Process Development, Assessment and Control of Reactive Dividing

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Process Development, Assessment and Control of Reactive DividingWall Column with Vapor Recompression for Producing N-Propyl Acetate Zemin Feng, Weifeng Shen, Gade Pandu Rangaiah, Liping Lv, and Lichun Dong Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.8b05122 • Publication Date (Web): 14 Dec 2018 Downloaded from http://pubs.acs.org on December 21, 2018

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Industrial & Engineering Chemistry Research

Process Development, Assessment and Control of Reactive DividingWall Column with Vapor Recompression for Producing N-Propyl Acetate Zemin Feng,a, b Weifeng Shen,b G.P. Rangaiah,a,* Liping Lv,c Lichun Dongb,c,d,* a

Department of Chemical and Biomolecular Engineering, National University of Singapore,

Singapore 117576 b School

of Chemistry and Chemical Engineering, Chongqing University, Chongqing 400044, China

c School

of Chemistry and Chemical Engineering, Collaborative Innovation Center for Green

Development in Wuling Mountain Area, Research Center for Environmental Monitoring, Hazard Prevention of Three Gorges Reservoir, Yangtze Normal University, Fuling 408100, Chongqing, China d Key

Laboratory of Low-grade Energy Utilization Technologies & Systems of the Ministry of

Education, Chongqing University, Chongqing, 40004, PR China

* Corresponding author: G.P. Rangaiah, Email: [email protected] Lichun Dong, Email: [email protected]

Abstract: As a combination of the conventional reactive distillation (RD) and diving-wall column, reactive diving-wall column (RDWC) is the highly thermally integrated process that has the advantages of higher thermodynamic efficiency, lower capital cost and smaller equipment size. In this study, the conceptual design of four different RD processes, i.e. the conventional RD, RDWC, heat integrated RDWC (HIRDWC) and vapor recompression heat pump assisted RDWC (VRHP-RDWC), was presented for the production of n-propyl acetate via the esterification of n-propanol with acetic acid. The results indicate that, compared with that of conventional RD process, the total annual cost of RDWC, HIRDWC and VRHP-RDWC intensified processes is reduced by 10.44%, 19.40% and 74.54%, respectively, while their thermodynamic efficiency is 9.96%, 15.52% and 25.53%, respectively, which are also significantly higher than that of conventional RD process (9.25%). Subsequently, since the VRHP-RDWC process exhibits the most favorable performance for intensifying the conventional RD process, two alternative control strategies were developed and assessed for the operation of VRHP-RDWC. Control performances demonstrate that the challenging 1

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VRHP-RDWC process can be operated smoothly under large disturbances of feed flowrate, water impurity in acetic acid feed and n-propanol feed as well as for set-point changes in temperature controllers. Keywords: Reactive distillation; N-propyl acetate; Reactive dividing-wall column; Heat integrated distillation; Heat pump; PI control

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INTRODUCTION Distillation has been the most widely used technology for the separation of

multicomponent mixtures in chemical and petrochemical industries; however, it is highly energy-intensive with thermodynamic efficiency of only 5-20%, resulting in the distillation processes accounting for more than 60% of the total energy consumption in chemical industry.1 Accordingly, various energy-efficient intensification technologies, e.g. heat-pump (HP) assisted distillation column,2-4 dividing-wall column (DWC),5,6 and internally heatintegrated distillation column (HIDiC),7 have been developed to reduce the energy consumption of distillation columns. Among of them, DWC has received the most attentions since it can save up to 30% of energy consumption and capital cost as well as reduced equipment space compared to conventional two-column sequence.8 DWC integrates the two columns of its thermodynamically equivalent Petlyuk system into one column shell by placing a vertical dividing-wall inside the shell. Further, it has a higher thermodynamic efficiency by avoiding the forced remixing of middle boiling components in the first column of the conventional two- column sequence.9 Reactive distillation (RD) is another important process intensification technology for reducing the energy consumption and capital cost of the conventional reaction-separationrecycle processes by integrating the formation and subsequent separation of products into a single column. This technology is very beneficial for reversible equilibrium-limited reactions, whose conversion can be increased by up to 100% due to the continuous removal of the formed products in the reactive zone of the RD column.10 On account of the multiple advantages of RD and DWC, the integration of RD and DWC resulting in reactive DWC

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(RDWC) has attracted considerable attention in the past three decades since it can bring the synergistic effect of the two technologies. The first experimental study of RDWC was reported by Sander et al.11 in 2007, who designed a lab scale RDWC using a 4-column sequence at a BASF laboratory. It is for the hydrolysis of methyl acetate and water, and the reactive zone of the RDWC was packed with Sulzer Katapak-SP 11 filled with Amberlyst 48 catalyst. Their lab scale RDWC was further scaled up to a 220 mm diameter industrial scale column at Sulzer Chemtech, in which the MellapakPlus 752Y and Katapak-SP 11 were used for the non-reactive and reactive zones, respectively. Sander et al. concluded that the operation of RDWC is possible and believed that the study can be served as the basis for further investigation of RDWC technology. After this, several experimental studies for various reaction system were implemented by other researchers, which further demonstrates that the operation of RDWC is feasible and practical.12-17 Among these, Egger and Firg15,17 experimentally investigated steady state and dynamic operation of a pilot RDWC column using PI controllers concluding that the RDWC can be safety and reliable operated for a wide range of operating conditions. Also, they proposed a rigorous mathematical model of the RDWC, validated it against the experimental data on conversion, temperature profile and product concentrations, and concluded that the experimentally verified model shows very good agreement with the experimental data and can serve as the solid basis for the design of RDWC. The process design, optimization and control of the RDWC for various reaction systems were investigated by employing rigorous simulation using its thermodynamic equivalent model. Kiss et al.18 proposed a novel biodiesel process for the synthesis of fatty acid methyl ester based on a RDWC, which can save the energy requirement by over 25% compared to the conventional process. Santaella et al.19 used a multi-objective optimization algorithm to design a RDWC for tri-ethyl citrate production for optimizing economics and controllability simultaneously; their results show that the proposed RDWC can save around 40% energy and 11% capital costs. Zheng et al.20 introduced RDWC for the synthesis of diethyl carbonate via transesterification of dimethyl carbonate with ethanol, demonstrating savings of 18.7% of energy and 13.9% of total annual cost (TAC) compared to the conventional RD process. Li et

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al.21 showed that the RDWC technology can save the energy requirement of hydrolysis of methyl acetate by 20.1% compared to the RD process. In summary, these studies indicate that the RDWC can save up to 40% energy consumption and 15% TAC compared to conventional RD process, which indicate huge potential for the intensification of RD processes by employing RDWC technology. The recent review work by Weinfeld et al.22 summarizes the advances in RDWC technology ranging from the shortcut conceptual design, process design and optimization to experimental studies. In the operation of DWCs, for providing the required reflux, large amount of condensation latent heat of the overhead vapor stream is usually removed by cooling water as the lower quality waste heat, which leads to lower energy utilization and thermodynamic efficiencies. For this reason, HP is an attractive technology that can fully or partially recover the condensation latent heat of the overhead vapor stream by updating the low-temperature thermal energy to high-temperature thermal energy using external mechanical energy. The typical HP technologies include vapor recompression, flashing of column bottoms stream to a lower pressure and later vapor compression (referred to as bottoms flash), closed-cycle compression, absorption heat pump and absorption heat transformer. Chew et al.23 investigated the application of these various HP technologies in a number of DWC processes, and they found that vapor recompression or bottoms flash is usually more economical. Li et al.2 demonstrated that, by employing the vapor recompression HP (VRHP) to azeotropic DWC, the total operating cost and TAC can be reduced by 48% and 32%, respectively. Luo et al.24 showed that the energy requirement for bioethanol separation drops from 2.07 kWh/kg (7.45 MJ/kg) for the classical three-column separation sequence to 1.24 kWh/kg (4.46 MJ/kg) for the VRHP assisted extractive DWC process. Chen et al.25 proposed a VRHP assisted DWC for separating ethylene diamine from water, demonstrating that the energy consumption and TAC can be saved respectively by 43% and 24% compared to the conventional azeotropic distillation process. All these studies conformed that, although additional capital cost is required for the addition of VRHP system to the distillation process, the VRHP assisted DWC can provide significant economic advantages from the reduced operating cost and environmental benefit due to the improved thermodynamic efficiency.

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Up to now, although the application of VRHP in DWC has been extensively studied, the reports on the combination of VRHP and RDWC are surprisingly scanty. Feng et al.26 proposed an energy-efficient process to produce n-propyl propionate via the esterification of 1-propanol and propanoic acid by incorporating VRHP assisted RDWC (VRHP-RDWC), wherein the VRHP is used to compress the overhead vapor stream to provide high temperature thermal energy to the intermediate or bottom reboiler of RDWC. Their results indicated that the proposed VRHP-RDWC can save 57% of operating cost and 31% of TAC with payback period of 5 years, compared to the conventional RD process. Shi et al.27 applied VRHP-RDWC to intensify the synthesis of 1,3-dichlorohydrin, 2,3-dichlorohydrin, and byproduct water along with the intermediate products of 1-monochlorohydrin and 2monochlorohydrin through the reaction of glycerol and hydrochloric acid; they showed that the VRHP-RDWC process can reduce TAC by 50% with payback period of 3 years compared to the conventional RD process. However, due to the highly integrated physical space structure and coupled heat exchanger network, controllability and operability of the VRHP-RDWC process is much more challenging than the conventional RD distillation process. In the present study, the conceptual design of four different RD processes, i.e. the conventional RD, intensified RDWC, heat integrated RDWC and VRHP assisted RDWC, for producing n-propyl acetate via the esterification of n-propanol with acetic acid was proposed and analyzed. TAC was utilized as the economic criterion to optimize the developed processes using the mixed-integer nonlinear programming (MINLP), and the thermodynamic efficiency was compared by implementing exergy analysis. The results demonstrated the superiority of VRHP-RDWC over the other processes. Subsequently, controllability and operability of VRHP-RDWC process were evaluated under two alternative control strategies. According to our knowledge, this is the first report on the evaluation of the controllability and operability of VRHP-RDWC process. The rest of this paper is organized as follows. In section 2, reaction kinetics and thermodynamic data are introduced. In section 3, economic and thermodynamic evaluation criteria are described, and the MINLP method employed to optimize the proposed RD

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processes is outlined. In section 4, conventional RD and RDWC processes for producing PrAc via the esterification of n-propanol with acetic acid were developed and assessed, and then the heat integrated RDWC and VRHP assisted RDWC processes were designed to intensify the proposed RDWC process. Subsequently, the controllability and operability of the intensified VRHP-RDWC process are evaluated in section 5. In section 6, main findings of this study were summarized.

2

REACTION KINETICS AND THERMODYNAMIC DATA

2.1

Reaction Kinetics N-propyl acetate (PrAc) is widely used as a solvent for liquid rotogravure, flexographic

printing inks, coatings, wood lacquers, aerosol sprays, nail care and cosmetics, and also as a flavor additive and fragrance due to its characteristics of colorless and odor of pears. It is commonly produced by the esterification of n-propanol (PrOH) with acetic acid (AA) together with water as the byproduct. The reversible esterification reaction is performed in liquid phase and is expressed as follows: CH3COOH + HOCH2CH2CH3 CH3COOCH2CH2CH3 + H2O Acetic acid

N ― propanol

(1)

N ― propyl acetate

The reaction kinetics of the reversible esterification of n-propanol with acetic acid to PrAc over various homogeneous catalyst28 and heterogeneous catalysts28-33 has been extensively studied. However, the utilization of the homogeneous catalyst in RD process can make the subsequent product separation more energy intensive and complex. On the other hand, the heterogeneous catalyst is more suitable for RD process due to the convenient installation and easy separation of the formed product. Thus, the Langmuir-HinshelwoodHougen-Watson (LHHW) kinetic model (Eq. 2) provided by Huang et al.,33 who performed the reversible reaction in liquid phase and catalyzed by ion exchange resin Amberlyst 15, is used in the present study. 𝛼AA𝛼PrOH ― 𝑟 = 𝑀cat𝑣𝑖𝑘𝑓

1 𝛼 𝛼 𝐾eq PrAc H2O

(2) 2

(1 + 𝐾s1𝛼AA + 𝐾s2𝛼PrOH + 𝐾s3𝛼PrAc + 𝐾s4𝛼H2O)

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Here, 𝑀𝑐𝑎𝑡 is the catalyst weight, 𝑣𝑖, 𝛼𝑖 and 𝐾si are respectively the stoichiometric coefficient, activity and adsorption constant of component i; and 𝑘𝑓 and 𝐾eq are respectively the forward reaction rate constant and equilibrium constant, which are given by: 𝐸𝑓 𝑘𝑓 = 𝑘0exp ― 𝑅𝑇 𝐾eq = exp

( ) (

)

―475.42 + 4.3223 𝑇

(3) (4)

Here, 𝑘0 and 𝐸𝑓 are respectively the pre-exponential factor and activation energy of forward reaction; T (K) is the absolute reaction temperature; R (8.314 kJ∙kmol-1∙K-1) is the ideal gas constant. The kinetic parameters of Eqs. 2 and 3 are summarized in Table 1, and their mean values were used in the present study. The reaction kinetic model was coded in a FORTRAN subroutine that was inserted into Aspen Plus to calculate reaction rate of each component in the RD column. Table 1. Kinetic model parameters33 Parameter k0 (mol∙s-1∙g-1) Ef (J∙mol-1) Ks1 Ks2 Ks3 Ks4 2.2

value 2.7163×106 ± 1.498% 5.4710×104 ± 0.002% 4.5383 ± 0.910% 7.9183 ± 0.834% 1.4662 ± 0.910% 8.7032 ± 0.910%

Thermodynamic Data For accurately predicting non-ideal vapor-liquid equilibrium (VLE) and vapor-liquid-

liquid equilibrium of the quaternary system, the UNIQUAC-HOC model was used to calculate the activity coefficients of all components in the reactive system, and the HaydenO’Connell model is used to predict the vapor nonideality. In the present study, binary interaction parameters (Table 2) used in UNIQUAC-HOC model were taken from the Aspen 9.0 databank. The system has four azeotropes that are shown in Table 3 together with the normal boiling point of the four pure components and topologic features of the system. As can be seen in this table, predicted azeotropic compositions and temperatures of the system are in

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good agreement with the experimental data, which conforms the accuracy of UNIQUACHOC model in Aspen built-in databank.

Table 2. UNIQUAC-HOC binary parameters for the reactive system component i AA AA PrOH PrOH PrAc AA

component j PrOH H2O PrAc H2O H2O PrAc

source APV90 VLE-HOC APV90 VLE-HOC APV90 VLE-HOC APV90 VLE-HOC APV90 LLE-ASPEN APV90 VLE-HOC

aij 0 0.7446 0 1.8387 1.9620 0

aji 0 0.0042 0 -2.4087 -1.8897 0

bij 125.9370 -615.2640 -3.7516 -668.9700 -1049.1300 174.4040

bji -117.8320 196.8990 -78.6062 620.7850 450.6320 -374.5850

Table 3. Boiling points, azeotrope data and topologic features of the reactive system34 Components AA PrAc N-propanol H2O

Composition (mole) 1 1 1 1 0.480/0.520* PrAc/H2O (A) 0.480/0.520# 0.613/0.387* PrOH/PrAc (B) 0.633/0.367# 0.432/0.568* PrOH /H2O (C) 0.406/0.594# 0.081/0.394/0.525* PrOH /PrAc/H2O (D) 0.121/0.373/0.506#

Boiling point (°C) 117.90 101.44 97.20 100.02 82.40* 82.75# 94.48* 94.75# 86.95* 87.64# 82.21* 82.31#

Singular point stable saddle saddle saddle saddle saddle saddle unstable node

Note: * - experimental data from the reference 34; # - predicted data using the binary parameters taken from the Aspen built-in databank.

3

EVALUATION CRITERIA

3.1

Economic Evaluation In the present study, simulation of the various RD processes was implemented in Aspen

Plus 9.0. Optimization was carried out using an algorithm in Matlab 2017a, which was directly interfaced with Aspen Plus via COM technique. Before optimizing the process, the RD processes were designed under the following assumptions. First, the composition of npropanol feed is set as 99.5 mol% n-propanol and 0.5 mol% water that is close to the

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industrial product purity of n-propanol, while the acetic acid feed composition is set as 95 mol% acetic acid and 5 mol% water.35 Second, the product purity of PrAc should be more than 99 mol% while keeping the impurity of acetic acid in PrAc below 100 ppm.35 Third, the installed catalyst Amberlyst 15 with a density of 770 kg/m3 occupies half of the tray and column base holdup. The weir height is assumed as 0.1016 m while the diameter of the RD column diameter is 1.22 m based on the base-case simulation results, which give tray liquid holdup of 0.107 m3; the column base liquid holdup is taken to be 10 times the tray holdup.35 Fourth, the life time of the solid catalyst is assumed to be 3 months.36 Fifth, the pressure drop of the columns is assumed as 0.0068 bar per stage/tray.37 Sixth, the tray spacing in the columns is assumed as 0.6096 m. Finally, column diameters are calculated by tray sizing in Aspen Plus with the maximum flooding of 80%. Under the above assumptions, TAC is minimized to obtain the optimal steady-state process; it is defined by: TCC TAC = + AOC payback period

(5)

Here, the total capital cost (TCC) is the sum of capital costs of columns, trays, decanter, condensers, reboilers, heat exchangers and compressor, while the cost of pipes, pumps and valves and the subsequent wastewater treatment cost is neglected; the annual operation cost (AOC) includes the annual cost of steam, cooling water, electricity and catalyst. In the present study, the payback period is assumed to be 3 years while the operation time is set as 8000 hours per year. The capital costs were calculated by using the correlations provided by Douglas,38 which are given in the Appendix. Heat transfer coefficient of condensers and heat exchangers was assumed to be 0.852 kW/(K∙m2) while that of reboilers was assumed as 0.568 kW/(K∙m2).39 Marshall & Swift index (M&S) was assumed as 1468.6 in 2012.24 Efficiency of the compressor and its motor is assumed to be 0.8 and 0.9, respectively.4 The maximum boiling point of the reaction system is 117.9°C at 1 atm, which makes it possible to use low pressure steam (LPS) to supply thermal energy for reboiler duties. The prices of LPS, cooling water, catalyst and electricity are listed in Table 4.

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Table 4 Utilities and their unit prices Utility Low pressure steam (160°C, 5 barg) Cooling water (inlet: 30°C; outlet: 40°C) Catalyst Electricity

Unit price 13.2800 $/GJ40 0.3540 $/GJ40 7.7162 $/kg41 0.0600 $/kWh40

For the optimization of the RD process, decision variables for the objective function in Eq. 5 include both discrete variables (i.e. number of stages in columns) and continuous variables (i.e. flowrates, ratio and pressure). Therefore, TAC minimization in Eq. 5 is a typical MINLP problem, and it can be mathematically stated as follows. min TAC = 𝑓(𝑥) 𝑥

(6a)

s.t. ℎ(𝑥) = 0

(6b)

𝑔(𝑥) ≤ 0

(6c)

𝑥min ≤ 𝑥 ≤ 𝑥max

(6d)

Here, x is the decision variables vector; f(x) represents the equation defined in Eq. 5; h(x) represents the equality constraints, i.e., the black-box nonlinear process model built in Aspen Plus; g(x) represents the inequality constraints in the optimization problem; and Eq. 6d are the bounds on the decision variables. In the present study, inequality constraints of the RD process are: purity of the PrAc product should be more than 99 mol%, impurity of acetic acid in the product PrAc should be below 100 ppm, and the maximum operation temperature in the reactive section of the RD columns should be lower than 120°C, which is the allowable maximum temperature of the catalyst Amberlyst 15. Therefore, the inequality constraints, g(x) in Eq. 6c can be defined as: 𝑔1 0.99 ― 𝑥PrAc 𝑔(𝑥) = 𝑔2 = 𝑥AA ― PrAc ― 0.0001 (7) 𝑔3 𝑇RD ― 120

[][

]

Here, xPrAc is the molar purity of the product PrAc; xAA-PrAc is the molar impurity of acetic acid in the product PrAc; and TRD is the reboiler temperature of the RD column, which is the maximum operation temperature in the reactive section of the RD column.

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The MINLP problem in Eq. 6 cannot be solved by deterministic optimizers due to the difficulty of obtaining the required gradients of nonlinear equations embedded in Aspen Plus that is usually treated as a black model in the optimization problem. Therefore, to solve the black-box optimization problems, a suitable alternative is stochastic methods that are capable of robustly optimizing the complex RD process. The stochastic methods do not need the gradient information. In the present study, MINLP problems (Eq. 6) are solved by the optimization package: NOMAD that implements the “Mesh Adaptive Direct Search (MADS)” algorithm for constrained optimization,42 which was carried out on a 64 bit computer with an Intel Core i7 processor at 3.4 GHz with 16 GB of RAM and running Windows 7. 3.2

Exergy Analysis Entropy generation and exergy analysis, based on the second law of thermodynamics, are

the fundamental tools for analyzing and improving the thermodynamic efficiency of a chemical process. In thermodynamics, the exergy of a system is the maximum useful work possible during a process that brings the system into equilibrium with its surroundings. Briefly, exergy is the energy that is available to be used; therefore, it can be used to measure the quality and utility efficiency of the energy that cannot be measured by the energy balance, i.e., first law of thermodynamics. By referring to the book of Seader and Henley,43 the exergy analysis can be implemented by Eqs. 8-13. The minimum separation work, WSEP, is the shaft work required to conduct an ideal reversible chemical process; it can be calculated by: 𝑊SEP =

∑ 𝐸𝑥 ― ∑ 𝐸𝑥 𝑖

out of system

𝑖

(8)

in to system

Here, Ex refers to the stream availability, also, named exergy, which can be defined by: (9)

𝐸𝑥 = 𝐻 ― 𝑇0𝑆

Here, H is the enthalpy of the stream, S is the entropy of the stream and T0 is the absolute temperature (K) of the surroundings. Since the enthalpy and entropy of each stream can be obtained from the simulation results in Aspen Plus, thus, exergy change of the system, ΔEx, can be easily calculated by (10)

∆𝐸𝑥 = ∆𝐻 ― 𝑇0∆𝑆

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Here, ΔH and ΔS are respectively the enthalpy change and entropy change of a system. For a real chemical process, the required separation work is usually more than the minimum separation work due to the irreversibility of the process that leads to the entropy production of the system, ΔSirr. This lost work is defined as the exergy loss, ∆Exloss: (11)

∆𝐸𝑥loss = 𝑇0Δ𝑆irr The exergy loss, ∆Exloss, can also be calculated from: 𝑇0 ∆𝐸𝑥loss = 𝐸𝑥𝑖 + 𝑄𝑗 1 ― + 𝑊𝑘 ― 𝑇𝑠𝑗

∑[

in to system

) ]

(

∑ [𝐸𝑥 + 𝑄 𝑙

out of system

𝑚

(

1―

𝑇0 𝑇𝑠𝑚

) ] + 𝑊𝑛

(12)

Here, Q is the heat flow into or out of the system; W is the shaft work crossing the boundary of the system; Ts is the temperature of the heat source; i and l respectively represent the streams into or out of the system; j and m respectively represent the heat into or out of the system; k and n respectively represent the shaft work on or by the system. The thermodynamic efficiency of the system, η, can be defined as: 𝑊SEP 𝜂= 𝑊SEP + ∆𝐸𝑥loss

(13)

In the present study, the surrounding temperature and pressure are assumed as 25°C and 1 atm, respectively. Average temperature of cooling water is assumed as 35°C (based on its inlet and outlet temperatures of 30°C and 40°C, respectively) while the temperature of LPS is 160°C.

4

STEADY-STATE PROCESS DESIGN

4.1

Conventional RD Process Figure 1 shows the residue curve map and liquid-liquid equilibria of the quaternary

system for producing PrAc via the esterification of n-propanol (PrOH) with acetic acid (AA), wherein acetic acid is the heaviest component with normal boiling point of 117.9°C at 1 atm while the ternary azeotrope of PrAc/PrOH/H2O is an unstable node with boiling point of 82.31oC at 1 atm. Here, composition of each azeotrope (A, B, C and D shown in Figure 1) is given in Table 3. Moreover, azeotropes of PrAc/H2O (with boiling point of 82.75oC at 1 atm) and PrAc/PrOH (with boiling point of 94.75oC at 1 atm) are saddle nodes whose boiling points are also lower than the boiling points of two feed reactants. Therefore, in the RD 12


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process for the esterification of n-propanol with acetic acid to produce PrAc, the fresh reactants, n-propanol and acetic acid can be directly fed into the RD column base resulting no bottoms outlet stream. Meanwhile, the ternary azeotrope of PrAc/PrOH/H2O can be obtained at the top of the RD column, and then it can be naturally separated, after condensation, into two liquid phases in the decanter. The aqueous phase with mostly water can be directly drawn-out of the system. A portion of the organic phase is sent back to the top of the RD column as reflux while the rest of the organic phase is fed into the top of the product column to obtain pure PrAc product (boiling point 101.4oC) at its bottoms stream. The vapor stream from the top of the product column is condensed and fed into the decanter. In summary, as the conceptual design flowsheet shown in Figure 2, the conventional RD process for producing PrAc consists of one RD column (RDC) with decanter and a stripper (i.e. product column, PC) used to obtain the high purity PrAc at its bottoms stream.

Figure 1. Ternary residue curve maps (RCMs) and liquid-liquid equilibria (LLEs) of the quaternary reactive system

The operating pressure of the column is an important process condition; a higher pressure can lead to a higher reaction temperature in RD column and consequently increase reaction rates; however, the more expensive steam may be required to supply the thermal energy to the reboiler of the RD column. Furthermore, the maximum reaction temperature is also constrained by the allowable maximum operating temperature of the catalyst, Amberlyst

13



15 in the RD column. Hence, operating pressure of the RD column is considered as a decision variable in the optimization problem in order to obtain the optimal temperature in the reactive section of the RD column. In the present study, the operating pressure (P) of the RD column is set equal to that of the product column in the conventional RD process.

Figure 2. Conceptual design flowsheet of conventional RD process

Figure 3. Reduction of TAC with number of iterations for optimizing (a) conventional RD (CRD), (b) RDWC and (c) VRHP-RDWC Since the molar flowrate ratio of the two feed reactants (FR) should be determined in order to satisfy the reaction stoichiometry, the feed flowrate of acetic acid is optimized to

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indirectly obtain the optimal FR. In addition, number of theoretical stages of the columns should also be determined optimally. Thus, for the conventional RD (CRD) process, there are 6 design degrees of freedom to be determined by solving the optimization problem in Eq. 6 with the constraints in Eq. 7. They are: numbers of theoretical stage of the rectifying section (N1) and reactive section (Nrn) of the RD column and theoretical stages of the product column (N2) including the respective reboiler, flowrate of the organic phase fed to the RD column top (L1, kmol/h), operating pressure (P, bar), flowrate of the overhead vapor from the product column top (D2, kmol/h) and the feed flowrate of acetic acid (FAA, kmol/h). These decision variables and their bounds were listed in Table 5. Table 5. Bounds on decision variables of the optimization problems Process CRD

Variables vector, x

[𝑁1,𝑁rn,𝑁2,𝐿1,𝑃,𝐷2,𝐹AA]

Lower and upper bounds 𝑥𝑇Lower = [10, 20, 4, 115, 1, 30, 50]

𝑇

𝑇

𝑥𝑇Upper = [25, 35, 15, 140, 1.5, 40, 53] 𝑥𝑇Lower = [10, 15, 2, 2, 110, 1, 225]

𝑇

RDWC

[𝑁1,𝑁rn,𝑁2t,𝑁2b,𝐹L,𝑃,𝑉top]

VRHPRDWC

[𝑅v, 𝑅cp, ∆𝑇hex3,𝑄R ― PC,𝐹L]𝑇

𝑇

𝑇

𝑥𝑇Upper = [25, 30, 10, 10, 130, 1.5, 240] 𝑥𝑇Lower = [0.05, 3, 5,152,118]𝑇 𝑥𝑇Upper = [0.08, 4, 35,158,122]

𝑇

𝑇

The MINLP problem for the conventional RD process was solved using NOMAD package in MATLAB. The reduction of TAC with number of iterations for optimizing CRD process is shown in Figure 3a, which confirms the convergence of the optimization method. Plots in Figure 3 show only the feasible solutions of each optimization problem. Moreover, the initial value may affect convergence of the optimization method. However, practically the same optimal solution was obtained for each of conventional RD, RDWC and VRHP-RDWC processes by using 5 different initial values. This confirms that the global solution for the optimization of the various processes was obtained by the optimizer used; computation time for each optimization run ranged from 30 to 100 minutes depending on the process. Figure 4 presents the optimal conventional RD process with detailed information including stream data, heat duties, equipment sizes and operating conditions. The RD column has a total of 42 stages, 17 for separation section and 25 for reactive section including reboiler

15



whereas the product column has 6 stages including reboiler. The product purity of PrAc is 99.01 mol%, satisfying the purity constraint, and the acetic acid impurity in product PrAc is 0.00865 mol%, below the specified impurity constraint in Eq. 7. In addition, the byproduct water purity is 99.63 mol%; this stream has 0.0027 mol% PrAC and 0.10 mol% acetic acid, mainly due to the liquid-liquid equilibrium between organic and aqueous phases in the decanter. The optimal operating pressure is 1.10 bar, giving maximum operating temperature of 115.43°C in the reactive section of the RD column, which is lower than the allowable maximum temperature (120oC) of Amberlyst 15 catalyst. Reboiler duties of RD and product columns are 2472.71 kW and 406.37 kW, respectively, while the condenser duties of the two columns are 2296.20 kW and 362.87 kW, respectively. Since both the feed reactants and the formed products need to be vaporized in the RD column, it is natural that the corresponding reboiler duty is large.

Figure 4. Optimal flowsheet of conventional RD process It can be observed from the composition profiles of the RD column in conventional RD process (Figure 5a) that the mole fraction of water firstly decreases from the first stage to

16


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third stage, and then increases with the stage number; on the other hand, the mole fraction of PrAc slightly increases from the first stage to third stage, and then decreases with increasing stage number. These are due to remixing that results in lower thermodynamic efficiency of the RD columns since more energy will be required in order to re-purify the remixed mixture. The remixing effects can be effectively eliminated by using RDWC configuration (described in the next section) that can improve the thermodynamic efficiency of the conventional RD process, resulting in lower energy consumption of the process. The remixing effects in conventional RD and RDWC are described and compared in detail in the next section.

Figure 5. Composition profiles in (a) RD column of conventional RD (CRD-RD), (b) product column of conventional RD (CRD-PC), (c) RD column of RDWC (RDWC-RD) and (b) product column of RDWC (RDWC-PC) 4.2

RDWC Configuration The conceptual flowsheet of RDWC and its thermodynamically equivalent model are

shown in Figure 6. The overhead vapor stream of RDWC (Vtop) is condensed using cooling water and then fed into the decanter, where the condensate separates into two liquid phases;

17



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the aqueous phase is drawn-out of the system while the organic phase is returned as reflux to the top of RDWC.

Figure 6. Conceptual flowsheets of (a) RDWC and (b) its thermodynamically equivalent model In the design of RDWC, the molar flowrate ratio of the two feed reactants (FR) is maintained same as that of the conventional RD process (in section 4.1) in order to compare the economic performance and thermal efficiency with the conventional RD process as the base case. For the liquid split ratio (βL) to either of the side of the dividing wall can significantly affect the energy consumption and capital cost of the RDWC, therefore, the liquid flowrate to left side of the dividing wall (FL, liquid molar flowrate fed to the top of the RD column) is chosen as a decision variable to indirectly obtain the optimal βL. Thus, decision variables for RDWC optimization are: numbers of theoretical stages of the separation section (N1) and reactive section (Nrn) of the RD column, numbers of theoretical stages of the top section (N2t) and lower section (N2b) of the product column, liquid flowrate to the left side of the dividing wall (FL), operating pressure (P) and overhead vapor flowrate (Vtop). These decision variables and their bounds are given in Table 5. Since there is only one column shell for RDWC, the equivalent diameter, Dd (defined by Eq. 14) is used to calculate TCC of RDWC. 𝐷d = 2 𝐷21 + 𝐷22

(14)

18


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Here, D1 and D2 are respectively the diameter of the RD column and the lower section of the product column.

Figure 7. Optimal configuration of RDWC Figure 3b shows the reduction of TAC with number of iterations for optimizing RDWC process, which confirms the convergence of optimization method. Figure 7 shows the optimal RWDC configuration with the information of stream data, operating conditions and duties. The optimized operating pressure of RDWC is 1.20 bar, slightly more than that of conventional RD process. In RDWC, the RD column has a total of 41 stages, 21 for separation section and 20 for reactive section including the reboiler while the product column has 9 stages including the reboiler but excluding the condenser. The liquid flowrate from the third stage is separated into two parts, which flow down to their respective sides of the dividing wall. Product purity of PrAc in RDWC is 99.01 mol%, same as that in conventional RD process; acetic acid impurity in PrAc product is 0.00727 mol%, smaller than the value in conventional RD process. Water purity of 99.47 mol% in RDWC process is slightly smaller than that in conventional RD process due to the different composition of the inlet stream to the decanter; average mole fraction of PrAc in the two inlet streams together to the decanter in conventional RD process is 0.6571, higher than PrAc mole fraction (0.6138) of the overhead

19



stream in RDWC process. Reboiler duties of the RD and product columns in RDWC are respectively 2432.51 kW and 154.98 kW, which are lower than their respective values in conventional RD process. The condenser duty of RDWC is 2354.18 kW, which is also lower than the total condenser duty of 2659.07 kW in conventional RD process. The reboiler temperature of the RD column is 117.22°C, meeting the constraint on the allowable maximum temperature (120oC) of Amberlyst 15. The composition profiles of the RD column in RDWC (Figure 5c) indicate that the remixing effect occurring in the RD column of the conventional RD (CRD) process is eliminated in the intensified RDWC process. Here, the stage number in Figures 5c and 5d refers to the stages of the RD column (RDC) and product column (PC) in Figure 7, respectively; in other words, stage 1 in RDC is different from stage 1 in PC. As shown in Figures 5b and 5d, PrAc mole fraction on the first stage of the product column in both RD and RDWC processes are very close (0.8933 for CRD versus 0.8975 for RDWC), which means the returned organic liquid phase from the decanter in both processes has similar composition. On the other hand, PrAc mole fraction on the first stage of the RD column in RDWC is more than that in the conventional RD process (0.9473 in Figure 5c for RDWC versus 0.9279 in Figure 5a for CRD); further, it is very close to the maximum mole fraction of PrAc (0.9494 in Figure 5a) on the third stage of the RD column in conventional RD process. These results imply that the returned low purity organic liquid phase (with PrAc mole fraction of 0.8933 in Figure 5b) from the decanter needs to be purified in both RD and product columns in the conventional RD process. Especially for the RD column in the conventional RD process, the returned organic liquid phase from the decanter firstly mixes with the high purity vapor from the lower stages and then it is re-purified to higher purity liquid from the first stage to third stage, which refers to remixing effect and results in lower thermodynamic efficiency of the RD column. However, for the RDWC process, the returned lower purity organic phase liquid from the decanter is directly purified in the separation section (stages from 1 to 3 in Figure 5d) above the dividing wall in RDWC, and then goes down to the two sides of the dividing-wall, which effectively avoids remixing effect in the RD column and results in higher thermodynamic efficiency than the conventional RD process.

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Moreover, the RDWC process can save 10.08% steam and 11.47% cooling water compared to the conventional RD process. These results demonstrate the advantages of RDWC over conventional RD process in terms of improved thermodynamic efficiency and lower energy consumption. However, from the point of thermal integration, the overhead vapor stream from RDWC top and the product PrAc can be used to preheat the two feed reactants; then, both the reboiler and condenser duties of the RDWC will decrease. 4.3

Heat Integrated RDWC (HIRDWC) As shown in Figure 7, temperatures of the overhead vapor stream and the product PrAc

stream in RDWC process are 94.18°C and 107.98°C, respectively, which are much higher than the temperature of the two feed reactants fed into the RD column. Therefore, in the present study, two heat exchangers were added in the heat integrated RDWC (HIRDWC) process to preheat the two feed reactants by using the energy of the overhead vapor stream and PrAc product. For all the heat exchanges, the temperature approach is assumed as 5°C to provide sufficient temperature driving force for heat transfer.44,45 Preheating of the two feed reactants mainly affects the reboiler duty of the RD column, and it has negligible effect on the optimal configuration of RDWC. Therefore, HIRDWC configuration can be directly obtained by adding preheaters in the optimal RDWC configuration without further optimization. Figure 8 shows the optimal flowsheet of the HIRDWC process, wherein, the RDWC design parameters remain same as those in the optimal RDWC process. Before feeding into the column base, acetic acid feed is preheated to 89.17°C by the overhead vapor stream from the RDWC top; the vapor stream is simultaneously partially condensed to 93.66°C and then sent to the condenser for further cooling/condensation using cooling water. On the other hand, the n-propanol feed is preheated to 102.97°C by the product PrAc stream that is simultaneously cooled to 49.02°C before leaving the system. By preheating the feed streams, the reboiler duty of the RD column is reduced from 2432.51 kW in RDWC to 2150.67 kW in HIRDWC while the condenser duty is reduced from 2354.18 kW in RDWC to 2253.02 kW in HIRDWC. In other words, HIRDWC can save 10.89% steam and 4.30% cooling water compared to RDWC without heat integration.

21



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Figure 8. Optimal configuration of RDWC with feed pre-heating 4.4

VRHP Assisted RDWC (VRHP-RDWC) In HIRDWC process, a large amount of steam is required to supply 2150.67 kW of

thermal energy to the reboiler of the RD column, while the 2253.02 kW of condensing latent heat is removed by cooling water as waste heat, which can be partially utilized by using the VRHP to further improve the thermodynamic efficiency of the RDWC process. Therefore, the VRHP assisted RDWC (VRHP-RDWC) process is designed and discussed in this section. The feasibility of a HP system can be evaluated by the coefficient of performance (COP) defined as:46 COP =

𝑇c 𝑄 1 = = 𝑊 𝜂 𝑐 𝑇r ― 𝑇c

(15)

Here, Q is the reboiler duty, W is the work required for the HP, ηc is the Carnot efficiency, Tr is the reboiler temperature in K and Tc is the condenser temperature in K. According to Plesu et al.46, use of HP in the distillation process is favorable when COP is 10 or higher; this is considering production cost of electricity is more than 5 times that of thermal energy and the Carnot efficiency is 0.5. The reboiler temperature of the RD column and the condenser temperature in RDWC process are 117.23°C (390.38 K) and 94.18°C (367.33 K), respectively. These give a COP of 15.94 by Eq. 15, which indicates that the utilization of VRHP can significantly improve the thermodynamic efficiency of the RDWC process. 22


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Figure 9. The compressor discharge temperature (Tdis) and the condensing temperature (Tcon) of the compressed vapor stream at various discharge pressures of the compressor (Pdis)

Since the required reboiler duty of the RD column is much larger than that of the product column in RDWC, the recovered latent heat of the overhead vapor stream is used via HP to supply thermal energy for the reboiler of the RD column, while the reboiler duty of the product column is still supplied by low pressure steam. Furthermore, the compression ratio of the compressor in HP system is an important operating condition that can significantly affect the discharge temperature (Tdis) and pressure (Pdis) of the compressor. It can be observed from the variation of Tdis and the condensing temperature of the compressed vapor stream (Tcon) at various Pdis (Figure 9) that when Pdis increases to 4.2 bar, Tcon increases to 122.35oC, thus the Tcon is 5oC higher than the reboler temperature of the RD column (TRD) of 117.23°C to satisfy the required driving force for heat transfer in the reboiler of the RD column. This makes VRHP-RDWC possible to utilize the condensing latent heat of the compressed overhead vapor to supply required energy for the reboiler of the RD column. After the compressed overhead vapor condenses in the RD column reboiler, condensate temperature is still high and can be used to preheat the overhead vapor stream before feeding it into the compressor. This can further increase Tdis, leading to the enhancement of the residual heat recovery of the condensate from RD column reboiler. However, Tdis cannot exceed 150°C for safe operation of the compressor,47 which is included as another constraint for the optimization of VRHPRDWC process.

23



Figure 10. Optimal flow diagram of VRHP-RDWC The flowsheet diagram of VRHP-RDWC process is shown in Figure 10, wherein, the overhead vapor stream is divided into two parts: a small portion is directly condensed using cooling water and then fed into decanter while the rest is preheated by the condensate from the RD column reboiler and then compressed. The compressed vapor at higher pressure and temperature outflowing from the compressor is then used to supply the thermal energy required for the reboiler of the RD column by releasing the latent heat; the generated condensate is firstly used to preheat the overhead vapor stream before the compressor, and then the acetic acid feed. In addition, the n-propanol feed is preheated by the product PrAc outflowing from the product column reboiler. Here, the theoretical stages, side withdraw, feed locations and liquid holdup for installing catalyst are maintained same as those in the optimized RDWC process. However, the reboiler duty of the product column and the liquid flowrate to left side of the dividing-wall (FL) are considered as the decision variables to be optimized to meet the purity and impurity constraints of the product PrAc defined in Eq. 7. Besides the aforementioned two decision 24


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variables, there are three remaining operating conditions that should be determined for the optimization of VRHP-RDWC process: the overhead vapor split ratio (Rv, the fraction of the overhead vapor stream directly going to the decanter after condensing), pressure ratio of the compressor (Rcp, the ratio of discharge pressure to suction pressure) and the temperature approach of the heat exchanger 3 (ΔThex3). Here, ΔThex3 refers to the temperature difference between the hot stream inlet and cold stream outlet of the heat exchanger 3 in Figure 10. Since the temperature difference between the hot stream outlet and cold stream inlet of the heat exchanger 3 is much larger than 5°C, only ΔThex3 for the heat exchanger 3 needs to be determined. Moreover, ΔThex3 is limited by the suction temperature of the compressor because higher suction temperature will lead to higher Tdis. Thus, ΔThex3 is optimized to satisfy the constraints on Tdis and the minimum temperature approach of 5°C. Here, the temperature approach of other two heat exchangers is also set as 5°C. The temperature approach of the RD column reboiler is the temperature difference between the condensation temperature of the compressed overhead vapor stream and reboiler temperature, while the temperature approach of the product column reboiler is the temperature difference between steam condensation temperature and reboiler temperature (Eqs. A5 and A6 in Appendix). For the RD column reboiler in VRHP-RDWC process, the temperature approach was optimized to reduce the capital cost of the RD column reboiler, which has large duty and consequently large size and high cost. Here, the minimum temperature approach of the RD column reboiler is also set as 5°C and considered as the constraint expressed in Eq. 16d while the maximum temperature approach of the RD column reboiler is indirectly constrained by the maximum discharge temperature and pressure ratio of the compressor and the overhead vapor split ratio. Thus, the decision variables vector for the VRHP-RDWC optimization is 𝑥 =

[𝑅v, 𝑅cp, ∆𝑇hex3,𝑄R ― PC,𝐹L]𝑇, and its MINLP problem can be expressed as: min 𝑇𝐴𝐶(𝑥)

(16a)

𝑥

Subject to: Eq. 7

(16b)

𝑇dis ≤ 150

(16c)

𝑇RD + 5 ≤ 𝑇con

(16d)

25



Here, Tdis is the compressor discharge temperature of the compressor that should be below the allowable maximum temperature (150°C); Tcon is the condensation temperature of the compressed overhead vapor stream that should be 5°C higher than the reboiler temperature of the RD column (TRD) in RDWC to meet the required temperature driving force for heat transfer. The typical pressure ratio of the compressor is 2.5-4.0 due to the mechanical and thermodynamic efficiency limitation.24 Moreover, for assuring the controllability of the VRHP-RDWC process, the vapor split ratio, Rv assumed to be 5% or higher in the present study. These and other bounds on the decision variables are given in Table 5. The reduction of TAC with number of iterations for optimizing VRHP-RDWC process is shown in Figure 3c, which confirms the convergence of optimization method. The optimal flowsheet of VRHP-RDWC with detailed operating conditions is shown in Figure. 10. Here, 215.07 kmol/h of the overhead vapor from RDWC top is preheated from 94.13°C to 97.08°C by the condensate from the reboiler of the RD column and then compressed to 4.80 bar giving Tdis of 146.85°C that satisfies the constraint on the Tdis. The Tcon of 127.24°C is higher more than 5°C of the reboiler temperature of 117.15°C of the RD column, which satisfies the required driving force for heat transfer in the RD column reboiler. In VRHP-RDWC process, the reboiler duty of the RD column is reduced from 2432.51 kW of RDWC and 2150.67 kW of HIRDWC to 2075.04 kW due to the addition of preheaters for preheating the feed reactants. More importantly, the entire reboiler duty of the RD column in VRHP-RDWC is supplied by the hot outlet stream of VRHP, and it only needs 155.25 kW of low pressure steam for the reboiler of the product column. The acetic acid feed is preheated to 120.43°C by the condensate from the RD column reboiler while the n-propanol feed is preheated to 102.98°C by the product PrAc stream. The pressure ratio of the compressor is 4.00, which requires 357.85 kW of work input. 4.5

Comparison of Steady-State Results Detailed comparison of the economic performance and thermodynamic efficiency of the

proposed conventional RD, RDWC, HIRDWC and VRHP-RDWC processes is given in Table 6. Compared to the conventional RD process, steam consumption of RDWC, HIRDWC and VRHP-RDWC decreases by 10.08%, 19.87% and 94.60%, respectively, while the cooling

26


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water consumption decreases by 11.47%, 15.27% and 84.03%, respectively. However, VRHP-RDWC requires 357.85 kW of electricity to drive the compressor. In comparison with the conventional RD process, AOC of RDWC, HIRDWC and VRHP-RDWC decreases by 10.44%, 19.40% and 74.54%, respectively, and TCC of RDWC, HIRDWC and VRHPRDWC increases by 0.49%, 0.11% and 67.92%, respectively, which is due to the addition of heat exchangers and compressor. The RD column diameter (1.27 m) in RDWC is slightly larger than that (1.24 m) in the conventional RD process, which leads to small increase in TCC. Overall, TAC of RDWC, HIRDWC and VRHP-RDWC decreases by 6.68%, 12.7% and 25.59%, respectively, than that of the conventional RD process. These results show the great potential of the VRHP-RDWC process for producing PrAc in terms of the reduction of the operating cost and capital cost. Table 6. Comparison of costs and thermodynamic efficiency of RD, RDWC, HRDWC and VRHP-RDWC processes for producing n-propyl acetate Item (103 $/year) Column Shells Trays Condensers Reboilers Decanter Heat exchangers Compressor Catalyst Steam Cooling water Electricity AOC TCC TAC WSEP (MJ/h) Exloss (MJ/h) η (%)

RD 908.29 94.84 224.55 456.52 222.19 0.00 0.00 86.37 1100.57 27.11 0.00 1214.00 (0%) 1906.39 (0%) 1849.51 (0%) 270.20 2649.52 9.25

RDWC 1023.73 99.98 171.23 419.27 201.54 0.00 0.00 73.67 989.62 24.00 0.00 1087.29 (−10.44%) 1915.75 (+0.49%) 1725.88 (−6.68%) 261.73 2366.17 9.96

HRDWC 955.76 90.62 166.90 391.18 201.54 102.42 0.00 73.67 881.84 22.97 0.00 978.49 (−19.40%) 1908.41 (+0.11%) 1614.62 (−12.70%) 360.59 1962.19 15.52

VRHP-RDWC 957.49 90.94 73.37 896.41 200.51 150.35 832.19 73.67 59.38 4.33 171.77 309.14 (−74.54%) 3201.27 (+67.92%) 1376.23 (−25.59%) 360.66 1052.17 25.53

Moreover, the thermodynamic efficiency of the conventional RD, RDWC, HIRDWC and VRHP-RDWC processes is 9.25%, 9.96%, 15.52% and 25.53%, respectively, demonstrating that the integration of VRHP with RDWC can significantly improve the thermodynamic efficiency of the RD process. The improvement of the thermodynamic efficiency for RDWC is mainly due to the elimination of the remixing effect occurring in the

27



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RD column in conventional RD process. Moreover, the residual heat of the product PrAc and the partial condensation heat of the overhead vapor stream are recovered and used to preheat the feed streams; this heat integration reduces the reboiler duty of the RD column in HIRDWC to 2150.67 kW from 2432.51 kW in RDWC, thus improving thermodynamic efficiency of the process. For the VRHP-RDWC process, the RD column reboiler duty is supplied by the heat pump, which requires only 357.85 kW of external work input for the compressor and also results in requiring only 155.25 kW of steam for the product column reboiler. This steam quantity is much less than that by RDWC and HIRDWC processes, which gives significantly higher thermodynamic efficiency of VRHP-RDWC. For further comparing the energy consumption of the various processes, by considering the thermoelectric efficiency, we assume that the required external work input of the compressor (Qcom) can be converted into the equivalent thermal energy by a factor 3.1 Hence, the equivalent total thermal energy consumption of the reboilers for VRHP-RDWC process (QR-VRHP) can be calculated by (17)

𝑄R ― VRHP = 𝑄R ― PC + 3𝑄com

Here, QR-PC is the reboiler duty of the product column in VRHP-RDWC. Thus, the total equivalent thermal energy for the reboilers of the VRHP-RDWC is 1228.80 kW, which further indicates that even though VRHP-RDWC requires 357.85 kW of electricity to drive the compressor motor, its total thermal energy consumption is still much smaller than 2879.08 kW for conventional RD, 2587.49 kW for RDWC, 2304.78 kW for HIRDWC. This further conforms the higher thermodynamic efficiency of VRHP-RDWC than other processes in Table 6.

5

PLANT-WIDE CONTROL SCHEME In this section, controllability of the proposed VRHP-RDWC process is tested and

evaluated using PI controllers by introducing disturbances in feed flowrate and composition. The sumps of the columns for RDWC were sized to give 10 min liquid holdup with 50% liquid level. However, the decanter was sized to provide 20 min liquid holdup with 50% liquid level for good separation of the two liquid phases. After adding the required pumps and

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valves into the steady-state flowsheet, Aspen Plus model was exported to Aspen Dynamics 9.0 to obtain the dynamic model of the process using pressure-driven simulation. 5.1

Design of the Overall Control Structure The overall plant control structure of a process normally consists of two levels; the first

level is the material balance control including pressure, flow and level control loops while the second level is the quality control that is used to maintain the product purity. The composition controller can maintain the product purity at the desired value without any offset; however, in industrial practice, tray temperature controller is commonly used in the quality control loop instead of composition controller due to the large time delay, higher capital and maintenance cost of the composition analyzer. There are two important control degrees of freedom that have significant effect on the product purity and energy consumption of the VRHP-RDWC process. One of them is the stoichiometric balance of feeds/reactants that can be manipulated by the sensitive tray temperature of the RD column. The other one is the liquid split ratio (βL) above the dividing wall that can be used to minimize the energy consumption for the operation of RDWC. Lee at al.48 proposed a control structure for a thermally-coupled RD process for the esterification of acetic acid and isopropanol, wherein the feed ratio is manipulated by the sensitive tray temperature of the RD column while βL is fixed. Chen et al.49 investigated the control performance of a thermally-coupled RD process for producing methyl valerate. In their control strategy also, the feed ratio is manipulated by the sensitive tray temperature of the RD column; however, the reflux rate of the RD column is ratioed to the sum of the two feed flowrates (FL/SF) instead of the fixed βL. Their results demonstrate that the product purities can be controlled at their respective specifications with smaller offsets using both control schemes with fixed βL or FL/SF; however, relatively large transient deviations occur in the dynamic responses of the product purities for the control strategy proposed by Chen et al..49 By referring to the previous relevant works,48,49 the sensitive tray temperature of the RD column of RDWC is used to manipulate the molar flowrate ratio of the n-propanol feed to acetic acid feed (FR) while the sensitive tray temperature of the product column of RWDC is used to manipulate its reboiler duty (QR-PC). The open-loop sensitivity analysis is

29



implemented to select the sensitive tray temperatures of the RD and product columns by introducing ±0.1% step change in each manipulated variable while keeping the others at their nominal steady-state values. Plots a and b in Figure 11 show these sensitive analysis results for ±0.1% step change in FR and QR-PC, respectively. The temperature of 18th stage of the RD column (TR18) is most sensitive to FR while the temperature of 6th stage of the product column (TP6) is most sensitive to QR-PC. Therefore, TR18 and TP6 can be used to manipulate FR and QRPC,

respectively.

Figure 11. Open-loop sensitivity test results for ± 0.1% step change in (a) FR and (b) QR2

Subsequently, two control schemes with fixed βL (CS1) or FL/SF (CS2) are designed and evaluated in the following section. The overall control schemes are shown in Figure 12, wherein the green dashed line shows the control scheme with fixed βL while the red dashed line shows the control scheme with fixed FL/SF; blue dashed line is applicable to both CS1 and CS2. All control loops in Figure 12 are specified as below. (1)The feed flowrates of n-propanol (FC1) and acetic acid (FC2) are flow controlled with a ratio FR of 1.036 at the nominal steady-state. (2)The operating pressure of RDWC is controlled at 1.2 bar by manipulating the vapor

30


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flowrate to Condenser 1 (PC1); this will guarantee the compressor inlet pressure at the desired value. (3)The discharge pressure of the compressor is controlled at 4.8 bar by manipulating the opening of the pressure-relief valve with direct action (PC3). Thus, suction and discharge pressures of the compressor are maintained at their specified values, and the condensing temperature of the compressed vapor stream can remain at the desired value to satisfy the required temperature driving force for heat transfer in the reboiler of the RD column.

Figure 12. Overall control structure of VRHP-RDWC process (4)The decanter pressure is controlled at 1.19 bar by manipulating the duty of Condenser 2 (PC2). (5)The organic phase level in the decanter is controlled by the reflux rate of the organic phase to the top of RDWC (LC3) while the aqueous phase level is controlled by the draw-out water flow (LC4). (6)The sump level in product column is controlled by PrAc flow (LC2).

31



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(7)The sump level in the RD column is controlled by manipulating the compressor power (LC1), since there is no outflow from the RD column bottom and the reboiler duty is indirectly supplied by the input work of the compressor. (8)The TR18 is controlled at 106.81°C by manipulating FR with reverse action (TC1) while the TP6 is controlled at 104.02°C by manipulating QR-PC with reverse action (TC2). (9)Temperature of the overhead vapor after condensing is controlled at 82.67°C by manipulating the duty of Condenser 1 (TC3). In the present study, proportional-integral (PI) controllers with proportional gain of 𝐾c = 20 and integral time of 𝜏I = 12 min are used in all pressure control loops while PI controllers with 𝐾c = 0.5 and 𝜏I = 0.3 min are used in all flow control loops. However, for level control loops, P-only controllers with 𝐾c = 2 are used. Moreover, PI controllers are used in all temperature control loops, which are tuned by closed-loop ATV test and the tuning parameters (listed in Table 7) are calculated by the Tyreus-Luyben PI tuning rules,50 all carried out within Aspen Dynamics. Table 7. Parameters of the temperature controllers in VRHP-RDWC process Control loop Controlled and manipulated variables

5.2

Gain, Kc

Integral time, τI (min)

TC1

TR18 and FR

4.70

118.93

TC2

TP6 and QR-PC

4.78

7.66

TC3

TCond1 and QCond1

0.44

0.264

Dynamic Performance Evaluation Changes in feed flowrate and impurity of reactants are the main disturbances for the

operation of VRHP-RDWC process. Therefore, in the present study, control performance of the proposed control structures was tested and evaluated by introducing each of these disturbances at the time of one hour. The simulation time is set as 30 hours for all tests. The composition responses are plotted for the entire simulation time due to their slow dynamics. Responses of the remaining variables are fast and hence they are plotted until 12 hours only, to clearly compare the transient performance under various disturbances.

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Figure 13 shows the dynamic responses of CS1 (with fixed βL) and CS2 (with fixed FL/SF) under ±20% changes in n-propanol feed flowrate (FPrOH ± 20%). It can be observed that these responses are relatively slow (requiring about 30 hours for settling), since the esterification reaction is typically a slow reaction. The purity of the product PrAc (xPrAc) and byproduct water (xH2O) takes about 30 hours to reach the new steady state with about ±0.1% offset when the n-propanol feed flowrate disturbance occurs. Moreover, acetic acid impurity in the product PrAc (xAA-PrAc) is maintained close to 100 ppm after the system reaches the new steady state. Flowrate of the product PrAc (FPrAc), byproduct water (FH2O), the reflux rate to the RD column top (FL) and acetic acid feed (FAA) increase or decrease simultaneously when the n-propanol feed flowrate increases or decreases, which indicate that both control schemes can drive the system to meet the optimal stochiometric feed balance.

Figure 13. Dynamic responses under n-propanol flow rate (FPrOH) disturbances 33



Figure 13j shows that the condensing temperature of the compressed vapor stream (Tcon) returns to the initial temperature after small transient oscillations when the n-propanol feed flowrate disturbance occurs. This demonstrates that both the proposed control schemes can maintain the required temperature driving force for heat transfer in the reboiler of the RD column. The dynamic responses of CS2 display large transient deviations and oscillations compared to the dynamic responses of CS1. This is because the control structure with fixed FL/SF (CS2) can respond fast to the feed flowrate change, which results in larger transient deviations. On the other hand, the control structure with fixed βL (CS1) is not sensitive to the feed flowrate changes, thus resulting in relatively smooth responses for rejecting feed flowrate disturbances. Therefore, even though CS1 displays slower dynamic responses, it still shows better control performance than CS2 for rejecting the feed flowrate disturbances. Next, the performance of the proposed control schemes under the composition disturbance of acetic acid feed is studied. Figure 14 shows the dynamic responses of CS1 and CS2 when the mole fraction of acetic acid in feed stream (xF,AA) increases from 0.95 to 1.0 (i.e., no water) or decreases from 0.95 to 0.9 (i.e., water impurity increases). When the system reaches the new steady state, the PrAc purity can return to the desired value with a smaller offset (maximum of 0.01 mol%) while acetic acid impurity in the product PrAc is close to the specified maximum of 100 ppm. Dynamic responses of the purity of the product PrAc and byproduct water for CS1 display relatively large transient deviations but with smaller offsets, compared to those for CS2. It can also be seen in Figure 14 that the feed flowrate of acetic acid (FAA) smoothly decreases or increases, to meet the required stochiometric feed balance, resulting in the byproduct water flowrate decrease or increase simultaneously. Moreover, the product PrAc flowrate returns to its initial value after oscillating with a small amplitude, since the feed flowrate of n-propanol is fixed. The reflux rate of the RD column (FL) also changes and reaches a new steady state under the composition disturbances in acetic acid feed. The system reaches a new optimal steady-state feed ratio to meet the stoichiometric balance of the feed reactants, which leads to the change in the total feed flowrate, and so FL will be changed. An interesting observation is that CS1 and CS2 give slightly different power (Qcom) and steam

34


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consumptions (QR-PC) (Figures 14k and 14l) although the obtained product purity is not much different, which are mainly caused by different βL after the system reaches a new steady state. This observation further confirms that βL has significant effect on the energy consumption of RDWC, which can be optimized by using a multivariable control strategy such as model predictive control to achieve optimal operation.

Figure 14. Dynamic responses under acetic acid mole fraction changes in feed stream (xF,AA) from 0.95 to 1 and 0.95 to 0.9 In addition, the low purity n-propanol is easier to obtain and is cheaper than high purity n-propanol in industrial production. Therefore, in this section, the performance of the proposed control schemes for rejecting large water impurity in n-propanol feed was investigated by changing the n-propanol mole fraction in feed stream (xPrOH) from 0.995 to 0.9.

35



The dynamic responses of CS1 and CS2 in Figure 15 demonstrate that the purity of the product PrAc and the byproduct water slightly increase for both control schemes when the mole fraction of water in n-propanol feed increases; increase in water purity is mainly due to the liquid-liquid equilibrium in decanter. Moreover, the increase of the purities of product PrAc and byproduct water for CS2 is slightly larger than that for CS1, this is because more liquid returns to the RD column as the reflux for CS2 than that for CS1, which makes more water go up to the top of the RD column in the form of the binary azeotrope of PrAc-water. Consequently, more n-propanol can be converted into product PrAc, leading to the decrease of the PrOH mole fraction in the overhead vapor stream. The similar phenomenon was also observed in the dynamic responses of a thermally coupled RD process for producing isopropyl acetate via the esterification of acetic acid with isopropanol.48

Figure 15. Dynamic responses under the mole fraction change of n-propanol in feed stream from 0.995 to 0.9

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Moreover, Figure 15 also shows that when the water impurity in n-propanol feed increases, acetic acid impurity in the product PrAc remains below the specified maximum of 100 ppm, while the flowrate of the product PrAc and the acetic acid feed decrease simultaneously due to the required stoichiometric feed balance and the lower flowrate of pure n-propanol into the system. The required compressor power and steam also decrease in both CS1 and CS2 due to the reduction in the flowrate of the feed reactants. While CS1 requires lesser electricity for driving the compressor than CS2 (333.63 kW for CS1 versus 335.13 kW for CS2); CS1 requires more low pressure steam to supply the thermal energy for the product column reboiler than CS2 (0.537 GJ/h for CS1 versus 0.525 GJ/h for CS2, see profiles of Qcom and QR-PC in Figure 15) when the system reaches a new steady state. The main reason for this observation is that the reflux rate to the RD column (FL in Fig. 15) in CS1 is smaller than that in CS2, leading to lower electricity requirement and larger steam consumption in CS1. In other words, the βL value in CS1 is smaller than that in CS2, which furtherly indicates that the βL value has significant effect on the energy consumption for the RDWC operation. Overall, although more steam is required in CS1 than that in CS2, CS1 is still better for the VRHPRDWC operation because the cost of total steam required is much smaller than the cost of electricity required for the operation of the proposed VRHP-RDWC. Figure 16 shows the dynamic responses of temperature controllers for 1°C increase and decrease in their set points as well as the corresponding dynamic responses of product purity. It demonstrates that both the temperature controllers of TR18 and TP6 can effectively track their set point changes, and the control strategies of CS1 and CS2 can bring the process to the new steady state within 30 hours. In particular, TR18 controller needs long time to reach steady state because the esterification of acetic acid and n-propanol is a slow reaction. The product PrAc purity is sensitive to the change in TP6 with faster dynamics but not to the change in TR18. In contrast, the byproduct water purity is not sensitive to the change in TP6 but to the change in TR18. This is because the byproduct water purity is mainly determined by the liquid-liquid equilibrium in the decanter, and changes in TR18 lead to composition variation of the overhead vapor stream that affects the liquid-liquid equilibrium in decanter. Furthermore, TP6 is the sensitive tray temperature of the product column and can be used to control the PrAc product

37



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quality. In summary, performance results in Figure 16 demonstrate that the PrAc product purity can be effectively controlled by the proposed control strategy.

Figure 16. Dynamic responses of TR18 and TP6 temperature controllers for 1°C increase and decrease in their respective set point, and corresponding dynamic responses of product purity

The integral of absolute error (IAE) is used as the quantitative criterion to evaluate performance of the two control schemes, which can be calculated by: 𝑇

IAE =

∫ |𝑒(𝑡)|𝑑𝑡

(15)

0

Here, e(t) is the deviation of the controlled variable from its set point and T is the total simulation time. The IAE values at T = 30 h (including one hour before the disturbance was introduced) for the dynamic responses of the purity of the product PrAc and the byproduct water are 38


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summarized in Figure 17, which shows that the IAE values for the flowrate disturbances of npropanol feed (FPrOH) is much larger than those for feed composition disturbances. This because large change in n-propanol feed flowrate leads to relatively large offsets (panels a and b in Figure 13), which result in larger IAE values. Moreover, as shown in Figure 17, IAE values for acetic acid feed composition (xF,AA) disturbances are much smaller than those for npropanol feed composition (xF,PrOH). The two control schemes give similar control performance for the disturbances in n-propanol feed flowrate and increase of water in acetic acid feed. However, IAE values for CS1 are lower than those for CS2 under the disturbances of acetic acid feed purity and water impurity in n-propanol feed. Although the steady state absolute deviation of product PrAc purity (Figure 14a) in CS2 is smaller than that in CS1, the dynamic responses in CS2 exhibit larger oscillations than those in CS1, which leads to smaller IAE values for CS1 than those for CS2.

Figure 17. IAE values of the product PrAc purity under various disturbances tested Overall, the above findings demonstrate that the VRHP-RDWC can be operated smoothly under various disturbances including large increase of water impurity in the feeds, by using both the control schemes. Compared to CS2, CS1 is more suitable for the operation of VRHP-RDWC due to its attractive potential for energy savings and smooth dynamic response.

6

CONCLUSIONS In this study, four different reactive distillation (RD) processes, i.e. conventional RD,

reactive dividing-wall column (RDWC), heat integrated RDWC (HIRDWC) and vapor recompression heat pump assisted RDWC (VRHP-RDWC) were systematically designed and compared for producing n-propyl acetate via the esterification of acetic acid with n-propanol.

39



Compared to the conventional RD process, the annual operating cost of RDWC, HIRDWC and VRHP-RDWC processes decreases by 10.44%, 19.40% and 74.54%, respectively, while the total capital cost of RDWC, HIRDWC and VRHP-RDWC increases by 0.49%, 0.11% and 67.92%, respectively. Overall, the total annual cost of RDWC, HIRDWC and VRHP-RDWC decreases by 6.68%, 12.70% and 25.59%, respectively, compared to that of the conventional RD process. Also, thermodynamic efficiency of conventional RD, RDWC, HIRDWC and VRHP-RDWC is 9.25%, 9.96%, 15.52% and 25.53%, respectively, demonstrating that VRHP-RDWC is a very promising, energy saving process for producing n-propyl acetate via the esterification of acetic acid with n-propanol. Moreover, two alternative control strategies are proposed for the operation of VRHPRDWC, wherein, the sensitive temperature of the RD column is used to manipulate feed ratio in order to maintain the optimal stoichiometric balance of the feed reactants, while the sensitive tray temperature of the product column is used to manipulate its reboiler duty in order to maintain n-propyl acetate purity. Simulation results show that the control strategy with fixed liquid split ratio (CS1) provides better performance than the control strategy with fixed flowrate ratio of reflux rate to the sum of the two feed reactant flowrates (CS2), in terms of smoother dynamic response and lower energy consumption of VRHP-RDWC process.

ACKNOWLEDGMENTS This work is financially supported by the National Natural Science Foundation of China (Nos. 21776025, 21606026, 21878028), the Fundamental Research Funds for the Central Universities (No. 106112017CDJQJ228809), and the Chongqing Research Program of Basic Research and Frontier Technology (No. CSTC2016JCYJA0474). The authors acknowledge the China Scholarship Council (CSC file no: 201706050043) for providing scholarship to Feng Z. to conduct research at the National University of Singapore.

CORRESPONDING AUTHORS G.P. Rangaiah, Email: [email protected] Lichun Dong, Email: [email protected] Notes: 40


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The authors declare no competing financial interest.

APPENDIX Stainless steel is used as the material of all equipment in the present study. 1. Height of Column (A1)

𝐿𝐶 = 1.2 × 𝐻s × (𝑁T ― 1)

Here, Hs [m] is the tray space, and NT is the total number of theoretical trays including its reboiler but excluding condenser. 2. Cost of Column Shell Column shell cost [$] =

M&S × 937.636 × 𝐷1.066 × 𝐿0.802 × (2.18 + 3.67) 𝐶 𝐶 280

(A2)

Here, DC [m] is the diameter of column. 3. Cost of Column Tray Column tray cost [$] =

M&S × 92.473 × 𝐷1.55 × 𝐿𝐶 × (1.0 + 0.0 + 1.7) 𝐶 280

(A3)

4. Reboiler Heat Exchanger Area 𝐴𝑅 [m2] =

𝑄R

(A4)

𝑈R × ∆𝑇R

Here, QR [kW] is the reboiler duty; the heat transfer coefficient of the reboilers, UR is 580 W/(K∙m2). When low-pressure steam at 160oC is used for thermal energy, ΔTR is given by (A5)

∆𝑇R = 160 ― 𝑇R

Here, TR is the reboiler temperature (i.e., the boiling temperature of the liquid in the reboiler). For the VRHP-RDWC process, ΔTR is calculated by (A6)

∆𝑇R = 𝑇con ― 𝑇R

Here, Tcon is the condensation temperature of the compressed overhead vapor stream. Note that steam condensation temperature, TR and Tcon remain constant throughout the reboiler, and hence log mean temperature difference simplifies to equations A5 and A6. 5. Heat Exchanger Area of Condenser and Heat Exchangers 𝑄c 𝐴𝑐 [m2] = 𝑈c × ∆𝑇c

(A7)

Here, QC [kW] is the heat duty of heat exchangers/condenser, the heat transfer coefficient, UC is assumed as 840 W/(K∙m2), and the ΔTC is given by 41



∆𝑇C =

Page 42 of 48

(𝑇hin ― 𝑇cout) ― (𝑇hout ― 𝑇cin) ln

(

𝑇hin ― 𝑇cout

)

(A8)

𝑇hout ― 𝑇cin

Here, Thin and Thout are respectively the inlet and outlet temperatures of the high temperature stream; and Tcin and Tcout are respectively the inlet and outlet temperatures of the low temperature stream. For the cooling water, Tcin and Tcout are respectively 30oC and 40°C. 6. Cost of Heat Exchangers Heat exchanger cost [$] =

M&S × 474.668 × 𝐴0.65 × (2.29 + 𝐹𝑐) 280

(A9)

Reboilers are assumed as kettle reboiler with 𝐹𝑐 = 1.35 × 3.75, and condensers and other exchanger are assumed as fixed-tube heat exchanger with 𝐹𝑐 = 0.8 × 3.75. 7. Decanter Cost Decanter cost [$] =

M&S × 937.636 × 𝐷1.066 × 𝐿0.802 × (2.18 + 3.67) d d 280

(A10)

Here, diameter (Dd) and length (Ld) of the decanter are given by 𝑉d [m ] = 2 × 𝐹v × 20/60 = 𝜋 × 3

𝐷2d 4

× 𝐿d

(A11)

Here, Vd is the volume of the decanter which is calculated by giving residence time of 20 min for the two-liquid phase separation, and Fv [m3/h] is the total flowrate of the inlet streams. Ld is assumed as three times the Dd. 8. Compressor Cost Compressor cost [$] =

M&S × 410.852 × 𝑄0.82 com × (2.18 + 1) 280

(A12)

Here, Qcom [kW] is the compressor horse power. 9. Steam Cost (A13)

Steam cost [$/year] = 𝐶s × 𝑄R × 8000

Here, CS [$/GJ] is the price of steam, QR [GJ/h] is the reboiler duty and 8000 is the number of operation hours per year. 10. Cooling Water Cost Cooling water cost [$/year] = 𝐶w × 𝑄C × 8000 Here, Cw [$/GJ] is the price of cooling water and QC [GJ/h] is the condenser duty. 11. Electricity Cost 42


(A14)

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Electricity cost [$/year] = 𝐶e × 𝑄com × 8000

(A15)

Here, Ce [$/kWh] is the price of electricity and Qcom [kW] is the compressor horse power. 12. Catalyst Cost Catalyst cost [$/year] = catalyst loding [kg] × 7.7162 × 4

(A16)

Here, 7.7162 $/kg is the cost of catalyst (Table 4) and 4 is the number of catalyst replacements per year (for the assumed catalyst life time of 3 months).

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Process Development, Assessment and Control of Reactive Dividing

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