Enhanced Simulated Moving Bed Reactor Process for Butyl Acrylate

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Enhanced Simulated Moving Bed Reactor Process for Butyl Acrylate Synthesis: Process Analysis and Optimization Dânia S. M. Constantino, Rui P. V. Faria, Carla S. M. Pereira, José Miguel Loureiro, and Alirio Egidio Rodrigues Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.6b02474 • Publication Date (Web): 16 Sep 2016 Downloaded from http://pubs.acs.org on September 16, 2016

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Industrial & Engineering Chemistry Research

Enhanced Simulated Moving Bed Reactor Process for Butyl Acrylate Synthesis: Process Analysis and Optimization

Dânia S.M. Constantino, Rui P. V. Faria*, Carla S.M. Pereira, José M. Loureiro and Alírio. E. Rodrigues Laboratory of Separation and Reaction Engineering - Laboratory of Catalysis and Materials (LSRELCM), Department of Chemical Engineering, Faculty of Engineering, University of Porto, Rua Dr. Roberto Frias, 4200-465 Porto, Portugal

*Corresponding author. Tel.: +351225081578; fax: +351225081674. E-mail address: [email protected] (Rui P.V. Faria).

Abstract A novel process design based on the simulated moving bed technology for the synthesis of butyl acrylate (BAc) was investigated in order to get a more competitive industrial process. For that, a fixedbed reactor was coupled with a simulated moving bed reactor (SMBR). Reactive separation regions were determined for different conditions of process configurations and feed compositions allowing to find the optimal operating parameters. Furthermore, the process integration was accomplished and the desorbent (n-butanol) recovery was also investigated ensuring the minimal BAc purity required (99.5 % w/w). The viability and competitiveness of this process were evaluated after an economic analysis which showed that it required similar production costs and energy consumption as well for the highest

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production capacity when compared with other state-of-the-art processes for the BAc synthesis, presented so far.

Keywords Butyl acrylate, process design, process intensification, process optimization, chromatographic reactors.

1. Introduction Market research studies have reported a significant growth of the global demand of butyl acrylate (BAc) during the last decade and this trend is expected to increase even further in the coming years1. Therefore, nowadays the BAc synthesis is a very attractive research and development topic with particular focus on the improvement of the conventional multistage process which consists of two reactors homogeneously catalyzed and three distillation columns for the recovery of the reactants and purification of the final product2. According to the literature2, the main challenges of this process include the presence of azeotropes and the high risk of polymerization at high temperatures (> 398 K). Moreover, this system involves an equilibrium-limited reaction, the esterification of n-butanol (nBuOH) with acrylic acid (AAc). Chromatographic separation processes combined with reaction operated in a continuous mode in a single unit are a very attractive strategy from an industrial point of view, because it enables overcoming the conversions imposed by the thermodynamic equilibrium (since reaction and separation occur simultaneously) while the energy costs and equipment size are reduced3,4. Furthermore, lower temperatures are generally required in comparison with the reactive distillation (RD) technology, for instance. Meanwhile, Simulated Moving Bed Reactor (SMBR) has been intensively investigated for several chemicals production disclosing good sustainability5-15. A detailed study of BAc synthesis over Amberlyst-15 ion exchange resin in SMBR was reported in our previous publication5, where very promising results were presented revealing a competitive production capacity at industrial scale (51,500 tBAc/y). For that, the ideal operating parameters to achieve high separation 2 ACS Paragon Plus Environment

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performance with reduced costs were found according to the reactive-separation region (RSR) determined. The data showed that the separation region (SR) is much larger than the respective RSR allowing to conclude that the process is severely limited by the chemical reaction due to the slow kinetic of this system, as mentioned in literature16. Accordingly, the reactants, n-BuOH and AAc, require long residence time in sections 2 and 3 of the SMBR (between the extract and raffinate streams) in order to completely convert the limiting reactant (AAc). Otherwise, this compound will preferentially contaminate the raffinate stream, since the resin selectivity between AAc and BAc is smaller than the between AAc and water. This fact implies that the process can only operate under low feed flow rates leading to a negligible or not competitive production capacity. Recently, SeidelMorgenstern and his group reported a system for the continuous synthesis and purification of complex reaction mixtures using the SMB technology. A layout based on a tubular reactor with a multi-column chromatographic separation process was successfully studied for the nucleophilic aromatic substitution reaction of 2,4-difluoronitrobenzene with morpholine under continuous flow conditions17. However, as far as our knowledge goes, up to now no esterification reaction has been investigated with this methodology. In this work, the study of a similar configuration is performed for the first time in an attempt to overcome the limitations imposed by the slow reaction kinetics of the esterification reaction between n-BuOH and AAc, instead of the nucleophilic aromatic substitution reaction studied by SeidelMorgenstern et al.17. The main objective is to attain a better performance (less desorbent consumption, DesC and higher productivity, PROD with the same target product purity, ≥ 99.5%) than the obtained with the previously studied technologies (Reactive distillation (RD) and the conventional SMBR processes). This novel concept consists in coupling a fixed-bed reactor (FBR) with a SMBR unit. Thus, the reaction takes places firstly in the FBR and after the reaction equilibrium is reached, the outlet stream of FBR is introduced in the SMBR unit, which works, essentially, as separation unit turning it into a more efficient process (with higher purity with less costs) according to the SR determined in advance for this system5. This design strategy tends to be even more efficient since it is 3 ACS Paragon Plus Environment


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possible to achieve the complete conversion of unreacted AAc from the FBR, by extending the reaction with the eluent/reactant (n-BuOH) that is inside of the SMBR, while the separation occurs, simultaneously. Other advantage of this configuration is that it allows operate under higher feed flow rates than in the previous process (single SMBR) leading to higher production capacities and lower desorbent consumptions. Finally, the process optimization was performed considering the energy consumption for different eluent recovery strategies.

2. Results and Discussion 2.1. Process Design An industrial process for the production of BAc using the SMBR technology has recently been reported in the open literature as an sustainable strategy for the production of this ester5. Despite the promising performance of this technology some limitations have also been detected by the authors. In fact, when the RSR of a SMBR unit fed by an equimolar mixture of AAc and n-BuOH (the reactants for producing BAc) was compared with a hypothetical SR of a purely separative SMB unit fed by an equimolar of BAc and water (the products of this esterification reaction) it was clear that the performance of the unit was severely conditioned by the equilibrium-limited reaction. Indeed, according to the literature16, this system presents a very slow kinetics, 1 kc,363K  3.51103 mol.g A15 .min 1 , so long residence times in the reaction zone of the SMBR (sections

2 and 3) are required which implies working under low feed flow rates in order to achieve a reasonable reactants conversions, leading to a lower production capacity. Therefore, the SMB presents a better performance than the SMBR. It has a much higher productivity and it needs less desorbent to reach the same final product purity. As a consequence of these results, it was considered that the process should be revaluated in order to find a new process design that could take full advantage of the potential of the SMBR technology for the synthesis of BAc. The new process design proposal consists in adding a


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FBR before the SMBR, as shown in Figure 1 (PDB), keeping the remaining units as suggested in the previous process design studied5. The esterification reaction between AAc and n-BuOH will take place, firstly, in the FBR and, according to Figure 1, the outlet stream of the FBR at equilibrium conditions is directly connected to the SMBR where the separation of the extract and raffinate streams will occur. Furthermore, almost all unreacted AAc from the first reactor (FBR) will be consumed in the second one (SMBR), where there is n-BuOH in excess, since this reactant is also used as eluent/desorbent. Then, a pervaporation unit (PERV) and a distillation column (DC) are used as separation units for the outlet streams treatment, extract and raffinate, respectively. Thereby, the operating parameters of each unit were optimized to assess if it would be possible to achieve a better performance than the obtained previously with PDA (conventional SMBR process)5: same raffinate purity (PR ≥ 99.5 %) with higher productivity and less desorbent consumption. Each unit was optimized in terms of dimensions and operating parameters, considering different scenarios for eluent recovery, which will be discussed ahead. However, the dimensions of SMBR at industrial scale obtained in PDA were kept in order to set a reference for performance data comparison (12 columns with a length of 0.615 m and a diameter of 2.23 m).



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Figure 1. New process plant design for BAc synthesis (PDB).

2.2. SMBR Optimization The determination of the design parameters is extremely important to evaluate the process feasibility. Therefore, firstly, it is extremely important to know its design and configuration. Basically, it is composed by several fixed-beds interconnected (12 in this case) which are packed with a heterogeneous catalyst that acts as adsorbent, simultaneously. The SMBR technology is based on the True Moving Bed Reactor (TMBR) technology where a counter-current between the solid and the fluid phases is promoted, but in SMBR case the solid phase movement is simulated by the periodical shift of the outlet and inlet streams in regular time intervals set by the operator (switching time), according to the desired solid velocity. The reactants (n-Butanol and AAc) are fed to the reactor and one of them (n-butanol in this case) is also used as desorbent for adsorbent regeneration. Thenceforward, the products (BAc and Water). Due to their different affinities towards the stationary phase, a mixture containing the desorbent and the most retained product (n-Butanol and Water) is obtained in the extract stream and a mixture containing the desorbent and the less retained product (nButanol and BAc) is obtained in the raffinate. The position of the inlet and outlet streams of the SMBR divides the unit into four sections, which may consist of different numbers of columns (fixedbeds) allowing the use of several configurations: In Section 1, the solid is regenerated by desorption of the most adsorbed product (Water) using the desorbent (n-Butanol). It is located between the extract and the desorbent nodes. In Section 2, located between the extract and feed nodes, and section 3, located between the feed and raffinate nodes, the reaction takes place and the products (BAc and Water) are separated as they are formed.


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In Section 4 the desorbent is regenerated by adsorption of the less retained product (BAc) before being recycle to section 1, is located between the raffinate and desorbent nodes18. As previously mentioned, in order to compare the production capacity of the preceding and actual process designs (PDA and PDB, respectively), equivalent SMBR dimensions were considered in both studies. In the previous case, the optimization was performed by determining the RSR at different conditions. The best performance was achieved by using a configuration 2-4-4-2, two columns for section 1, four columns for section 2 and 3 and two columns for section 4, respectively, and a switching time of 3.1 minutes at 363 K, reaching a limiting reactant conversion of 99.5 % and a raffinate purity of 99.9 %. The annual production capacity reached with PDA was about of 51,500 tonBAc/year, consuming 1.0 Ln-BuOH/kgBAc (after recycling all possible eluent from the raffinate and the extract streams). In this case, the same procedure was followed for the process optimization. The RSR were determined for different SMBR configurations, considering that the feed stream is composed by a reaction mixture in equilibrium conditions coming from the FBR instead of an equimolar composition of the reactants as considered in PDA. For each configuration, the optimal point of the respective RSR was determined by changing 𝛾2 and 𝛾3 (ratio of interstitial liquid and simulated solid velocities in the section 2 and 3 of the SMBR, respectively), so that the maximum feed flow rate that allows to obtain the desired separation (BAc purity ≥ 99.5 %, free solvent basis) is achieved, which corresponds to the vertex of the region18. The RSR determined for PDB are presented in the Figure 2, as well as the RSR obtained for PDA in order to allow a direct comparison between the two process designs. From this figure, it is possible to conclude that coupling the FBR to the SMBR can enhance the BAc production. Comparing the different configurations studied for the PDB it is noticeable that the most efficient performance is achieved when the configuration 2-3-5-2 is used. Due to the relatively low selectivity of the resin between BAc and AAc it is necessary to increase the number of columns in section 3 in order to avoid the contamination of the raffinate stream with AAc. That way, this configuration is more suitable for this system allowing to work with higher feed flow rate ensuring the required purity specifications leading to higher productivity and, accordingly, less desorbent 7 ACS Paragon Plus Environment


consumption. On the contrary, for PDA, both configurations presented similar performances, and so four columns in section 3 is enough to achieve the maximum BAc-AAc separation and AAc conversion for the optimum feed compositions in PDA (equimolar reactants mixture). PDA_2442

PDB_2442

PDB_2352

2.5

2.3 2.1

g3

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1.9 1.7 1.5 1.5

1.7

1.9

g2

2.1

2.3

2.5

Figure 2. Optimization of Reactive Separation Region (RSR) for Process Design B (PDB-enhanced SMBR process) at two different configurations (2-4-4-2 and 2-3-5-2) and optimal RSR for PDA (conventional SMBR process). Aiming to maximize the process performance, the optimal operating parameters were further investigated by performing a sensitivity analysis to the flow rate ratios between the liquid and the solid phases of sections 1 and 4, critical sections for the SMBR operation. The Equilibrium Theory, the most widely used methodology for the determination of these parameters, assuming that no chemical reaction takes place in that sections (1 and 4), determines a value of 9.10 for the flow rate ratio in section 1 and 1.52 for section 4 according to the equations (1) and (2), respectively, as described in our previous work5. The sensitivity analysis was carried out by varying 𝛾1 from -6 % to 4 % and 𝛾4 from 2 % to 2 %. Since the SMBR dimensions and the switching time (𝑡 ∗ ) were kept constant, the solid velocity (𝑈𝑠 ) remains constant. So, the flow rate ratios in each section were changed by varying the liquid flow rate in that section, according to the equations from (3) to (8). For each pair of 𝛾1 and 𝛾4 a new RSR was determined. The respective optimal points in terms of productivity and desorbent 8 ACS Paragon Plus Environment

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consumption are presented in Figure 3, which shows that the effect on the process productivity promoted by changing 𝛾1 in the range of -4% to 4% relative to the equilibrium theory value (keeping 𝛾4 ) is almost negligible; however, looking to the desorbent consumption (right), there is an evident minimum consumption when 𝛾1 is decreased by 4 % and the Equilibrium Theory value is kept for 𝛾4 .

g1 

qD (1   )    p  (1   p ) 1   C p , D1

   

(1)

g4 

qC (1   )    p  (1   p ) 4   C p ,C4

   

(2)

us 

L , where L is the column length t*

gj 

vj us



Qj A us

, where  j and

(3)

Q j represent the liquid velocity and the volumetric (4)

flow rate in the section j, respectively. In turn, the flow rate of each section is given by the following equations:

Q1  QD  QRe c

(5)



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Q2  Q1  QX

(6)

Q3  Q2  QF

(7)

Q4  Q3  QR  QRe c

(8)

PROD (kg BAc/(L ads.day)

DesC (L But/kg BAc)

12.0

17.0

11.5

16.5

11.0

16.0

10.5

15.5

10.0

15.0 4%

2%

0%

-2% Ɣ1

-4%

-6%

4%

2%

0% -2% -4% -6% Ɣ1

Figure 3. Sensitivity analysis to 𝛾1 and 𝛾4 using the configuration 2-3-5-2: Productivity data (kBAc/LAds.day) on left and desorbent consumption (LBuOH/kgBAc) on right. The operating and performance parameters obtained along this study are summarized in Table 1 together with the respective values previously determined for the PDA (first column in same table). The second and third columns of Table 1 correspond to the vertex of the RSR (optimal point) using the configurations 2-4-4-2 and 2-3-5-2, respectively. With the 2-3-5-2 configuration it is possible to increase the productivity by 25 (%) and to decrease the desorbent consumption by approximately 20 %. After the sensitivity analysis performed to 𝛾1 and 𝛾4 , it was possible to decrease the desorbent consumption even further (5 % less BuOH is needed). Nevertheless, if these operating parameters were implemented at real industrial scale, this process would be working under a very sensitive set of conditions with respect to the raffinate purity. Any disturbance on the optimal operating parameters (such as slight variations of the feed stream composition or fluctuations on the flow rate of the inlet and outlet streams of the SMBR, for instance) might lead to the contamination of the raffinate, making it impossible to obtain BAc with the required specifications (minimum purity of 99.5 %). Therefore, a 10 ACS Paragon Plus Environment

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security factor (S.F.) must be applied to some of the most relevant operating parameters (feed and extract flow rates) in order to get a reliable set of operating conditions, which is presented on right column in Table 1 ( 2-3-5-2 after sensitivity analysis with S.F.). Accordingly, in this case, a S.F. of (0.5) % and (-1.5) % were applied to 𝛾2 and 𝛾3 , respectively. Table 1. Operating and performance parameters optimized for PDB vs PDA. PDA 2-4-4-2

2-4-4-2

2-3-5-2

QF (L/min)

118

110

138

PDB 2-3-5-2 after sensitivity analysis 138

QX (L/min)

2315

2265

2256

2145

2148

QD (L/min)

2397

2390

2390

2280

2280

QRec (L/min)

475

478

478

478

478

Cn-ButOH,F (mol/L) CAAc,F (mol/L) CBAc,F (mol/L) CWater,F (mol/L)

5.76 5.76 -

0.57 4.56 3.47 3.47

0.57 4.56 3.47 3.47

0.57 4.56 3.47 3.47

0.57 4.56 3.47 3.47

PX (%) PR (%)

99.3 99.8

99.5 99.6

99.6 99.6

99.6 99.5

99.6 99.7

Conv. (%)

99.5

98.9

99.0

99.0

99.3

PROD (kgBAc/(Lads.day)

7.20

9.39

11.8

11.8

11.1

DesC (LBut/kgBAc)

27.6

20.8

16.5

15.7

16.7

Parameters/Configuration

2-3-5-2 after sensitivity analysis with S.F. 130

The optimal operating conditions of the PDB (on the right column of Table 1) were used to simulate the operation and obtain the respective concentration profiles which are presented in the Figure 4 together with the concentration profile obtained with the PDA. Comparing both profiles, it can be concluded that there was an evident extension of the reaction in PDB, as intended. In section 3, it is possible to observe a sudden decrease in n-BuOH concentration, since this is the section where the reaction between the AAc and n-BuOH occurs yielding BAc and water. Furthermore, according to the adsorption data already published19, Amberlyst-15 is more selective to water than to BAc and because



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of that this compound is present inside the unit at higher concentrations. Indeed, this is one of the advantages of SMBR since the reaction and adsorption occur simultaneously.

Figure 4. Concentration profiles in the SMBR unit at cyclic steady state (21st cycle) and 363 K using 𝑡 ∗ = 3.1 min (all operating conditions are presented in Table 1). Grey profile corresponds to the PDA case study5.

2.3. Process Integration and Optimization The process integration was studied in order to assess the feasibility of the global process design suggested in this work. For that, all remaining units presented in Figure 1 were dimensioned: the FBR and the separation units to treat the outlet streams of SMBR. All data are presented below. 2.3.1

Dimensioning of FBR

A FBR operating at 363 K was simulated (according to the mathematical model described in our previous work 19) assuming that a n-BuOH/AAc mixture with a molar ratio of (1:2) was fed at the same optimal feed flow rate found for the SMBR, 130 L/min. Contrarily to the PDA, in this case it is possible to use AAc in excess on feed stream, since it enters diluted in SMBR (outlet stream of FBR)


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avoiding raffinate contamination with AAc, as shown in the concentration profiles presented in Figure 4. The respective dimensions needed to reach the equilibrium conversion were determined and they are presented in Table 2, together with the characteristics of the adsorbent/catalyst used. The ratio between the length and the diameter (L/D) of each column of the SMBR was kept for FBR to avoid differences on the bed properties. The reaction and adsorption data for this system using the ion exchange resin Amberlyst-15 as stationary phase is available in the open literature16,19. The concentration histories are shown in Figure 5. Initially the reactor is saturated with the eluent (n-BuOH) that acts also as reactant. As the feed mixture enters the reactor and the AAc is consumed producing BAc and water. When the stationary stage is achieved, a mixture almost at equilibrium conditions is obtained, composed by 4.8 % of nBuOH, 37.8 % of AAc and 28.7 % of BAc and water. This stream is then fed to the SMBR. Table 2. Characteristics of the FBR and resin (Amberlyst-15) used in simulation runs. Column length (L), m

2.00

Column internal diameter (Di), m

7.25

Resin particle porosity (єp) Bed porosity (є)

0.360 0.400

Resin Particle radius (rp), µm

375

Bulk density (ρb), kg m

390

-3



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Figure 5. Concentration histories at the outlet of the FBR initially saturated with n-butanol and fed with a mixtures of n-BuOH/AAc (Cn-BUOH,F = 4.06 mol/L and CAAc,F = 8.06 mol/L); QF = 130 L/min and T = 363 K. 2.3.2

Eluent Recovery and Economic Evaluation

Additional separation units were studied to treat the outlet streams of the SMBR, the raffinate and extract streams, composed by n-BuOH and BAc and n-BuOH and water, respectively. After analyzing the concentration profile of the SMBR and determining its outlet streams compositions it is possible to design the complementary separation units in order to get the desired purity product (BAc, ≥ 99.5 %). Furthermore, through the process integration will be possible to study the eluent recovery enabling to evaluate the process sustainability. The raffinate stream of the SMBR is composed by 49.1 % (n/n) of n-BuOH and 50.7 % (n/n) of BAc. The remaining composition corresponds to impurities of AAc and water (0.02%). A distillation column (DC) was considered to treat this stream and a similar procedure to the previous work5 was followed aiming to determine the optimized parameters required to perform the desired separation. Accordingly, Aspen Plus (Version 8.6) was used to simulate this unit using the RadFrac method. In summary, the inputs for this simulation method are the volumetric flow rate (Q), temperature (T), number of trays (N), feed stage (NF), reflux ratio (Ractual) and the bottom to feed ratio besides the molar feed composition. The DC parameters and streams compositions are summarized in the Supporting Information, Tables S1 and S2, respectively. The final parameters were calculated based on equations from (9) to (11) and are presented in Table 3, which shows that the dimensions of the DC required for this process (PDB) are smaller than the dimensions of the DC determined for PDA (34.9 m x 1.55 m instead of 39.7 m x 1.92 m), reducing, simultaneously, the required energy consumption. These facts are due to the differences in the molar compositions of the raffinate stream obtained in SMBR of the different design processes, since BAc is more concentrated in that stream of PDB.


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Under optimized conditions, the BAc purity in the DC bottom stream reached 99.7 % while the nBuOH obtained on top stream presented a purity of 95.6 %. Furthermore, it is possible to conclude that PDB is able to produce 67,000 tBAc/year by feeding 262 L/min (total raffinate flow rate) to this DC. This represents a production capacity 30 % higher than the production capacity of PDA. It is important to note that despite of the high temperatures verified in the reboiler, the risk of polymerization is low or negligible since the polymerization reaction of AAc and BAc is initiated by radicals present in the mixture20 and it is unlikely to occur in DC of this work, because its function is limited to the separation of n-Butanol and BAc in the presence of only trace amounts of AAc..

Nactual  N / Eff

(9)

Htower  Htray  Nactual

(10)

Ltower  1.2  Htower

(11)

The variables Htower, Ltower and Htray represent the tower height, the required length of the column (estimated to be 20 % higher than the required just for trays) and height of each tray (assumed to be 0.7 m), respectively, according to the literature21. Table 3. Final parameters required for the distillation unit. Nactual 41.4

Htower /m 29.0

Ltower/m 34.9

Dtower/m 1.55

Htray/m 0.70

The extract stream from SMBR unit presents a composition of 95.4 % of n-BuOH and 4.6 % of water. Since the amount of water is less than 10 %22, a pervaporation unit (PERV) was considered for the nBuOH dehydration. This unit was investigated according to the mathematical model described in a previous work5. The permeance values, Qmemb,i ,were calculated and the values obtained at 363 K were 0.207 kg/ (m2 h bar) for n-BuOH and 16.4 kg/ (m2 h bar) for water on literature data23. After that, the PERV simulation 15 ACS Paragon Plus Environment


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was performed by feeding a flow rate of 2148 L/min which corresponds to the extract flow rate from SMBR unit. As result, a module of 785 m2 is needed, allowing to remove 97.70 % of water and to reach 99.97 wt % of n-BuOH purity (retentate). The parameters required to determine the permeance as well as the molar flux permeate, 𝐽𝑖 , for each compound, which are essential to carry out the PERV process simulation, are summarized in the Supporting Information (Table S3). The n-BuOH recovery was investigated taking into account the recycling of the outlet streams obtained from the separation units. However, two different scenarios were considered, ER1 and ER2, which are presented in Figure 6 and Figure 7, respectively. The most relevant stream for the recovery eluent is the outlet stream from the PERV unit, representing the highest saving in eluent and energy consumption. Basically, the scenario ER1 consists in recycling the n-BuOH from the PERV unit (11) to the eluent stream of SMBR (stream 6). In the second scenario (ER2), besides the eluent recovery strategy adopted in scenario ER1, 48.5 % of the top stream of the DC is reintroduced in the process through the feed stream of the FBR. The remaining is discarded to ensure the required molar feed ratio of n-BuOH and AAc (1:2), in order to get the same molar feed composition for the SMBR. That way, the total amount of fresh n-BuOH required in the process (which is introduced through stream 1 and 8) can be significantly reduced. The different flow rates and the respective concentrations for each stream according to the different scenarios ER1 and ER2 are shown in Supporting Information (Tables S4 and S5, respectively) as well as the global performances parameters including the energy and desorbent consumption (Table S6). The results shows that it is possible to save about 91% of the eluent in relation to the open process (without recycle of n-BuOH) while the scenario ER2 allows to recover about 93%. Accordingly, the energy costs were estimated for each scenario, considering the enthalpies of each stream. The reactors were considered to operate under isothermal conditions, the reboiler duty was obtained from the simulations results in Aspen Plus (Version 8.6) and the energy required for the PERV unit was determined by using the non-isothermal model described in a previous work5. All


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thermophysical data necessary were found in the open literature24. As expected, the ER2 showed to be the most economic choice, presenting a cost of 1.85  103 kJ/kgBAc against 1.93  103 kJ/kgBAc for the ER1 and 3.77  103 kJ/kgBAc for the open process, requiring practically the same energy of PDA with similar strategy for eluent recovery (1.80  103 kJ/kgBAc)5, although this actual process (PDB) presents higher production capacity (30 %).

n-BuOH

(1)

AAc

(2)

(3) (12)

(FN)

Water FBR

(5)

Perv (11)

(4)

SMBR

(6)

(8) (EN) (9)

(7)

n-BuOH (E) n-BuOH 96 %

DC

(10)

BAc

Figure 6. Configuration for n-BuOH recycle using the outlet stream of the pervaporation unit (Scenario ER1). FN and EN represent the feed and eluent nodes, respectively.



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n-BuOH

(1)

Page 18 of 28

AAc

(2)

(3) (FN)

(12)

(9a)

Water

FBR (5)

Perv (11)

(4)

SMBR

(6)

(EN)

(9) DC

(8)

n-BuOH (E)

(7) (10)

BAc

Figure 7. Configuration for n-BuOH recycle using the outlet stream of the pervaporation unit and 48.5 % of the top stream of the distillation column (Scenario ER2). To access the viability of the process presented in this work, an economic analysis was performed by determining the different costs involved in this process in order to compare it with others processes presented so far. The total CAPEX includes the investment for the FBR (7.9 %), SMBR (10.2 %), PERV (24.5 %), DC (15.9 %), trays (1.8 %), reboiler (5.8%), condenser (6.0 %), heat exchangers (25.4 %), all pumps required (1.4 %) and a tank for the feed mixture (1.2 %). All costs were determined based on equations described in Dimian A. C.25, with exception of the cost of the pumps which were achieved by following a procedure based on Timmerhaus and Peters26, while the yearly cost of the pumps were estimated based on Turton et al.27. All the equipment was considered to be made of stainless steel. The cost of reactants, CREACTANTS, was estimated based on the prices of AAc (614.5 €/t) and n-BuOH (628.8 €/t) given by the literature2. The costs associated with utilities required by the process, CUtilities, namely, steam (72.2%), cooling water (4.1%), waste water treatment (22.4 %) 18 ACS Paragon Plus Environment

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and electricity (1.3 %) were also considered. For that, the prices reported by Moraru et al.28 were taken into account: 7.78 (6 bar) $/GJ for steam, 0.72 $/GJ for cooling water and 0.5 $/kgorganics for waste water. An exchange rate of 1.11 ($/€) was used. For the electricity cost estimation a value of 0.125 €/kWh was considered according to the Eurostat data (half-yearly electricity prices) for industry from European Commission home page29 .The Marshall and Swift equipment cost index of 2011 considered was 1536.5 according to the literature28. The production costs, CProduction, were also calculated and include the amortization costs (0.9 %), CAmortization, the utilities costs (2.4 %), the reactants costs (95.0 %) as well as the packing of the reactors (1.7 %). The price of Amberlyst-15 was calculated based on the Sigma-Aldrich’s price (70.4 €/kg). The amount of catalyst/adsorbent required is 43.5 tons and a replacement was considered at the end of 3 years. Finally, the cost of the product, C Product (€/t), was obtained from equation (12) and CAmortization from equation (13). The payback period of all processes were standardized for 7 years.

CPr oduct  CPr oduction / prod . capacity

(12)

CAmortization  CAPEX / payback period

(13)

All the costs associated with the different processes that have been proposed for the BAc synthesis are presented in Table 5: SMBR (PDA)5, RD with decanter2, an industrial process for BAc synthesis recently published based on Reactor-Separation-Recycle28 and the process proposed in this work comprising a FBR and a SMBR (PDB). According to the results, the CAPEX of PDB is slightly larger than our previously studied process (PDA) since it requires an additional reactor (FBR). Furthermore, in terms of CAPEX, the values show that the RD with a decanter and the Reactor-Separation-Recycle processes are more expansive than any of the SMBR based processes, PDA and PDB. Despite the fact that different methodologies might have been used for the estimation of the production costs, the process proposed in this work seems to be competitive presenting a final cost of the product significantly lower than its current price, 1350 €/tBAc2 and an energy consumption similar to the



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Page 20 of 28

remaining processes for a much larger production capacity (67,000 tBAc/y). From all the considered processes PDA seems to be the most economically attractive solution however the differences between the values estimated are relatively small, with a variation of only 20%, approximately, which might be within the range of error of the methodologies applied for the costs estimations. Table 5. Economic analysis for the different processes for BAc synthesis.

PDA5

PDB (This work)

RD with decanter2

ReactorSeparationRecycle28

Prod. capacity (  103) (tBAc/y)

51.5

67.0

20.0

20.5

Energy consumption (  10 ) (kJ/kgBAc)

1.80

1.87

1.69

-

CAPEX (€/tBAc/y)

69.5

57.6

238

191

CReactants (€/tBAc)

725

866

743

707

CUtilities (€/tBAc)

21.7

22.1

76.6

101

CAmortization (€/tBAc)

9.94

8.23

34.4

27.3

CProduct (€/tBAc)

762

911

855

836

Process parameters

3

3. Conclusions An enhanced SMBR process for BAc synthesis over Amberlyst-15 was studied by coupling a FBR to a SMBR unit operating at 363 K. The viability of this novel process design was investigated at industrial scale, by modeling the process according to the flow diagram proposed. The design and operating parameters were found ensuring the desired separation and complete regeneration of the resin and eluent in sections 1 and 4, respectively. The Reactive Separation Regions were determined for different SMBR configurations and the unit was optimized by performing a sensitivity analysis to the flow rates ratios determined from the Equilibrium Theory. The best performance was obtained, by using a molar feed ratio of (1:2) (n-BuOH:AAc) in the FBR and a 2-3-5-2 configuration in the SMBR, achieving the maximum productivity with the minimum desorbent/eluent (n-BuOH) consumption for BAc purity required. The most suitable operating flow rates found were: QD = 2280 L/min, QRec = 478 L/min, QF = 130 L/min and QX = 2148 L/min. 20 ACS Paragon Plus Environment

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A Distillation Column and a Membrane Pervaporation unit were dimensioned and optimized to treat the SMBR raffinate and the extract streams, respectively. In this case, a PERV unit of 785 m2 and a DC with 34.9 m of length (42 trays) and 1.55 m of diameter are enough to achieve the purification specifications. Smaller separation units are required for this process design when compared with the conventional SMBR process, since both raffinate and extract flow rates are lower in PDB. However, the BAc concentration in the raffinate stream is higher when the FBR is used before the SMBR, which simplifies the purification step and increases the process productivity. Additionally, the global process was studied considering the eluent recovery. The most economically attractive process was achieved by recycling the n-BuOH from the PERV unit to the eluent stream of the SMBR and using 48.5 % of the top current of DC in the FBR feed (scenario ER2). According to these results, it is possible to reach an eluent recovery of 93 %. Finally, an economic analysis was performed showing the viability of this process design and its competitiveness, since it presented the highest production capacity for a similar energy demand and production costs comparing with other state-of-the-art processes for the BAc synthesis, presented so far. Supporting Information The DC input parameters and respective streams compositions are summarized as well as the parameters required to determine the permeance and the molar flux permeate for each compound. Furthermore, the different flow rates and the respective concentrations for each stream according to the different eluent recovery scenarios, ER1 and ER2, are shown in Supporting Information. Finally, the global performances parameters including the energy and desorbent consumption are compared for the different processes studied. This information is available free of charge via the Internet at http://pubs.acs.org.



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Acknowledgments The authors acknowledge financial support provided by Fundação para a Ciência e a Tecnologia (FCT, Portugal) and FEDER under Program COMPETE through the Project FCOMP-01-0124-FEDER027458 (FCT Project EXCL/QEQ-PRS/0308/2012). This work was financially supported by Project POCI-01-0145-FEDER-006984 - Associated Laboratory LSRE-LCM funded by FEDER funds through COMPETE2020 - Programa Operacional Competitividade e Internacionalização (POCI) - and by national funds through FCT. This work was also co-financed by QREN, ON2 and FEDER (Project NORTE-07-0162-FEDER-000050 and Project NORTE-07-0124-FEDER-0000007).

Notation Abbreviations AAc Acrylic Acid

-

BAc

Butyl Acrylate

-

Conv

Reaction Conversion

%

DC

Distillation Column

-

DesC

Desorbent consumption

LBut/kgBAc

FBR

Fixed-Bed Reactor

-

gPROMS

General Process Modeling System

-

n-BuOH

n-butanol

-

PERV

Pervaporation

-

PDA

Process Design A

-

PDB

Process Design B

-

PR

Raffinate purity

%

PX

Extract purity

%

Prod

Productivity

kgBac/(Lads day)

RD

Reactive Distillation

-

RSR

Reactive Separation Region

-

SF

Security Factor

-

SMB

Simulated Moving Bed

-

SMBR

Simulated Moving Bed Reactor

-

SR

Separation Region

-


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Symbols area

m2

liquid phase concentration inside the particle

mol/L

activation energy

J/mol

tray efficiency

-

tower height

m

height of each tray

m

permeate molar flux

kg/ (m2 h)

column length of reactors

m

tower length

m

number of trays

-

feed tray

-

solid phase concentration

mol/Lres

volumetric flow rate

L/min

permeance of membrane

kg/ (m2 h bar)

flux ratio

-

particle radius

m

switching time

min

temperature

K

Solid velocity

m/s



Liquid velocity

m/s

x

liquid phase molar fraction

-

vapor phase molar fraction

-

A

Cp E Eff

H tower H tray

J

L Ltower

N NF q Q

Qmemb

R

rp t* T

us

y

Greek Letters interstitial velocities ratio g bulk porosity 

p b  m 

-

catalyst/adsorbent particle porosity

-

bulk density

Kg/m3

Selectivity relative to pervaporation process

-



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Subscripts 0

relative to initial conditions

D

desorbent

F

relative to feed

i

relative to a specie

j

relative to a section of SMBR unit

out

at the outlet of the fixed-bed column

p

relative to particle

ret

retentate

perm

permeate

References (1) TranTech Consultant, Inc. Home Page. http://chemplan.biz/home.aspx. Chemical Profile: Butyl Acrylate (accessed November 2016). (2) Niesbach, A.; Kuhlmann, H.; Keller, T.; Lutze, P.; Górak, A. Optimisation of Industrial-Scale nButyl Acrylate Production using Reactive Distillation. Chem. Eng. Sci. 2013, 100, 360. (3) Stankiewicz, A. I.; Moulijn, J. A. Process intensification: Transforming Chemical Engineering. Chem. Eng. Prog. 2000, 96, 22. (4) Stankiewicz, A. Reactive Separations for Process Intensification: an Industrial Perspective. Chem. Eng. Process. 2003, 42, 137. (5) Constantino, D. S. M.; Pereira, C. S. M.; Faria, R. P. V.; Loureiro, J. M.; Rodrigues, A. E. Simulated Moving Bed Reactor for Butyl Acrylate Synthesis: From Pilot to Industrial Scale. Chem. Eng. Process. 2015, 97, 153. (6) Rodrigues, A. E.; Silva, V. M. T. M. Industrial process for acetals production in a simulated moving bed reactor. U.S. Patent 7,488,851, 2009.


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(7) Pereira, C. S. M.; Zabka, M.; Silva, V. M. T. M.; Rodrigues, A. E. A Novel Process for the Ethyl Lactate Synthesis in a Simulated Moving Bed Reactor (SMBR). Chem. Eng. Sci. 2009, 64, 3301. (8) Minceva, M.; Gomes, P. S.; Meshko, V.; Rodrigues, A. E. Simulated Moving Bed Reactor for Isomerization and Separation of P-xylene. Chem. Eng. J. 2008, 140, 305. (9) Pereira, C. S. M.; Gomes, P. S.; Gandi, G. K.; Silva, V. M. T. M.; Rodrigues, A. E. Multifunctional Reactor for the Synthesis of Dimethylacetal. Ind. Eng. Chem. Res. 2007, 47, 3515. (10) Ströhlein, G.; Mazzotti, M.; Morbidelli, M. Optimal Operation of Simulated-Moving-Bed Reactors for Nonlinear Adsorption Isotherms and Equilibrium Reactions. Chem. Eng. Sci. 2005, 60, 1525. (11) Silva, V. M. T. M.; Rodrigues, A. E. Novel Process for Diethylacetal Synthesis. AIChE J. 2005, 51, 2752. (12) Yu, W.; Hidajat, K.; Ray, A. K. Modeling, Simulation, and Experimental Study of a Simulated Moving Bed Reactor for the Synthesis of Methyl Acetate Ester. Ind. Eng. Chem. Res. 2003, 42, 6743. (13) Zhang, Z.; Hidajat, K.; Ray, A. K. Application of Simulated Countercurrent Moving-Bed Chromatographic Reactor for MTBE Synthesis. Ind. Eng. Chem. Res. 2001, 40, 5305. (14) Kawase, M.; Inoue, Y.; Araki, T.; Hashimoto, K. The Simulated Moving-Bed Reactor for Production of Bisphenol A. Catal. Today 1999, 48, 199. (15) Gyani, V. C.; Reddy, B.; Bhat, R.; Mahajani, S. Simulated Moving Bed Reactor for the Synthesis of 2-Ethylhexyl Acetate. Part I: Experiments and Simulations. Ind. Eng. Chem. Res. 2014, 53, 15811. (16) Ostaniewicz-Cydzik, A. M.; Pereira, C. S. M.; Molga, E.; Rodrigues, A. E. Reaction Kinetics and Thermodynamic Equilibrium for Butyl Acrylate Synthesis from n-Butanol and Acrylic Acid. Ind. Eng. Chem. Res. 2014, 53, 6647.



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(17) Lee, J. W.; Horváth, Z.; O’Brien, A. G.; Seeberger, P. H.; Seidel-Morgenstern, A. Design and Optimization of Coupling a Continuously Operated Reactor with Simulated Moving Bed Chromatography. Chem. Eng. J. 2014, 251, 355. (18) Rodrigues, A. E.; Pereira, C.; Minceva, M.; Pais, L. S.; Ribeiro, A.; Ribeiro, A. M.; Silva, M.; Graça, N.; Santos, J. C. Simulated Moving Bed Technology: Principles, Design and Process Applications; Elsevier Science: Oxford, 2015. (19) Constantino, D. S. M.; Pereira, C. S. M.; Faria, R. P. V.; Ferreira, A. F. P.; Loureiro, J. M.; Rodrigues, A. E. Synthesis of Butyl Acrylate in a Fixed-Bed Adsorptive Reactor over Amberlyst 15. AIChE J. 2015, 61, 1263. (20) Niesbach, A.; Daniels, J.; Schröter, B.; Lutze, P.; Górak, A. The Inhibition of Acrylic Acid and Acrylate Ester Polymerisation in a Heterogeneously Catalysed Pilot-Scale reactive Distillation Column, Chem. Eng. Sci. 2013, 88, 95. (21) Douglas, J. M. Conceptual design of chemical processes; McGraw-Hill: New York, 1988. (22) Kujawski, W. Application of Pervaporation and Vapor Permeation in Environmental Protection. Pol. J. Environ. Stud. 2000, 9, 13. (23) Sommer, S.; Melin, T. Influence of Operation Parameters on the Separation of Mixtures by Pervaporation and Vapor Permeation with Inorganic Membranes. Part 1: Dehydration of Solvents. Chem. Eng. Sci. 2005, 60, 4509. (24) Yaws, C. L. Yaws' Thermophysical Properties of Chemicals and Hydrocarbons (Electronic Edition); Knovel: 2010. (25) Dimian, A. C. Integrated Design and Simulation of Chemical Processes; Elsevier Science: Amesterdam, 2003. (26) Peters, M. S.; Timmerhaus, K. D., Plant Design and Economics for Chemical Engineers; McGraw-Hill: New York, 1991. 26 ACS Paragon Plus Environment

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(27) Turton, R.; Bailie R. C.; Whiting, W. B.; Shaeiwitz, J. A.; Bhattacharyya, D. Analysis, Synthesis, and Design of Chemical Processes; Prentice Hall: New Jersey, 2012. (28) Moraru, M. D.; Milea, A.; Bîldea, C. S. Design and Economic Evaluation of a Process for n-Butyl Acrylate Production. U.P.B. Sci. Bull, Series B. 2016, 78, 113. (29) European Commission Home Page. http://ec.europa.eu (accessed August 2016).

List of Figure Captions Figure 1. New process plant design for BAc synthesis (PDB). Figure 2. Optimization of Reactive Separation Region (RSR) for Process Design B (PDB-enhanced SMBR process) at two different configurations (2-4-4-2 and 2-3-5-2) and optimal RSR for PDA (conventional SMBR process). Figure 3. Sensitivity analysis to 𝜸𝟏 and 𝜸𝟒 using the configuration 2-3-5-2: Productivity data (kBAc/LAds.day) on left and desorbent consumption (LBuOH/kgBAc) on right. Figure 4 - Concentration profiles in the SMBR unit at cyclic steady state (21st cycle) and 363 K using 𝒕∗= 3.1 min (all operating conditions are presented in Table 1). Grey profile corresponds to the PDA case study (5). Figure 5 - Concentration histories at the outlet of the FBR initially saturated with n-butanol and fed with a mixtures of n-BuOH/AAc (Cn-BUOH,F = 4.06 mol/L and CAAc,F = 8.06 mol/L); QF = 130 L/min and T = 363 K. Figure 6 - Configuration for n-BuOH recycle using the outlet stream of the pervaporation unit (Scenario ER1). FN and EN represent the feed and eluent nodes, respectively. Figure 7 - Configuration for n-BuOH recycle using the outlet stream of the pervaporation unit and the top stream of the distillation column (Scenario ER2). 27 ACS Paragon Plus Environment


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For Table of Contents Only


Enhanced Simulated Moving Bed Reactor Process for Butyl Acrylate

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