) Add. Exchanger Cost (106$) C2 Column Column Vessel Cost (106$) Exchanger Capital Costs (106$) Energy Costs (106$/year) Add. Exchanger Cost (106$) C3 Column Column Vessel Cost (106$) Exchanger Capital Costs (106$) Energy Costs (106$/year) Add. Exchanger Cost (106$) Total Capital Cost (106$) Total Energy Cost (106$/year) TAC (106$/year)

Heat Integration Sequences

Base Case

Proposed RD Configuration

Seq-1

Seq-2

Seq-3

Seq-4

0.1091 0.120 0.095 0.128 0.442 -

0.073 0.968 0.187 0.379 -

0.073 0.968 0.156 0.278 0.045

0.073 0.968 0.159 0.281 0.045

0.073 0.968 0.183 0.377 0.000

0.073 0.968 0.183 0.377 0.000

0.567 0.213 0.224 0.000

0.071 0.081 0.225 0.000

0.071 0.082 0.225 0.000

0.067 0.079 0.216 0.006

0.068 0.079 0.217 0.005

0.068 0.079 0.217 0.005

0.249 0.128 0.141 0.000 1.608 0.808 1.344

0.156 0.079 0.055 0.000 1.615 0.660 1.198

0.153 0.048 0.007 0.021 1.616 0.511 1.049

0.154 0.078 0.051 0.000 1.629 0.548 1.091

0.153 0.048 0.007 0.013 1.588 0.600 1.130

0.153 0.054 0.010 0.020 1.601 0.604 1.138

The performance comparison in terms of savings of the proposed RD configuration and heat integration sequences with the base case is illustrated in Table 7. Table 7 Performance comparison in terms of savings Heat Integration Sequences Base Case Proposed Savings compared to proposed RD configuration in costs (%) Configuration Configuration Seq-1 Seq-2 Seq-3 Seq-4 32.18 25.63 25.21 Energy 18.29 36.74 18.81 15.90 15.32 TAC 10.82 21.89

14. Conclusions The chemical 2-methoxy-2-methylheptane exhibits decreased solubility in water and has the ability to replace MTBE as a fuel additive, which causes environmental contamination. To meet the growing market demand for 2-methoxy-2-methylheptane, the design of an intensified configuration based on reactive distillation was proposed. A desired yield of MMH in the conventional process is usually



obtained by maintaining a large reactor size or by using an excess of MH. Both strategies are expensive in terms of capital and operating costs. The proposed configuration uses the RD technology to obtain the desired MMH yield at significantly lower capital and operating costs. The results demonstrate that the RD configuration can significantly reduce the energy consumption, and TAC by up to 18.2, and 10.8 %, respectively, compared to the conventional reactor column sequence. Furthermore, different heat integration sequences were investigated to further reduce the energy requirement. The heat integration sequence (1) realized savings of 36.7, and 21.8 %, for the energy, and total annual cost, respectively. Future work would include the development of an effective control structure for the proposed configuration. Acknowledgments This work was supported by Engineering Development Research Center (EDRC), funded by the Ministry of Trade, Industry & Energy (MOTIE) (No. N0000990), and by the Basic Science Research Program through the National Research Foundation of Korea (NRF), funded by the Ministry of Education (2015R1D1A3A01015621), and by the Priority Research Centers Program through the National Research Foundation of Korea (NRF), funded by the Ministry of Education (2014R1A6A1031189).

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Table of Contents (TOC)


Design of an Intensified Reactive Distillation Configuration for 2

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