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Article Cite This: Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

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Reaction−Separation−Recycle Processes for 2‑Ethylhexyl Acrylate Production: Design, Control, and Economic Evaluation Mihai Daniel Moraru*,†,‡ and Costin Sorin Bildea‡ †

Department of Process Technology and Development, Hexion, Seattleweg 17, 3195 ND, Pernis, The Netherlands Department of Chemical and Biochemical Engineering, University Politehnica of Bucharest, Str. Gh. Polizu 1-7, 011061, Bucharest, Romania

‡

S Supporting Information *

ABSTRACT: 2-Ethylhexyl acrylate is commercially produced from acrylic acid and 2-ethylhexanol using sulfuric acid as catalyst. Employing solid catalysts can eliminate corrosion, increase selectivity, and make easier the separation and product purification. Process development studies for a complete plant have not been reported in the literature. In this paper, the design, control, and economic evaluation of three reaction−separation−recycle processes are developed. The reaction takes place in a fixed-bed reactor using Amberlyst 70 as catalyst. The acrylate is recovered by distillation at the desired purity. One of the processes obtains high-purity wastewater by removing the water in a distillation−decanter system. The other two obtain low-purity water in a flashdecanter, and a decanter-only system, respectively. All three processes are controllable, the control system showing robustness when an increase or decrease in production capacity is required, or when the fresh reactants become contaminated. The economic analysis shows attractive economic potential and other key economic indicators.

1. INTRODUCTION 2-Ethylhexyl acrylate (2-EHA) is an important bulk chemical, generally used as precursor in the production of acrylic polymers. This ester has similar uses as the n-butyl acrylate, finding its end-use applications in a variety of products as solvent- and water-based paints, adhesives, paper and textiles,1 and impregnating agents.2 Industrially, 2-EHA is produced by esterification of acrylic acid (AA) and 2-ethylhexanol (2-EH), with formation of water as byproduct. The reaction is performed batchwise (3 to 5 h reaction times) in the presence of an organic solvent and sulfuric acid as catalyst, followed by product separation and recovery in a distillation system.1 The use of strong homogeneous catalysts, however, has serious negative consequences as corrosion, loss of catalyst, and pollution of environment,3 and high downstream operation costs.2 Usually, all these lead to high maintenance costs and a continuously increasing difficulty to comply with environmental regulations. Therefore, using acidic solid catalysts eliminates corrosion, increases selectivity, and makes easier the acrylate separation and recovery.3 Recent publications2−5 studied the reaction © XXXX American Chemical Society

between AA and 2-EH on solid catalysis, using various ion exchange resins. From a process design and operation perspective, and despite its wide use as a large bulk chemical, process development studies for 2-EHA production using solid catalysts have not been reported in the literature. To fill this gap, the design, control and economic evaluation of three reaction-separationrecycle (RSR) processes are developed and presented in this study. The estimation of reaction kinetic parameters using experimental data reported in the literature by Komon et al.3 is made first. Then, the thermodynamic model and the key physical properties are presented. On the basis of a kinetic reactor model, a black-box separation section and recycle of unconverted reactants, a preliminary analysis of the RSR system is made. Considering that the system has one recycle (both reactants recycled together), a plantwide control structure is Received: November 16, 2017 Revised: January 31, 2018 Accepted: January 31, 2018

A

DOI: 10.1021/acs.iecr.7b04752 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

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Industrial & Engineering Chemistry Research

The experiments were performed in a batch reactor with a capacity of 100 mL. In a typical experiment, known amounts of AA and 2-EH were used. Each time, 5% mass catalyst was used. The evolution in time of 2-EHA mole fraction was read from Figures 7 and 8 of that paper (see the Supporting Information for the tabular data). Other details regarding these experiments can be found in the original reference. 2.1.2. Regression of Kinetic Parameters. As mentioned, Komon et al.3 used the UNIFAC model to calculate the liquid phase activities, while in this study the UNIQUAC model is employed (see section 2.2). The activities should be calculated by the same thermodynamic model used while deriving the kinetic parameters. As we calculate the activities with the UNIQUAC model, the same should be employed when kinetic parameters are regressed; thus, using the kinetic parameters determined using the UNIFAC model (in the original study3) and calculating the activities using UNIQUAC (in this study) is not consistent. Therefore, the kinetic parameters k0 and EA, in units of [kmol/(kgcat·s)] and [kJ/mol], respectively, and the two constants A [−] and B [K] in eq 5 (i.e., logarithm of the invers of the equilibrium constant), are regressed from the experimental data reported by Komon et al.3 In addition, one can remark in the units of r in eq 2, which has the same units as k0, the use of [kgcat]. Determining r in units of [kmol/(kgcat·s)] (i.e., per mass of catalyst) is very useful since this expression can be immediately used to obtain the component consumption/formation rates (Rj, in [kmol/s]) by multiplication of the reaction rate with the stoichiometric coefficient of component j (νj) and the mass of catalyst (mcat, in [kgcat]) in the reactor (eq 6). Note that mcat is the variable which determines the component consumption/formation rates.

proposed, and the steady state behavior is analyzed using bifurcation diagrams. The selection of an initial operating point provides the mass balance for the synthesis of the three separation systems. The nominal operating point of each process is selected based on rigorous sensitivity analyses. After sizing the relevant equipment, the dynamic behavior is analyzed when subjected to various process changes. The study ends with an economic evaluation, process comparison and general conclusions. The proposed plant has a capacity of 20 000 t/a of 2-EHA. This plant capacity was proposed by Niesbach et al.6 for a similar higher-acrylate process (i.e., n-butyl acrylate); for the same acrylate, Constantino, D.S.M., et al.7,8 proposed plant capacities in the same order of magnitude. The product specifications are ≥99.5% 2-EHA, ≤ 0.05% water, and ≤0.01% AA (all by mass), while the rest is 2-EH, which accounts as well for other impurities. Aspen Plus V8.4 (AP), Aspen Batch Modeler V8.4 (ABM), and Aspen Plus Dynamics V8.4 (APD) are used as efficient Computer-Aided Process Engineering (CAPE) tools to perform the analysis.

2. REACTION KINETICS AND THERMODYNAMICS Komon et al.3 determined the reaction kinetics making use of the UNIFAC (universal quasichemical functional-group activity coefficients) model to calculate the liquid-phase activities. In this study, the UNIQUAC (universal quasichemical) model is used. This thermodynamic model must be consistent with the thermodynamic model that was used when regressing the kinetic parameters. For this reason, a new set of kinetic parameters is regressed (see section 2.1.2) using the literature experimental data presented by Komon et al.3 (see section 2.1.1). 2.1. Reaction Kinetics. 2.1.1. Literature Experimental Data. Equation 1 describes the liquid phase esterification reaction of AA and 2-EH with formation of 2-EHA and water. Aspen Plus calculates an enthalpy of reaction of 21 kJ/mol (liquid phase reaction, 25 °C). The course of reaction was studied by Komon et al.3 For reaction on Amberlyst 70, that paper reports experimental data on 2-EHA formation, as well as regression of kinetic parameters for the pseudohomogeneous model described by eqs 2−4. Note that in our work, the form of the equilibrium constant described by eq 4 is replaced by the one described by eq 5.

r = k f (aacidaalcohol − (1/Keq)aestera water)

(2)

k f = k 0exp(−EA /(RT ))

(3)

Keq = exp(b1 + b2 /T + b3T )

(4)

ln(1/Keq) = A + B /T −1

R j = νjrmcat

(6)

ABM is used to make the regression. The setup of the regression case is as follows. Conveniently, and for maintaining consistency, a property file is generated using AP and imported in ABM. Then, the reaction stoichiometry is included, and the reaction is specified as kinetic using the power law-type kinetic expression. This consists of the forward reaction rate constant and driving force. Thus, the values of all parameters of the rate expression are introduced. The setup continues with introducing the experimental data. A typical entry requires the reactor conditions (initial charge: temperature, pressure, amount, and component composition of the reaction mixture AA + 2-EH) and the mole fraction in time of 2-EHA. The amount of catalyst is specified separately, and this value holds for all experiments considered in the regression. Table 1 presents the regression results, together with the initial values from which the regression calculation starts. The initial values are derived from rough back-calculations using data reported in the same reference. Figure 1 shows the comparison between the experimental and calculated molar fraction of 2-EHA, at different temperatures and initial molar ratios AA:2-EH. A very good agreement is observed.

(5) −1

where r [min ] is the reaction rate, kf [min ] is the forward rate constant, ai, (i = acid, alcohol, ester, water) are the liquid phase activities, while Keq [−] is the equilibrium constant. k0 [min−1] is the pre-exponential factor in the Arrhenius equation, and EA [kJ/mol] is the activation energy. b1 [−], b2 [K], and b3 [K−1] are constants in the equilibrium equation eq 4. A [−] and B [K] are constants in the equilibrium equation eq 5.

Table 1. Regressed Kinetic Parameters Using Experimental Data from Komon et al.3

B

parameter

k0/[kmol/(kgcat·s)]

EA/[kJ/mol]

A/[−]

B/[K]

initial guess regressed

500.0 722.7

50.12 51.77

−12.591 −8.5845

2923.4 2438.5


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Figure 1. Comparison between experimental (markers) and calculated (lines) data of 2-EHA mole fraction in time, for different temperatures and initial AA:2-EH mole ratios; the experimental data are taken from the paper of Komon et al. Figure adapted with permission from ref 3. Copyright 2013 Elsevier.

Table 2. Binary Interaction Parameters for the UNIQUAC Activity Coefficient Model component i

water

water

AA

water

AA

2-EHA

component j

AA

2-EH

2-EH

2-EHA

2-EHA

2-EH

sourcea

APV84 VLE-HOC

APV84 LLE-ASPEN

APV84 VLE-HOC

R-PCES

R-PCES

R-PCES

ai,j/[−] aj,i/[−] bi,j/[°C] bj,i/[°C]

0 0 −425.922 240.834

−0.7007 0.7694 −42.798 −518.336

0 0 −459.914 256.560

0 0 −214.978 −421.744

0 0 100.073 −212.211

0 0 −26.497 −17.605

a

The source refers to the databank name from where the binary interaction parameters are taken (i.e., APV84 VLE-HOC or APV84 LLE-ASPEN), or to the model that estimated the binary parameters (i.e., R-PCES, the Property Constant Estimation System of Aspen).

simulation. This method uses the UNIQUAC activity coefficient model for describing the liquid phase behavior, while the vapor phase is described by the HOC (HaydenO’Connell) equation of state model. Pure component physical property parameters are all available in the Aspen Plus databanks. Not all binary interaction

It is worth mentioning that the activation energy of this reaction is only 3.3% higher than that reported by Komon et al.3 2.2. Thermodynamics. 2.2.1. Thermodynamic Method. The UNIQ-HOC method in Aspen Plus (default naming) is used to calculate all properties required for the process C


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3. REACTION−SEPARATION−RECYCLE SYSTEM 3.1. System Structure. Figure 3 presents the reaction− separation−recycle (RSR) structure of the plant. The fresh feeds of 2-EH (F2‑EH,0) and AA (FAA,0), together with the recycle stream of unconverted reactants, are mixed and fed to a fixed-bed catalytic reactor operated adiabatically. The reactor effluent is sent to a separation section where product 2-EHA, byproduct water, and the recycle streams are obtained. The plant has one recycle, the reactants AA and 2-EH being recycled together. This choice is based on the boiling temperature of the singular points present in the system: the reactants have the boiling temperature between the lightest group of minimum boiling heterogeneous azeotropes (including water) and the product 2-EHA (Table 3 and Table 4). 3.2. Sensitivity Analyses. 3.2.1. Basis of Analyses. The sensitivity analyses presented in this section make use of a steady state model created in Aspen Plus. The reaction section is rigorously modeled using the reaction kinetics previously determined (section 2.1.2), while the separation section uses an ideal separation model with given separation factors. The reactor is represented by the RPLUG block (i.e., a plug-flow reactor model), which assumes perfect radial mixing and neglects the axial mixing. Given the reactor sizing (i.e., the mass of catalyst), the operating policy (adiabatic), the reaction kinetics, and the inlet stream specifications (i.e., temperature, pressure, and alcohol to acid molar flow rates), the RPLUG block solves rigorous mass and energy balance equations, together with phase-equilibrium relationships to calculate the condition of the outlet stream. The separation is modeled by the SEP block, which simply distributes the components present in the inlet stream to several outlet streams (recycles and products), according to the component-specific separation factors. The mass balance of fresh reactants and recycle is performed by the MIXER block. 3.2.2. Conversion and Recycle Composition. To analyze the influence of AA conversion and recycle composition on process performance, some preliminary specifications are made. The plant operates at 13.6 kmol/h fresh AA, setting the production capacity. The stream at the reactor inlet is set to 120 °C and a mole ratio 2-EH to AA of 3. The temperature is selected such that reduces the reactor size (the higher the temperature, the larger the reaction rate), while the mole ratio is selected based on the information that polymerization of AA can be reduced when the reactor is operated at higher alcohol to AA ratios.10 The separation section is specified using the separation factors for products outlets and reactants recycle. The separation factor (β) is component specific, and is defined as the ratio between the flow rate leaving the plant as product and the flow rate in the stream entering the separation section. A separation factor of 1 means complete removal from the plant, while 0 means complete recycle of component. The first analysis investigates the influence of AA per pass conversion on reactor size (i.e., amount of catalyst) and on separation section load (i.e., recycle flow rate). Figure 4 (left) presents the mass of catalyst (mcat) and the recycle flow rate versus AA conversion (XAA), for the case of a perfect separation: complete recycle of reactants (β2‑EH = 0, βAA = 0) and complete removal of products (βwater = 1, βn‑BA = 1). As expected, more catalyst is needed to achieve higher conversion, with the benefit of lower recycle flow rate. Clearly, a trade-off exists. To find the optimum conversion value, the reactor and separation costs (both, investment and operation) are needed,

parameters for the UNIQUAC activity coefficient model are present. Three pairs (water/AA, water/2-EH, and AA/2-EH) out of the six are available, while the other three pairs (water/2EHA, AA/2-EHA, 2-EHA/2-EH) are estimated using the UNIFAC group contribution method. For the latter pairs (i.e., the binary systems containing 2-EHA), neither parameters nor phase equilibrium data are found in literature. All parameters are given in Table 2. The association parameter for AA used in the equation of state is available in Aspen. 2.2.2. Azeotropy. An azeotrope search using the Distillation Synthesis tool reveals that this four-component system has three minimum boiling heterogeneous azeotropes, water being always one of the components in that mixture. Table 3 presents the calculated azeotropes at 0.2 and 1.013 bar, together with one experimentally determined azeotrope (at Table 3. Experimental and Calculated Azeotropic Data (Mole Based) Using the UNIQUAC-HOC Model, at 0.2 and 1.013 bar Az. no.

typea

Tb/[°C]

1 2 3

Het Het Het

59.7 59.7 59.8

2 2e 3

Het Het Het

99.1 99.1 99.6

a

water

2-EH

P = 0.2 bar 0.9839 0.0121 0.9848 0.0152 0.9893 P = 1.013 bar 0.9683 0.0317 0.9679 0.0321 0.9836

2-EHA 0.0040

refa

0.0107

calcd calcd calcd

0.0164

calcd Omota et al.9 calcd

Notation: Het, heterogeneous; calcd, calculated.

1.013 bar) reported in the literature.9 All these azeotropes have almost the same boiling point, very close to that of water, being the lightest singular points in the system. The calculations predict that at 1.013 bar the ternary heterogeneous azeotrope disappears. The boiling points of pure components are included in Table 4. Table 4. Calculated Boiling Points, in °C, at 0.2 and 1.013 bar P/[bar]

water

AA

2-EH

2-EHA

0.2 1.013

60.1 100.0

95.3 141.2

134.8 184.6

159.8 216.4

2.2.3. Phase Equilibria. Both the vapor−liquid equilibrium (VLE) and liquid−liquid equilibrium (LLE) use the same sets of binary interaction parameters. Liquid−liquid heterogeneity is present only in two binary systems, namely water/2-EH and water/2-EHA. For the first system, Omota et al.9 show that using only mutual solubility and azeotropic data is sufficient to accurately describe the whole range of vapor−liquid−liquid equilibria (VLLE). Since no experimental data are available for the latter system, we accept the hypothesis that the estimated set of binary parameters can describe accurately enough the VLLE. The process synthesis of the separation systems makes use of key phase equilibrium diagrams. Figure 2A presents the ternary LLE diagrams (mass-based) at 30 °C and 1.013 bar, for the three systems containing water. Figure 2B and Figure 2C present the VLE diagrams for the binaries AA/2-EH and AA/2EHA, respectively. D


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Figure 2. Phase equilibrium diagrams (mass based) for selected ternary and binary systems. Ternary liquid−liquid (A) at 30 °C and 1.013 bar for water/AA/2-EHA, water/AA/2-EH, and water/2-EH/2-EHA; minimum binary heterogeneous azeotropes: (◇) water/2-EHA (85.5/14.5% mass, Tb = 99.5 °C), (○) water/2-EH (80.9/19.1% mass, Tb = 99.1 °C). Binary vapor−liquid diagrams at 0.2 bar for AA/2-EHA (B) and 2-EH/2-EHA (C).

conversion interval (say, 20−75%). However, it is evident that a conversion around 70% ensures a small recycle and a reasonable amount of catalyst. Above this value, the amount of catalyst significantly increases, while the decrease in recycle flow rate is relatively low. The second analysis investigates the influence of recycle composition on reactor design, namely on the amount of catalyst necessary to achieve a certain conversion. Figure 4 (right) presents the mass of catalyst (mcat) required to achieve 71% AA conversion (XAA) as a function of the separation factor

Figure 3. Reaction−separation−recycle structure.

but these costs are difficult to estimate at this preliminary stage of design. The optimum value may very well lie within a large

Figure 4. Left: Amount of catalyst and recycle flow rate versus conversion. Right: Influence of recycle composition on the amount of catalyst needed to achieve 71% AA conversion. E


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Figure 5. Left: plantwide control structure; FAA,0 and F1/FAA,0 ratio fixed; F2‑EH,0 on level control. Right: XAA versus FAA,0/mcat ratio bifurcation diagram, at different values of the F1/FAA,0 set-point.

(βwater or β2‑EHA) for the case of complete reactants recycle (β2‑EH = 0, βAA = 0). As βWater or β2‑EHA decreases (βi = 1 means complete removal), the required amount of catalyst increases. In the interval 0.8−1, both water and 2-EHA recycle have a similar and a relatively small effect. Therefore, the recycle stream does not need to have a tight specification on water and 2-EHA concentrations. 3.2.3. Throughput. To analyze the throughput influence on process operation, represented here by XAA, choices are made in order to fix some key process variables. As in the previous case, the reactor operates adiabatically with an inlet temperature of 120 °C. Since in most of the cases the reactants have the largest share on the overall economics, it is fair to assume that they are completely recovered and recycled to the reaction section (β2‑EH = 0, βAA = 0); in addition, the products must be obtained within specifications, usually at high purity. The analysis on recycle composition showed that the recycle stream does not need to have tight specification (some contamination with products is allowed); thus, the separation factors for water and 2-EHA are selected to be βWater = β2‑EHA = 0.9. The last specification is in relation to the alcohol-to-acid ratio at the reactor inlet. To directly fix this ratio to a specific value, concentration measurements are required in practice. However, one indirect way to maintain the alcohol-to-acid ratio around a nominal value is by fixing the ratio F1 to FAA,0, as shown by the plantwide control structure in Figure 5 (left). This structure fixes FAA,0 and the F1/FAA,0 ratio at the reactor inlet. Since FAA,0 is added on flow control, Fn‑BuOH,0 is fed on level control to maintain the inventory. The dependence of the state variable XAA versus the model parameter FAA,0/mcat, for different values of the ratio set-point F1/FAA,0, is shown in Figure 5 (right). The process presents two steady states at low FAA,0 values, or no steady state beyond a critical FAA,0 value. For mcat, F1/FAA,0, βi fixed, and operation on the high-conversion branch, it is observed that increasing FAA,0 leads to a decrease in XAA. To explain the observed behavior, we note that the amount of catalyst sets a constraint on the rate at which reactants are transformed into products. Thus, if the flow rate of fresh reactant (FAA,0) exceeds the reactor capacity, reactants accumulation occurs, and no steady state can be reached. This particular control structure can be applied to systems involving two reactants that are recycled together. Its development and feasibility were discussed by Altimari and

Bildea11 (and extended in the chapter dedicated to plantwide control in the recent book of Dimian et al.12), and was recently applied to design and control of a similar acrylate process by Moraru and Bildea.13 Other plantwide control structures are possible, which can rely on concentration measurements. These usually introduce large dead-times (while flow rates are immediately measured). The proposed plantwide control structure does not require composition control for maintaining the component inventory. 3.3. Preliminary Operating Point. The sensitivity analyses made so far reveal a large process operating window, with a broad range of values for the key process parameters. At this stage of design, a rough selection of these variables is made so that a preliminary mass balance is obtained. This mass balance is necessary in order to start the synthesis of the separation system. Once designed, a detailed analysis of the RSR can be made, and a decision on the nominal operating point can be taken. This point should be far enough from the large sensitivity region (turning point of the Figure 5 diagram) to provide flexibility in throughput, to show robustness in face of perturbations, and to ensure safe operation. The selected preliminary operating point is marked by the empty-white dot presented in Figure 5 (right). The ratio setpoint (F1/FAA,0 = 8.8, mass-based) ensures an alcohol-to-acid molar ratio at the reactor inlet of 3. The AA conversion (XAA = 67.6%) and separation factors (β2‑EH = βAA = 0, βWater = β2‑EHA = 0.9) are such selected to keep a balance between reactor size (amount of catalyst) and size of the recycle (unconverted reactants). The fresh AA feed (FAA,0 = 13.6 kmol/h) is selected to meet the required capacity. Therefore, given the nominal flow rate, conversion, and ratio set-points, the amount of catalyst (mcat = 1000 kg) is determined. This point is fairly far away from the infeasibility area and has enough flexibility at changes in fresh AA feeds. The mass balance of this preliminary operating point is given in Table 5.

4. PROCESS DESIGN The process synthesis of the reaction−separation structure is based on a rigorous analysis of physical properties and key phase equilibrium diagrams. Once the overall structure is developed, the selection of the nominal operating point of the plant is based on a sensitivity analysis that aims to decrease the heating requirement (i.e., steam consumption) and to increase the wastewater purity (i.e., utilization of raw materials). F


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purity. The heating requirement is a direct measure of steam consumption, while the water purity reveals the degree of reactants utilization. In both cases, the impact on economics is large as both steam and reactants are expensive.14 It is worth mentioning that a complete process optimization is not attempted in this study. On the basis of this sensitivity analysis, the nominal operating point of the plant is selected. This point provides the basis for equipment sizing and economic calculations, as well as for determining the holdup of the main equipment required in dynamic simulations to assess the robustness of the control structure. 4.2. RSR-A Process. The complete process flow diagram is presented in Figure 6. The reactor (PFR) is a multitubular equipment with the catalyst inside the tubes. The separation section removes the water and acrylate from the process, and recovers the reactants in a single recycle stream. The recycle is mixed with the fresh alcohol and fed to the reactor. The fresh acid is fixed on flow control. The separation section is designed based on the mass balance of the initial operating point. The composition of the reactor outlet (see stream 2 in Table 5) shows that it contains about 2.8% water, 4.9% AA, 63.3% 2-EH, and 29.0% 2-EHA (all by mass). The thermodynamic analysis suggests that water can be fairly easily separated: the low-boiling heterogeneous azeotropes containing a large amount of water can be obtained as vapor distillates, condensed, and liquid−liquid separated; then, the water phase with a fairly low content of organics (see ternary diagrams in Figure 2A) is removed, while the organic phase is returned to the process. Therefore, the first separation step uses a reboiled stripping column (C-1) which obtains the water-containing heterogeneous azeotropes as top vapors. After condensation (COND-1) and liquid−liquid splitting (V-2), water is removed from the system at a relatively high purity. To increase the water purity, part of the water stream is mixed with the organic stream from this decanter (V-2) and sent as reflux

Table 5. Mass Balance of the Preliminary Operating Point (Figure 5, Right, White Open Dot) stream

0a

0b

1

2

3

4a

4b

mole flow/ [kmol/h] mole fraction water AA 2-EH 2-EHA mass flow/ [kg/h] mass fraction water AA 2-EH 2-EHA

13.6

13.6

83.3

83.3

56.1

13.6

13.6

0.181 0.078 0.560 0.181 9588

0.027 0.116 0.831 0.027 6843

1

1

0.018 0.241 0.723 0.018 9588

0.003 0.151 0.817 0.029

0.028 0.049 0.633 0.290

0.004 0.068 0.887 0.041

1

1

978

1767

1 1

244

1 2500

1

4.1. Design Procedure. The design procedure is valid for each of these processes. The detailed process structure of the plant is developed considering the reaction and separation sections coupled by recycle (i.e., the RSR system). In addition, the one-recycle structure is maintained, the reactants being recycled together. Thus, the characteristics of the RSR system previously described are preserved. While this design approach can easily fix the amount of catalyst and establish the size of the fixed-bed reactor for a given conversion and recycle composition, the design of the separation section requires a feed composition (reactor outlet) based on which the separation sequence can be developed and later on sized. A good starting point providing an initial mass balance is given by the preliminary operating point selected in section 3.3. Once the detailed process structure is known, a sensitivity analysis is made with respect to its key design variables which have an influence on the heating requirement and wastewater

Figure 6. Process flow diagram and plantwide control of the RSR-A process. G


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Figure 7. Sensitivity analysis and operating point of the RSR-A process; independent variables (x-axis) from top-left to bottom-right: NT,C‑1, NF,C-1, NT,C-2, NF,C‑2, mcat, fsplit. Dependent variables: sum of reboilers duty (primary y-axis) and wastewater purity (secondary y-axis).

NF,C‑2), the mass of catalyst in the reactor (mcat), and the split fraction ( fsplit) of the aqueous stream sent as reflux to the first column. The stripper’s (only) degree of freedom, namely the reboiler duty, was adjusted such that the mass fraction of water in the bottom stream 4 has a required value, set at 500 ppm mass. Meanwhile, fsplit is used to increase the water purity. Regarding the product column C-2, one DOF (bottoms to feed ratio) was used to hold the product purity at 0.995. The second DOF (reflux ratio) was used to achieve the desired product recovery of 0.9, namely the 2-EHA in the product stream/2EHA in the feed stream. Starting with some initial values of the independent variables, these values are varied one at a time (while the other ones remain fixed), and the values of the dependent variables (overall duty and the water purity) are recorded. When varying the number of stages (NT,C‑1 and NT,C‑2), the feed tray remains at the same relative location between the top and bottom of the column (i.e., NF/NT = constant). In the second step, new values of the independent variables are selected such that the

to the column. Since the ternary system AA/2-EH/2-EHA forms no azeotropes and the x−y diagrams of AA/2-EHA and 2-EH/2-EHA (Figure 2B,C) show an easy separation, the product 2-EHA can be obtained by distillation. However, some product is expected to be recycled together with the two reactants, since many separation stages are suggested (Figure 2C) for obtaining a 2-EH stream free of 2-EHA. Therefore, the bottom of the reboiled stripper (C-1) is fed to a distillation column (C-2) where 2-EHA is obtained as bottoms product (99.5%mass), while a mixture containing the two reactants and some 2-EHA is recycled to the reaction section. The sensitivity analysis aims to find a robust operating point, given the nominal production capacity of 20,000 t of acrylate per year. The operation at this point should have a relatively low steam consumption, namely the sum of columns reboilers (C-1, C-2) and preheater (HX-1) duties, and achieve a relatively high purity of the wastewater. The key design variables (i.e., independent variables) are the number of trays and the feed tray of each of the columns (NT,C‑1, NF,C‑1, NT,C‑2, H


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Industrial & Engineering Chemistry Research Table 6. Mass Balance of the RSR-A Process for the Nominal Operating Point stream name:

0a

0b

temp/[°C] pressure/[bar] mole flow/[kmol/h] mass flow/[kg/h] mass fraction water AA 2-EH 2-EHA mole fraction water AA 2-EH 2-EHA

25 1.5 13.6 978

25 1.5 13.7 1779

1 1

1 1

2

3

4

5

120 1.5 76.2 8892

1

113.9 1.5 76.2 8892

108.3 0.1 48.9 6135

120.5 0.145 62.6 8647

30 0.1 0.3 40

30 0.1 9.0 163

108.3 0.1 82.7 10364

30 0.1 13.6 245

30 0.22 13.7 2512

486 ppm 0.151 0.817 0.031

0.028 0.041 0.619 0.312

0.001 0.059 0.895 0.045

500 ppm 0.042 0.636 0.321

0.019 0.002 0.561 0.418

0.998 0.001 441 ppm 42 ppm

0.001 0.059 0.895 0.045

0.998 0.001 441 ppm 42 ppm

trace 0.5 ppm 0.005 0.995

0.003 0.244 0.733 0.020

0.181 0.066 0.555 0.198

0.005 0.103 0.861 0.031

0.004 0.080 0.675 0.241

0.141 0.004 0.561 0.295

1 257 ppm 61 ppm 4 ppm

0.005 0.103 0.861 0.031

1 257 ppm 61 ppm 4 ppm

trace 12 ppm 0.007 0.993

Table 7. Selected Design Variables for the RSR-A Process value

NT,C‑1 NF,C‑1 NT,C‑2 NF,C‑2 mcat/kg fsplit

10 7 25 17 1600 0.4

7

8

9

bottoms liquid, while the reactants are obtained as distillate stream which is recycled to the reaction section. The sensitivity analysis has the same objective. The dependent variables remain the same, namely the steam requirements (i.e., in this case, the sum of reboiler duty of the only column C-2 and preheater HX-1 duty) and the wastewater purity. The independent variables are the number of trays and the feed tray of the column C-2 (NT, NF), the mass of catalyst in the reactor (mcat), and the pressure (pflash) in the flash vessel V-4. One remark here is that the wastewater purity could not be set at a required value since no degrees of freedom exist: the purity of the water product is given by the vapor− liquid equilibrium in the flash, and by liquid−liquid equilibrium in the decanter. The procedure remains the same as well, varying the values of the independent variables one at a time while the others remain fixed. The overall duty and wastewater purity are recorded, and the new values for the variables are selected in the following step such that the duty reduces and the wastewater purity increases. In one particular case however (i.e., when varying the pflash), this is not possible since both dependent variables move in the same direction. In this situation, contrary to the previous case, the independent variable is selected such that a low duty is achieved. This is because the duty significantly increases at higher flash pressures, and the increase in water purity is not significant, not driving the process to achieve high purity wastewater (i.e., no significant amount of reactants recovery). This is a trade-off that can be investigated by solving an optimization problem. The procedure is repeated until no improvement in the dependent variables is observed. Figure 9 shows the results of the last step. The values of the independent variables marked by the full-black dots were selected in the previous step. A change (increase/decrease) in the variables does not lead to a significant improvement in duty and wastewater purity. Thus, this is selected as the nominal operating point. A detailed mass balance of this point is presented in Table 8. The selected independent variables are given in Table 9; they correspond to an overall duty of 2972 kW and 96.1% mass water purity. 4.4. RSR-C Process. Figure 10 shows the process flow diagram. As in the previous two cases, the reactor and the recycle structure remain unchanged. Since AA is in small concentration (see Table 5), the composition at the reactor outlet falls into a heterogeneous

duty reduces and the wastewater purity increases. In one particular case (i.e., when varying the fsplit), this is not possible since both dependent variables move in the same direction (i.e., both, duty and water purity increase, or decrease). In this situation, the independent variable is selected such that highpurity wastewater is achieved, since the increase in the overall duty is relatively small. This two-step procedure is repeated until no improvement in the dependent variables is observed. Figure 7 shows the results. The values of the independent variables marked by the full-black dots were selected in a previous sensitivity step. When these are again varied in a new step, no significant improvement in duty and wastewater purity can be observed. Therefore, this is selected as the nominal operating point of the process (note that in the multidimensional space of the independent variables, there is actually one single point). A detailed mass balance of this point is presented in Table 6. The selected values of the independent variables are given in Table 7, and correspond to an overall duty of 2645 kW and 99.9% mass water purity.

variable

6

4.3. RSR-B Process. The process flow diagram of this process alternative is presented in Figure 8. The reactor type remains the same as for the previous process, as well as the recycle structure which has the reactants recycled in a single stream. Again, the separation system is developed starting with the initial mass balance and making use of thermodynamic considerations. The water-containing heterogeneous azeotropes can also be obtained as vapors by flashing. In this way, the water is removed in a flash-decanter system. So, the reactor outlet stream is fed first to a flash vessel (V-4), condensed (COND1), and liquid−liquid phase separated (V-2). The water is eliminated, while the organic phase is mixed with the liquid stream of the flash vessel and undergoes the same separation in a distillation column (C-2). The acrylate product is obtained as I


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Figure 8. Process flow diagram and plantwide control of the RSR-B process.

Figure 9. Sensitivity analysis and operating point of the RSR-B process; independent variables (x-axis) from top-left to bottom-right: NT, NF, mcat, pflash. Dependent variables: reboiler duty (primary y-axis) and wastewater purity (secondary y-axis).

The dependent variables for the sensitivity analysis remain the same, while the independent variables are the number of trays and the feed tray of the column C-2 (NT, NF) and the mass of catalyst in the reactor (mcat). In this process, the water purity is given by the liquid−liquid equilibrium in the decanter, no variable being available to be manipulated.

region (Figure 2A). Therefore, one can directly exploit the heterogeneous area, since large immiscibility exists between water/2-EH and water/2-EHA. This suggests that water can be separated immediately after the reaction, in a liquid−liquid decanter (V-2). Again, 2-EHA is obtained as a bottom stream in a distillation column (C-2), while the reactants are recovered as distillate and recycled to the reaction section. J


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Industrial & Engineering Chemistry Research Table 8. Mass Balance of the RSR-B Process for the Nominal Operating Point stream name:

0a

0b

1

2

3

4

6

7


25 1.5 13.6 978

25 1.5 13.5 1762

120 1.013 81.8 9241

114.1 1.013 81.8 9241

70.9 0.1 54.7 6501

88.1 0.15 65.1 8654

30 1.5 3.1 335

85.9 0.18 68.2 8989

70.9 0.1 74.1 8809

30 1.5 13.6 252

30 0.22 13.5 2488

0.006 0.150 0.814 0.030

0.033 0.045 0.624 0.298

0.009 0.063 0.886 0.042

0.006 0.045 0.641 0.309

0.034 0.070 0.679 0.217

0.007 0.046 0.642 0.306

0.009 0.063 0.886 0.042

0.961 0.038 961 ppm 37 ppm

trace 5 ppm 0.005 0.995

0.040 0.235 0.733 0.020

0.204 0.071 0.555 0.198

0.060 0.104 0.861 0.031

0.041 0.082 0.675 0.241

0.202 0.105 0.561 0.295

0.048 0.083 0.650 0.219

0.060 0.104 0.861 0.031

0.990 0.010 61 ppm 4 ppm

trace 12 ppm 0.007 0.993

1 1

1 1

value

NT NF mcat/kg pflash/bar

25 17 1600 0.15

8

9

5. PROCESS CONTROL The control structure of each process preserves the plantwide control structure initially proposed (see Figure 5left). One can define the plantwide control as the control philosophy of the overall plant with emphasis on structural decisions.15 The primary objective is to achieve stable operation by keeping the mass inventory, and to maintain the product and wastewater purities at their set-point, despite any disturbances. Three types of disturbances (process changes) are studied to show the robustness of the overall process control scheme: plant throughput, contamination of fresh reactants, and temperature at the reactor inlet. The dynamic simulations require that the holdup of the major equipment (sumps, mixers, flash, reflux, and decanter vessels) is known. The sizing is described in section 6.1. Once these dimensions are introduced in Aspen Plus and the simulation converged, the flowsheet is exported in Aspen Plus Dynamics as a flow-driven simulation.

Table 9. Selected Design Variables for the RSR-B Process variable

5

Figure 11 shows the results of the last step of the sensitivity analysis. As in the previous cases, increasing or decreasing the values of the independent variables does not improve significantly the overall duty or the water purity. The mass balance of the nominal operating point is showed in Table 10. The values of the independent variables are given in Table 11, and correspond to an overall duty of 4151 kW and 96.8% mass water purity.

Figure 10. Process flow diagram and plantwide control of the RSR-C process. K


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manipulating the reboiler duty, sump level is controlled by the bottoms flow rate, and top pressure is controlled by the vapor flow rate. The organic level in decanter V-2 is controlled by the organic flow rate, while the aqueous level is controlled by the aqueous flow rate. The ratio between the two aqueous splits (i.e., the stream sent as reflux and the wastewater) is maintained constant. The feed temperature to V-2 (C-1 condensed vapor) is controlled by the duty of COND-1. In column C-2, the temperature on stage-22 is controlled by manipulating the reboiler duty, the level in the sump is controlled by the bottoms flow rate. At the top, the pressure is kept constant by the condenser COND-2 duty, and the level in the reflux vessel is controlled by the distillate flow rate. In the recycle vessel V-1 the level is controlled by the fresh alcohol. Most of the temperature, pressure, and level controllers are PI-type, tuned by specifying appropriate ranges of the controlled (PV) and manipulated variables (OP), and setting the gain and the integral time to 1% OP range/% PV range, and to estimated values of the process time constant, respectively. Key temperature controllers (bottoms C-1 and bottoms C-2) however are tuned using the closed loop Auto Tune Variation (ATV) method implemented in Aspen Plus Dynamics. With all the other control loops closed, the ATV method finds the stability limit using a relay of amplitude 5% of the controller output range (temperature). For each temperature control loop, a dead time of 1 min is considered. After finding the ultimate gain (Ku) and ultimate period (Tu), the Kc and Ti are calculated using the Tyreus−Luyben tuning rules. Detailed controller settings are shown in the Supporting Information. The simulations presented in section 5.1.2 show that in two specific situations, for +25% FAA,0, and +5 °C at the reactor inlet, the product purity drops from 99.5 to 99.1, and 99.2% mass, respectively. To maintain the purity, a concentration controller (QC) in cascade with the temperature controller of C-2 is implemented. A 30 min sampling time and a 30 min deadtime arising due to the concentration measurement are taken into account. The concentration controller has a set-point of 0.5% mass 2-EH + AA so that the acrylate product specification can be reached. The controller output is the setpoint to the temperature controller. The action of the concentration controller is reverse: when the sum 2-EH + AA concentration increases, the temperature set-point is increased forcing the reboiler duty to increase. Thus, as the bottoms temperature increases, more 2-EH and AA is vaporized, and a purer product is obtained. When used for tuning the same ATV method, a very large reset time is calculated (248 min) due to the large delay of the concentration measurements. However, when a faster analyzer is assumed in the control loop (for the purpose of tuning the controller), say 5 min sampling + 5 min deadtime, the ATV method with the Tyreus−Luyben rules find a more aggressive controller. Despite that these parameters are determined considering a relatively fast analyzer, the simulations results show that these parameters are conservative. The control system presents a good performance even for large measurement delays of 30 min. This relatively large time span expands the possible alternatives for concentration measurements that could be used in this control loop. One suggestion is the use of on-line measurement techniques, obtaining directly or indirectly the concentration of impurities (i.e., AA and 2EH). 5.1.2. Control Structure Testing. The performance of the control structure is tested by changing key operating parameters from their nominal values:

Figure 11. Sensitivity analysis and operating point of the RSR-C process. Independent variables (x-axis) from top to bottom: NT, NF, mcat. Dependent variables: reboiler duty (primary y-axis) and wastewater purity (secondary y-axis).

5.1. RSR-A Process. 5.1.1. Control Structure and Controllers Settings. The control structure is shown on the process flow diagram in Figure 6. At the plant level the fresh AA (FAA,0) is set on flow control (FC), setting in this way the plant capacity. The fresh 2-EH (F2‑EH,0) is added on level control (LC) to maintain its inventory, thus meeting the stoichiometry requirement of the reaction. The reactor is set to operate at a fixed F1/FAA,0 mass ratio of 8.09, ensuring a molar ratio 2-EH/ AA of about 3 at nominal operation. Therefore, the plantwide control strategy proposed at the previous design level (i.e., kinetic-based reactor/black box separation/recycle) is preserved. At the unit level, standard control applies. The temperature of the reactor inlet is controlled at 120 °C by manipulating the duty of the heat exchanger HX-1. A temperature in the bottom section of column C-1, on stage-8, is controlled by L


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Industrial & Engineering Chemistry Research Table 10. Mass Balance of the RSR-C Process for the Nominal Operating Point stream name:

0a

0b

1

2

3

4


25 1.5 13.6 978

25 1.5 13.6 1765

120 1.013 109.3 10954

115 1.013 109.3 10954

46.1 0.1 82.2 8211

30 1.013 95.7 10703

30 1.013 13.6 251

30 0.28 13.6 2492

0.03 0.147 0.798 0.025

0.052 0.059 0.638 0.252

0.04 0.077 0.849 0.034

0.031 0.059 0.652 0.257

0.967 0.032 833 ppm 40 ppm

trace 3 ppm 0.005 0.995

0.168 0.205 0.614 0.014

0.291 0.081 0.491 0.137

0.223 0.107 0.651 0.018

0.192 0.092 0.560 0.156

0.992 0.008 118 ppm 4 ppm

trace 9 ppm 0.007 0.993

1 1

1 1

value

NT NF mcat/kg

37 28 2200

9

The impurification with 5% mass water of the fresh acid (middle-left diagram) decreases the capacity with about 5%. The transition period is small, basically the same as the time period of ramping. The fresh alcohol decreases also by 5%, while the fresh acid and wastewater increases by about 15%. No practical change is observed in the acrylate and wastewater purities. For the impurification of the fresh alcohol (middleright diagram), the production and the fresh acid flow remains the same, while the fresh alcohol and wastewater streams increase by about 37%. There is no impact on the acrylate and wastewater purities. When changing the reactor inlet temperature with +5 °C (bottom-left diagram), the dynamic behavior is even less impacted; the transition period is very short while the flow rates, as expected, do not change. However, while the water purity remains unchanged the purity of the product decreases from 99.5 to 99.0% mass. This offset is eliminated when the concentration controller is used (dashed line in the bottom-left diagram). The −5 °C change of the reactor inlet brings no changes in any of the variables of interest (bottom-right diagram). 5.2. RSR-B process. 5.2.1. Control Structure and Controllers Settings. Figure 8 shows the plantwide control structure for the RSR-B process. This control structure is very similar to that of the RSR-A process. At the plant level, the fresh AA (FAA,0) sets the plant capacity, while the fresh 2-EH (F2‑EH,0) is added on level control to maintain its inventory. The reactor operates at a fixed F1/FAA,0 mass ratio of 8.45, ensuring a molar ratio 2-EH/AA of 3 at nominal operation. The control at the unit level is made in the same manner. Although the first column in the RSR-A process is replaced by a flash vessel (V-4) for water separation from the reactor outlet stream, it does not present a different control configuration: the pressure is controlled by the vapor flow rate, while the liquid level is controlled by manipulating the liquid flow rate. Again, all the control loops use PI controllers. The parameters of all controllers are given in the Supporting Information. The simulations presented in the following section 5.2.2) show that in one specific situation, for +25% FAA,0, the product purity drops from 99.5 to 99.2% mass. Therefore, a concentration controller (QC) in cascade with the temperature controller of C-2 is implemented. The tuning of this controller and the final selection of the parameters are made in the same

Table 11. Selected Design Variables for the RSR-C Process variable

8

• ±25% for FAA,0 (change in throughput) • 5% mass water in FAA,0, and F2‑EH,0 (contamination of fresh reactants) • ±5 °C for temperature at the reactor inlet (change in reactor operating conditions) Each change is introduced one at a time. After 2 h of steady state operation, the changes in flow rate and contamination of fresh reactants are introduced as a ramp, time for 1 h, while the temperature is modified in a step change, at once. Figure 12 presents the simulation results. All diagrams show selected plant flow rates on the primary y-axis, and products purity on the secondary y-axis. Each diagram corresponds to one change of only one operating parameter: top left/right diagrams correspond to +/−25% FAA,0; middle left/right diagrams to 5% mass water in FAA,0/F2‑EH,0; and bottom left/ right diagrams to +/−5 °C temperature at the reactor inlet. For the increase with +25% of FAA,0 (top-left diagram), from 13.6 to 17.0 kmol/h, the plant capacity (i.e., the acrylate flow rate) also increases by 25%. In about 5 h from the start of ramping the flow rate, the process reaches the new steady state. The increase of FAA,0 with +25% leads to an increase in the fresh alcohol and wastewater flow rates with the same percentage. The wastewater purity remains practically constant (99.8% mass). However, the acrylate purity decreases from 99.5 to 99.1% mass, being contaminated with alcohol. During the 5 h transition period from one operating point to another, the contamination is rather large, the acrylate purity reaching a minimum of even 88% mass. The offset of 0.4% mass is eliminated when the concentration controller is used, bringing the purity back to the desired value (dashed line in the top-left diagram). As expected, the decrease with 25% of FAA,0 (topright diagram) leads to a decrease in capacity with the same percentage, as well as a decrease, with the same number, of the fresh alcohol and wastewater streams. Regarding the acrylate purity, it basically suffers a very small increase (from 99.5 to 99.7% mass), while the wastewater purity remains at the same value (99.8% mass). M


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Figure 12. Dynamic simulation results for the RSR-A process: top left/right diagrams, ±25% FAA,0; middle left/right diagrams, 5% mass water in FAA,0/F2‑EH,0; bottom left/right diagrams, ±5 °C reactor inlet.

concentration controller is still required when capacity is increased by 25% (i.e., increase of FAA,0). 5.3. RSR-C Process. 5.3.1. Control Structure and Controllers Settings. Figure 10 shows the plantwide control structure for the RSR-C process. This control structure is very similar to those of the other two processes, RSR-A and RSR-B, and its description follows closely the other two previously presented. The plant capacity is set by the fresh AA flow rate, while the fresh 2-EH maintains the inventory. The reactor operates at a fixed F1/FAA,0 mass ratio of 10.2, ensuring a molar ratio 2-EH/AA of 3 at nominal operation. The column and the flash vessel are replaced by a decanter (V-2) for water removal, which has the two liquid levels controlled by manipulating the two outlet liquid flows. Since the other equipment remains the same, the control is made in the same manner. The parameters of all controllers are given in the Supporting Information. A concentration controller in cascade with the temperature controller of C-2 is needed when the fresh acid is ramped with +25% (see section 5.3.2). The tuning and the final selection of

manner as in the RSR-A process. The controller parameters (Kc and Ti) are given in the Supporting Information. 5.2.2. Control Structure Testing. For testing the control structure, the same operating parameters (throughput, contamination of fresh reactants, inlet reactor temperature) are changed. The same holds for the manner of introducing the changes. Figure 13 presents the simulation results. The arrangement of each diagram follows the same description as in the RSR-A process. As expected, the dynamic behavior is very similar to that of the RSR-A process, and one can closely follow the description previously presented and relate to the actual diagrams. Only a few differences between the two processes can be observed. One of them is that the wastewater shows larger deviations from the nominal operating point; however the starting point is completely different, since this process achieves low purity wastewater. Another difference is that increasing the reactor inlet temperature has no impact on the process dynamics, and concentration measurements are not required. However, a N


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Figure 13. Dynamic simulation results for the RSR-B process: top left/right diagrams, ±25% FAA,0; middle left/right diagrams, 5%mass water in FAA,0/F2‑EH,0; bottom left/right diagrams, ±5 °C reactor inlet.

6.1. Equipment Sizing. The sizing is based on the material and heat balance of the steady state simulations. The columns diameters are sized using the packing sizing tool of Aspen Plus. The sumps, mixers, and reflux vessels are sized based on a residence time of 10 min. The flash vessels and decanters are sized based on 10 min residence time and the half-full criteria (i.e., 20 min residence time). All heat exchangers are designed based on a heat transfer coefficient of 815 W/(m2·K). The reactor is a multitubular equipment. The catalyst is of spherical shape particles of 0.5 mm in diameter with a density of 1540 kg/m3 (770 kg/m3 bulk density, 0.5 bed voidage) accommodated inside the tubes. The main dimensions of all process equipment are shown on each process flow diagram: Figure 6, Figure 8, and Figure 10 for the RSR-A, -B, and -C, respectively. 6.2. Economics Basis. Equations 7−13 calculate the annual economic potential (EP) and other key economic indicators. The EP is the difference between the sales revenue (Cproduct) and the sum of the expenses of raw materials (Creactants), utilities (Cutilities), and three years depreciation (Cdepreciation) of the main fixed assets. The specific annual economic potential (EPspec) is

the controllers parameters are made in the same manner. The controller parameters Kc and Ti are given in Supporting Information. 5.3.2. Control Structure Testing. Figure 14 presents the simulations results for the same process changes. The arrangement of each diagram follows the previous descriptions. The dynamic behavior show that it takes more time for the flow rates to settle after a disturbance is introduced. All process changes do not bring unexpected behavior on product purities. A concentration controller is still required when the capacity is increased by 25%.

6. ECONOMIC EVALUATION AND PROCESS COMPARISON A rough economic evaluation and comparison between the three process alternatives is made. The evaluation uses generally accepted methods for capital and utility costs calculations. The process comparison is based on two major criteria: key economic indicators and process parameters. O


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Figure 14. Dynamic simulation results for the RSR-C process: top left/right diagrams, ±25% FAA,0; middle left/right diagrams, 5% mass water in FAA,0/F2‑EH,0; bottom left/right diagrams, ±5 °C reactor inlet.

Cspecproduction = Cspecreactants + Cspecutilities + Cspecdepreciation

obtained by dividing the EP to the annual capacity (P2‑EHA). Similarly, the other specific costs (i.e., all per tonne of product) of reactants (Cspec reactants), utilities (Cspec utilities), and depreciation (Cspec depreciation) are obtained by dividing each cost to P2‑EHA. The specific cost of production (Cspec production) is the sum of all specific costs. They are all expressed in USA dollar ($). EP = Cproduct − (Creactants + Cutilities + Cdepreciation)

(7)

EPspec = EP/P2 − EHA

(8)

Cdepreciation = C TCI/payback period

(9)

Cspecreactants = Creactants/P2 − EHA

(10)

Cspecutilities = Cutilities/P2 − EHA

(11)

Cspecdepreciation = Cdepreciation /P2 − EHA

(12)

(13)

The prices for product and reactants are taken from the MOLBASE, one of the largest platforms for chemicals buying and selling on line (http://www.molbase.com): 2258 $/t 2EHA, 1381 $/t AA, and 1738 $/t 2-EH; these are reference prices in July, 2017. Although not used in any calculations, it is worth mentioning that ICIS, another large source that shares chemical prices, gives similar values (https://www.icis.com): 2116 $/t 2-EHA and 1367 $/t 2-EH; these are indicative prices in July, 2017 (thus, the economic potential is for half 2017, since the costs of raw materials and product are the main economic drivers for the revenue calculation). No price for technical grade AA was found on this site. The cost of low-pressure and medium-pressure steam is 7.78 and 8.22 $/GJ, respectively.14 The cost of cooling water is 0.354 $/GJ.16 The wastewater cost is calculated based on concentration of organics and total flow rate, namely 500 $/t of organics17 and 27 $/t of wastewater,18 respectively. P


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Industrial & Engineering Chemistry Research The total capital investment (CTCI) used in calculating the depreciation cost is calculated using eq 14. This is the individual factored method of Guthrie expressed by eq 15, described in detail by Seider et al.,19 to which the costs for royalties and start-up are added (suggested by the same reference). C TCI = C TCI,Guthrie + Croyal + Cstartup

Table 12. Key Economic Indicators and Process Parameters of All Three Process Alternatives process capacity EP EPspec Cspec reactants Cspec utilities Cspec depreciation Cspec production

(14)

C TCI,Guthrie = C TPI + C WC = 1.18(C TBM + Csite + C buildings + Coffsitefacilities) + C WC C TCI = 1.6562 × C TBM

(15)

AA conversion 2-EH conversion water purity

(16)

One remark here is that the economic calculations take into account only the major equipment. In addition, the economic burden of vacuum operation is not considered; the assumption is that this plant is operated on a site that has available all utility systems, including vacuum capabilities. In addition, the vacuum level required in these processes (i.e., 0.1 bar) can be easily achieved using one-stage liquid-ring pump. The basic equipment dimensions, material of construction, and operating pressure provide the input for calculating the cost of total bare-module (CTBM), also known as the total installed equipment cost. For consistency reasons, the cost equations are taken from a single reference.18 The installed cost of vessels and columns (shells) is calculated with eq 17. M&S is the Marshall and Swift equipment cost index (1536.5 for 2011), D and H are diameter and height in meters, Fm is the material factor (Fm = 1 for CS, 2.25 for SS) and Fp is the pressure factor (Fp = 1 + 0.0074(P − 3.84) + 0.00023(P − 3.48)2). For the cost of shell and tubes heat exchangers, eq 18 is used. Since the tubular reactor is a shell and tube-like equipment, its installed cost is calculated using the same equation. Fm is the correction factor for material (Fm = 1 for CS/CS, 2.81 for CS/SS, and 3.75 for SS/SS), Fd is the correction for design type (Fd = 0.8 for fixed-tube sheet, 1.35 kettle reboiler) and Fp is the correction for the design pressure (Fp = 0 for design pressures less than 10 bar). The catalyst costs 50 $/kg.14 The cost of the MellapakPlus structured packing, 10000 $/m3, is taken as for Mellapak reported by Bildea et al.20

RSR-A Economic Indicators [t/a] 20,095 [$/a] 8,124,761 [$/t] 404 [$/t] 1,768 [$/t] 35 [$/t] 50 [$/t] 1,854 Process Parameters [%] 99.97 [%] 99.29 [%] 99.8

RSR-B

RSR-C

19,902 7,584,084 381 1,786 39 52 1,877

19,937 7,048,299 354 1,773 53 78 1,904

99.02 99.29 96.1

99.18 99.29 96.7

• Generally, the key economic and process indicators of the RSR-A process are slightly better than those of the other two alternatives • The EP (annual earnings) of the RSR-B and RSR-C processes are about 6.7% and 13.2% lower, respectively; these, are closely followed by the EPspec (earnings per ton of product) which shows values of 5.7% and 12.6% lower • The Cspec utilities is about 11% and 34.6% higher for the RSR-B and RSR-C processes; the Cspec depreciation is also larger, with about 3.3% and 35.7% • Overall, the sum of the specific cost of reactants, utilities, and depreciation, results in the Cspec production being also higher; this cost is 1.2% and 2.7% larger for the RSR-B and RSR-C process, respectively; note that the Cspec production is driven by the specific cost of reactants (Cspec reactants), which has by far the largest share • The RSR-A process achieves higher purity wastewater, namely 99.8% mass compared to 96.1 and 96.7% mass purities of the RSR-B and RSR-C processes; this is not surprising since, given the same feed, the RSR-A uses a distillation-decanter system to remove the water, while the RSR-B and RSR-C use a flash-decanter and a decanter-only system, respectively • The acid conversion (over the complete process) is lower with about 1% and 0.8% for the RSR-B and RSR-C processes, while the alcohol conversion is the same for all three processes; it is worth mentioning that for these two processes, the acid lost in the wastewater brings a penalty on the annual earnings of about 89,000 $/a and 106,000 $/a, respectively It should be remarked that reactive distillation with decanter separator for reactant recovery is an economically attractive alternative for esterification processes in general,21,22 for higheracrylates production in particular,6,23,24 leading to high process performance in terms of high acid conversion, high product purity, and low heat duty.

Cvessels/columns = (M&S/280)(957.9D1.066H 0.82)(2.18 + FmFp) (17)

C HX = (M&S/280)(474.7A0.65)(2.29 + Fm(Fd + Fp)) (18)

All the other costs in eqs 14 and 15 are factors of CTBM. The value of each factor is selected based on the description provided by Seider et al.19 The selection of these factors and the calculations made for the RSR-A process is given in the Supporting Information. For fixed values of these factors, the ratio CTCI/CTBM is constant, independent of the total basemodule cost (i.e., CTBM). Thus, for the other two process alternatives, the CTCI is calculated with eq 16. 6.3. Comparison. To aid the comparison between the three process alternatives developed in this study, the key economic indicators (calculated using the procedure outlined in section 6.1), and some important process parameters (retrieved from section 4), are collected in Table 12. On the basis of this information, the following comparisons are made:

7. CONCLUSIONS 2-Ethylhexyl acrylate production at the industrial scale using conventional reaction−separation−recycle systems and solid catalysts is feasible. Three process alternatives are evaluated. Each process employs a fixed bed reactor with Amberlyst A70, and standard distillation equipment for separation. The nominal operating points are selected based on rigorous sensitivity analyses, having the aim to reduce the heating Q



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requirement (steam consumption) and to obtain high wastewater purity (high degree of reactants utilization). The RSR-A process has slightly better key performance indicators. The reactor achieves a space-time-yield of 1.57 kgacrylate/kgcat/h, designed to meet a production capacity of 20 000 t/a. The separation section is operated at vacuum and has two distillation column systems (DC‑1/C‑2 = 0.6/2.0 m, NT,C‑1/C‑2 = 10/25 theoretical trays), both using MellapakPlus structured packing. The separation sequence is direct; the water is removed first, at 99.8% mass purity; second, the acid and alcohol are separated and recycled to the reaction section, while the acrylate is recovered at the required 99.5% mass purity. The dynamic simulations show a robust control structure when faced with various process changes: ±25% capacity, 5% mass contamination of fresh reactants, and ±5 °C temperature at the reactor inlet. To achieve tight control of the product purity, control based on concentration measurement is required. However, a rather slow analyzer is necessary, since the control loop is robust enough when 30 min sampling interval +30 min dead time are allocated to the concentration measurements. The economic analysis shows positive economic indicators. The net earnings are estimated at 8,125,000 $/a of acrylate. The product sales bring 45,375,000 $/a, while the overall production costs are 37,250,000 $/a. The largest share of the production costs is attributed to the cost of reactants with 35,535,000 $/a, followed by the cost of depreciation with 1,014,000 $/a, and the cost of utilities with 700,000 $/a. A breakdown of the depreciation cost shows that 52.8% is taken by columns, 32.0% by heat exchangers, 12.5% by the reactor, and 2.7% by other vessels. For utilities, 88.5% is payed for steam, 7.8% for wastewater treatment, and 3.8% for cooling water.

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REFERENCES

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ASSOCIATED CONTENT

* Supporting Information S

The Supporting Information is available free of charge on the ACS Publications website at DOI: 10.1021/acs.iecr.7b04752. Experimental data retrieved from literature supporting the kinetic parameters estimation; controller tuning parameters of all three processes: RSR-A, RSR-B, and RSR-C; calculation of the total capital investment (CTCI) from the total installed cost (CTBM) of process equipment (example for the RSR-A process) (PDF)

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AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. ORCID

Mihai Daniel Moraru: 0000-0001-9223-9913 Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS C.S.B. gratefully acknowledges the financial support of the European Commission through the European Regional Development Fund and of the Romanian state budget, under the Grant Agreement 155/25.11.2016 (Project POC P-37-449, acronym ASPiRE). R


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Design, Control, and Economic Evaluation - ACS Publications

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