Reactive Distillation for the Production of Methoxy Propyl Acetate

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Reactive Distillation for the Production of Methoxy Propyl Acetate: Experiments and Simulation Shambala Gadekar-Shinde, Bhoja Reddy, Mohammad Khan, Sanjay Chavan, Daulat Saini, and Sanjay M. Mahajani Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.6b03489 • Publication Date (Web): 10 Jan 2017 Downloaded from http://pubs.acs.org on January 14, 2017

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Industrial & Engineering Chemistry Research

Reactive Distillation for the Production of Methoxy Propyl Acetate: Experiments and Simulation

Shambala Gadekar-Shinde

a,d,e

, Bhoja Reddy b, Mohammad Khan b, Sanjay Chavan

a,e

, Daulat

saini c and Sanjay Mahajani *b a-Sinhgad College of Engineering, Pune, 411041, India b-Department of Chemical Engineering, Indian Institute of Technology, Bombay, Powai, Mumbai, 400 076, India c-National Chemical Laboratory, Pune, 411008, India d-Chemical Engineering Department, Bharati Vidyapeeth University College of Engineering (BVUCOE), Pune, 411043, India e- Savitribai Phule Pune University (SPPU), Pune, 411007, India

* Corresponding Author Email- [email protected]

; [email protected]

1 ACS Paragon Plus Environment


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Abstract

Propylene glycol methyl ether acetate is an industrially important solvent. In this work, we study the applicability of Reactive Distillation (RD) for its synthesis in the presence of ion exchange resin, Amberlyst 15, as a catalyst. Simultaneous separation of water during the course of reaction shifts the reaction in the forward direction, which renders cost effectiveness and compactness to the process. The presence of azeotrope between methoxy propanol and water complicates the separation, leading to a loss of reactants in the product streams, thereby hampering both conversion and purity. Toluene is thus used as an entrainer to further intensify the process. In this work, the intrinsic kinetic parameters are determined from batch reactor data, and used subsequently to simulate the column performance. An experimentally validated simulator is used to examine the effect of various operating and design parameters. Reactant and product losses are minimised to negligible levels and a significant increase in the conversion is realized. Keywords: Reactive distillation, entrainer, methoxy propyl acetate, methoxy propanol, process intensification


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1. Introduction Methoxy propyl acetate, also known as propylene glycol methyl ether acetate, is a high boiling industrial solvent, with applications in paints, inks, dyes, cleaning agents and photo-resist agent etc. The propylene glycol based products have the potential to replace ethylene glycol based products, which have toxic effects on human metabolism.1,2,3 Methoxy propyl acetate (MPA) is one such replacement candidate produced by the esterification of methoxy propanol (MP) with acetic acid, in the presence of an acid catalyst (see equation 1). The literature available on this reaction is scarce; a few articles have reported synthesis of MPA using catalysts like sulphuric acid, p-toluene sulphonic acid, titanium sulphate, heteropolyacids, tungstophosphoric acid and solid acid catalyst etc.

4,5,6,7

Most of these articles give an insight into the applicability and

performance of respective catalysts for MPA synthesis. The classical route of MPA synthesis involves esterification reaction in a kettle, followed by a separation steps, mostly in batch mode. This route is energy-intensive and also requires expensive down-stream operations for catalyst recovery/disposal. The aim of this study is to assess the technical feasibility of using an ion exchange resin (Amberlyst-15) catalyst in a Reactive Distillation (RD) process for the production of MPA. To the best of our knowledge, such work has not been reported till date in literature. Furthermore, none of the reported studies claim close to quantitative conversion.

CH3

O

O CH3

OCH3 OCH3

+

H3C

OH

H3C

O

+

H2O

(1)

OH

1-Methoxy 2- Propanol

Acetic Acid

1-Methoxy 2-Propyl Acetate



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RD, as a process intensification tool, has been successfully used in the production of numerous industrial chemicals.8-10 Simultaneous separation of the products during the course of the reaction can offer one or more benefits that include, positive shift in equilibrium conversion limit, improvement in selectivity, reduction in downstream processing and elimination of thermal instability of reactor. In a reactive distillation system, the presence of homogeneous azeotrope often leads to low conversions and poorer separation. The presence of heterogeneous azeotrope may, however be advantageously exploited to achieve better separation, as in the case of C4-C5 alcohols like nbutanol and iso-amyl alcohol.11 Despite being a C4 alcohol, simultaneous separation is difficult in case of MPA because of the peculiar vapour-liquid equilibrium characteristics, as explained later. The separation and hence the conversion in these types of systems may therefore be enhanced by adding a mass separating agent, called entrainer. Entrainer helps in the efficient removal of one of the products, namely, water in the case of esterification reactions, thereby shifting the reaction equilibrium in the forward direction. Entrainer enhanced RD is a well-researched topic. Dimian et al.12 have reported that an entrainer should form heterogeneous ternary azeotrope for esterification reactions involving C2-C4 alcohols. De Jong et al.13 studied cyclohexane and isopropyl acetate as suitable entrainers for fatty acid esterification in RD and concluded that the feasibility of entrainer based RD is strongly influenced by the kinetics of the reaction. Wang et al.14 studied production of butyl cellosolve acetate with cyclohexane, toluene, octane, ethylene dichloride as entrainers and found octane to be the most cost-effective entrainer. In our earlier work, ethylene dichloride was found to be a suitable entrainer for the synthesis of triacetin in esterification of glycerol15 and for esterification 4 ACS Paragon Plus Environment

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of ethylene glycol with acetic acid.16 Entrainer enhanced RD for the production of ethyl acetate has been reported to yield energy savings of up to 32%.17 In a few other studies from our group18, 19

we found toluene to be the best entrainer for the synthesis of 2 ethyl hexyl acetate and

acetalisation of glycerol with formaldehyde. The present work on methoxy-propyl acetate also uses toluene as an entrainer and evaluates its suitability for the reaction of interest in RD. The selection of entrainer is dependent on factors such as kinetics of the reaction, azeotrope composition, boiling point of azeotrope and VLE characteristics of the multicomponent mixture. A review of potentially important entrainer-based RD processes may be found elsewhere.16 In the present case, the experimental investigations for the selection of the most appropriate entrainer were carried out in a batch reactor, equipped with Dean and Stark assembly. Three entrainers, viz. cyclohexane, butyl acetate and toluene, were investigated. Maximum conversion was realised when toluene was used as an entrainer. Cyclohexane forms a heterogeneous azeotrope with water at 69.8 °C and a homogeneous minimum boiling azeotrope with acetic acid at 79.93 °C. As a result, significant loss of acetic acid was observed in the overhead aqueous phase. Butyl acetate forms a heterogeneous azeotrope with water at 90.9 °C, which is close to the dew/boiling point of the other azeotropes present in the system (see Table 1), thereby making the separation difficult. On the contrary, the temperature of toluene–water heterogeneous azeotrope (84.29 °C) is relatively less than any other azeotropes present in the mixture. Out of all the three entrainers, toluene was thus found to be the most suitable entrainer that offers higher conversion, thermal stability of the catalyst and lower acetic acid loss in the overhead aqueous phase.



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Table 1. Composition of azeotropes and boiling points (°C) of pure components and azeotropes in MP-AA-MPA-water-toluene mixture

Components

Boiling Point

Azeotrope composition

(ºC) Toluene-Water (azeo-heterogeneous)

84.29

mole % Toluene - 44.58; Water 55.42

MP-Water (azeo-homogeneous )

97.31

MP - 16.93 ; Water - 83.07

MPA-Water (azeo-heterogeneous)

97.90

MPA - 7.94; Water - 92.06

MPA-MP-Water (azeo-heterogeneous)

98.16

MPA - 4.51; MP - 3.15; Water - 92.34

Water

100

-

Toluene-Acetic acid (azeo-homogeneous )

103.99

MP-Toluene (azeo-homogeneous)

105

Toluene

110

-

Acetic acid

118

-

MP

118.5

-

MPA

146

-

Toluene - 56.34; Acetic acid - 43.66 MP - 49; Toluene – 51


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The article is organized as follows. The next section details the kinetic studies and the estimation of relevant parameters in the proposed kinetic model. This is followed by the Vapor-LiquidLiquid Equilibrium (VLLE) for the reaction mixture and conceptual design based on residue curve maps, justifying the use of entrainer for this reaction. The laboratory-scale experiments on the RD column are then described along with the simulations in Aspen Plus. The Aspen Plus model was validated against the experimental results and further used to examine the effect of different operating and design parameters and arrive at the best RD configuration.

2. Kinetic studies 2.1 Materials Methoxy propanol (99.5 wt %) was obtained from M/s. Sigma Aldrich. Acetic acid (99.5 wt %) and toluene (99.5 wt %) were supplied by M/s. Merck Ltd., India. The acid catalyst, i.e. dry Amberlyst-15 was supplied by M/s.Rohm and Hass, India. It is a commonly used ion exchange resin catalyst and the properties may be found elsewhere.20 Fresh catalyst was washed sequentially with distilled water, dilute hydrochloric acid, acetone and again with distilled water to remove any residual acid. It was further dried under vacuum at 70 °C for about 14-16 h in order to remove any moisture present in it. 2.2 Apparatus and analysis Experiments on kinetics were carried out in a kettle-type glass reactor of 250 ml capacity. About 60% of the reactor was filled with the reaction mixture and continuously stirred with the help of a mechanically agitated stirrer. Heating was provided by a constant temperature oil bath and the temperature of the reaction vessel was recorded with the help of calibrated temperature sensors, Pt-100. The reactor was equipped with a condenser, and chilled water was circulated to avoid 7 ACS Paragon Plus Environment


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losses, if any, due to evaporation. Samples were withdrawn after specific time intervals and immediately quenched, in order to cease any further reaction. The acid content in the sample was also measured by titrating it against standard alcoholic KOH solution. Analysis of the samples from RD experiments (Section 4) was performed using a Gas Chromatograph (GC) (Chemito 8610), equipped with the flame ionisation detector. The column used for analysis was 30 m long capillary column (BP-5; ID = 0.53 m) with nitrogen as a carrier gas. Injector and detector temperatures were maintained at 150 °C. The oven temperature (50 °C to 230 °C) was carefully programmed to achieve the best possible resolution. The oven temperature was increased at a rate of 10 °C /min till 90 °C, before being maintained at that temperature for 1 min. This was followed by reheating the oven at a rate of 30 °C/min until 230°C.

Iso-propyl alcohol was used as an external standard. The error in instrument and

sampling, together, was found to be ± 2%. Water in the samples cannot be detected by FID and hence, it was determined by material balance. The results obtained by GC were verified by independent titrations using alcoholic KOH solution. Elimination of mass transfer resistance Preliminary experiments were performed at different speeds of agitation (400-1000 rpm) to study the effect of external mass transfer resistance on the rate of reaction. Mass transfer resistance was found to be negligible at 1000 rpm, beyond which there is no further change in the rate. Hence, all the remaining runs were performed at 1000 rpm. These runs were performed over a range of mole ratio, catalyst loading and reaction temperature to generate the intrinsic kinetic data. The reproducibility of the kinetic experiments was within ±5% range. Figure 1 shows the effect of these parameters on the conversion vs. time plot.


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The effect of temperature, as depicted in Figure 1a, was studied over the range of 325-353 K and as expected, the reaction rate was found to be sensitive to the change in temperature. The effect of reactant molar ratio (acid: alcohol) in the feed was also studied over the range of 0.5-3. The fractional conversion of the limiting reactant increases if the other reactant is used in excess of stoichiometric requirement (Figure 1b). The effect of catalyst loading was found to be insignificant above 10% (w/w), when varied in the range of 5-15% (w/w), as shown in Figure 1c. The next section elaborates the development of the kinetics model and the estimation of the relevant parameters based on the data presented in this section.



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a)

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b)

Figure 1. a) Effect of temperature on conversion of acetic acid ( catalyst loading 5% (w/w), AA:MP ratio = 3:1 ) b) Effect of mole ratio of AA:MP on conversion of limiting reactant (catalyst loading 5% (w/w), temperature 353 K) c) Effect of catalyst loading on conversion of acetic acid (Temperature 353 K, mole ratio AA:MP = 3:1)

c) 2.3 Kinetic parameters A concentration based pseudo-homogeneous kinetic model was found to explain the data reasonably well. Equation 2 gives the expression for the overall rate of reaction. Mcat is the weight of the catalyst in the reaction mixture, n0 is initial number of moles of reactants in the 10 ACS Paragon Plus Environment

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mixture, Xa, Xb, Xc, Xd are the mole fractions of acetic acid, methoxy propanol, methoxy propyl acetate and water, respectively. It is a commonly used form of rate expression (see e.g. [21] for methyl acetate hydrolysis). The reaction is endothermic in nature. The rate equation is given by, {

– (

|

(2) (

)

|

)

Table 2 gives the values of the kinetic parameters estimated using MATLAB non-linear regression tool nlinfit. The activation energy for the forward reaction (66503 J/mol) is in close agreement with the literature value (61581 J/mol) reported recently.22 The equilibrium constant increases with temperature and the estimated heat of reaction is 41510 J/mol. The more details are available in the supporting information [S1]. Table 2. Kinetic parameters for Amberlyst-15 catalysed esterification of methoxy propanol with acetic acid

Parameters

95% confidence limit values

Ef (J/mol)

66503

Eb (J/mol)

24875

kfo(kmol/kg-sec)

3.7E +05

kbo (kmol/kg-sec)

0.393

MSE

2.4E-04



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3. Vapour liquid-liquid equilibrium and Residue curve map 3.1 Vapour liquid-liquid equilibrium Vapour liquid equilibrium data at atmospheric pressure for some of the binaries viz. MP-MPA, MPA-acetic acid, toluene-MPA, toluene-MP and MP-water, is not available in literature. The thermodynamic data for these binary pairs (Figure 2) is generated experimentally using a modified Othmer’s still apparatus. The details of the experimental setup and the operating procedure may be found elsewhere.23 The UNIQUAC interaction parameters were estimated using the generated experimental data by the new Britt-Luecke algorithm and Deming initialisation method in Aspen Plus (version 7.3). Maximum Likelihood is used as the objective function in the regression analysis. Table 3 depicts the value of UNIQUAC interaction parameters for ten binary mixtures involved in the study. Figure 2 shows the predicted T-xy plot, along with the observed experimental results. UNIFAC model was chosen for the remaining binaries. The binary mixture Water-MPA is heterogeneous in nature and the apparatus used, modified Othmer’s still, is proven to be accurate for homogeneous system. Hence all the five binaries, which are homogeneous in nature, were studied in this apparatus and UNIFAC values were used for the MPA-water system. As regards the MP-AA mixture, components MP and AA are the reactants and the reaction takes place, to some extent, even in the absence of catalyst at boiling temperature. Hence, generating true phase equilibrium data for this binary mixture in VLE apparatus is difficult. We therefore prefer to use UNIFAC model and parameters for MPAA binary as well. More information on UNIQUAC molecular volume and surface parameters is given in Table S1 of the supporting information.


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a)

b)

c)

d)

Figure 2. Comparison of the T-xy plots generated through the UNIQUAC Equation of State (EoS) and experimental data at 1 atm for the binaries a) MP-MPA b) MPA-acetic acid c) toluene-MPA d) toluene-MP and e) MP-water. e) .13 ACS Paragon Plus Environment


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Table 3. UNIQUAC binary interaction parameters for different component pairs (Temperature °C) Comp i

Compj

Aij

Aji

Bij

Bji

Source

Water

Acetic acid

4.20

0.75

196.90

-615.26

Aspen

Acetic acid

Toluene

0.298

-1.203

0

0

Aspen

Toluene

Water

0

0

-950.6

-350.21

Aspen

MP

Water

25.54

-16.59

-10000

6402.66

Exp (UNIQ)

Toluene

MP

6.526

0.84

-2268.06

-880.49

Exp(UNIQ)

MPA

Acetic acid

0

0

-96.87

-12.45

Exp (UNIQ)

Toluene

MPA

8.36

-13.58

-3108.88

5049.42

Exp (UNIQ)

MP

MPA

-1.04

-1.22

514.44

403.60

Exp (UNIQ)

Water

MPA

0

0

-325.42

79.08

UNIFAC

Acetic acid

MP

0

0

117.45

-47.12

UNIFAC

3.2 Residue curve maps Before proceeding to the experimentation in RD, it is necessary to understand the feasibility of RD with the help of Residue Curve Maps (RCM) for the ternary systems involved in MPA synthesis. These topological diagrams help us to identify the most feasible alternative for the separation of given mixture, involving azeotropes. The system of interest has several azeotropes present in it. The corresponding data has already been listed in Table 1. Figure 3 exhibits the RCMs of a few important ternary systems involved in the reaction mixture. It also shows the RCM for the n-butanol-n-butyl acetate-water mixture for comparison. It is revealed from Figure


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3a that, unlike the n-butanol esterification system (Figure 3b), there are three azeotropes; MPwater, MPA-water and MPA-MP-water, present very close to each other. The presence of these azeotropes in the RD column makes the separation task difficult and as a result, distillate may contain some MP, along with the product MPA. Furthermore, the alcohol-water azeotrope in the case of MP is the least volatile and homogeneous in nature. The n-butanol-water azeotrope, however, is heterogeneous and intermediate boiling as shown in Figure 3b. Therefore, in all probability, the overhead mixture in the case of MPA synthesis is homogeneous in nature. It is therefore necessary that we add an external mass separating agent that would form an azeotrope with a boiling point much less than all the other three azeotropes. The RCM in Figure 3c shows the presence of only two azeotropes MPA-water and toluene-water. There is a wide gap between the azeotropic temperature of toluene–water azeotrope and that of other azeotropes. This toluenewater azeotrope thus has the least boiling point among all the azeotropes present in system and hence facilitates separation to yield almost pure water in the aqueous phase of the overhead product.



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b)

a)

c)

Figure 3.

Residue curve maps for ternary mixtures a) MPA-MP-water b) butanol-butyl acetate

-water c) MPA-toluene-water 16 ACS Paragon Plus Environment

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4. Continuous RD 4.1 Experimental Setup Figure 4 shows schematic of the continuous RD column with decanter used in this work. The glass column is 3.18 m long with 50 mm internal diameter. It consists of rectifying and stripping sections each 0.9 m long, and the middle reactive section of length 1.38 m. The intermediate sections are interconnected using Teflon gasket assembly equipped with ports for thermocouple and sample collection. Stripping and rectifying sections are packed with HYFLUX packing from M/s. Evergreen India Ltd.24, which is a commercially available low-pressure-drop, structured packing made of stainless steel 304. The mesh structure formed due to knitting of steel wire ensures proper flow channel for liquid and vapour contact thereby offering high mass transfer rates and separation efficiency with low flow resistance. The NTSM (Number of theoretical stages per meter) value for this packing is 8 m-1, which corresponds to 14 non-reactive stages (7 each in rectifying and stripping sections) and the liquid hold up is 3-8% w/w. Reactive section (1.38m) consists of KATAPAK packing from Sulzer, with NTSM value of 3 m1

, loaded with 260 gm of pre-treated Amberlyst-15 catalyst. Thus the total number of theoretical

stages present in the column including reboiler is 19. Multiple temperature sensors, Pt-100, and sample ports are provided throughout the column at appropriate locations. The glass column is covered with a thick (2 cm) insulation of glass wool and then covered with a layer of asbestos rope (thickness: 1 cm) to avoid heat losses to the surrounding. The reboiler (capacity: 2 litres) is electrically heated and wattmeter is used to measure the power that can be regulated as desired. The exposed part of the reboiler is also externally covered with glass wool and asbestos rope to avoid heat losses to the surrounding. The feed is introduced to the column with the help of a



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calibrated peristaltic pump, and the distillate is collected in a vertical condenser at the top of the column with a Dean and Stark assembly as shown in Fig. 4. The organic and aqueous phases are well separated in this assembly. The aqueous phase formed is collected continuously, and the organic phase is recycled back to the column in the form of reflux. The decanter temperature was in the range 40 – 60 °C, depending on the extent of reboiler duty. The column was operated at atmospheric pressure. Start-up is the most crucial step in RD experiments, and appropriate initial composition of the mixture in the column leads to attainment of steady state in a shorter time. The experimental run was initiated by charging the reboiler with a mixture of all the four components. Once the reflux starts due to condensation of vapours rising in the column, feed comprising of acetic acid and MP in a stoichiometric ratio (room temperature, atmospheric pressure) was introduced at a fixed rate in the middle of the reactive section. Once the steady state is achieved, it was maintained for 2-3 hours and samples were withdrawn every half an hour to confirm steady state and were analysed with the help of a GC. Attainment of steady state was indicated by negligible variation in temperature and compositions throughout the column with respect to time.


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Figure 4. Continuous RD set-up. 4.2 Experiments with and without entrainer The feasibility of RD for MPA synthesis was first checked without any entrainer. Reboiler duty was fixed at 500 W. The feed with 1:1 molar ratio of reactants, is introduced at room temperature and atmospheric pressure at the middle of the reactive section (position 6), at a rate of 8 ml/min. 19 ACS Paragon Plus Environment


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The presence of homogeneous minimum boiling azeotrope (MP-water at 97.31 °C) resulted in a homogenous liquid condensate at the top, which contains significant amount of MP, thereby resulting in a relatively less conversion (70-80%). The bottom product purity was also limited due to the presence of unreacted reactants. Furthermore, the presence of one more ternary azeotrope between MPA-MP-water (98.16 °C) also led to the presence of product ester in the distillate. Since the boiling temperatures of all the three azeotropes are close to each other, the best operating window for this reaction is extremely narrow. In case of n-butanol, the ternary heterogeneous azeotrope, being least volatile, appears at the top and facilitates separation of water.11,25 Such a situation does not exist in the case of MP and hence, there is a need to explore the option of entrainer enhanced RD process. As mentioned before, the entrainer forms a minimum boiling azeotrope with water and facilitates separation. An experimental run was therefore carried out with all the operating parameters identical to the above specified conditions, except for the addition of toluene with the feed. As expected, the toluene-water heterogeneous minimum boiling azeotrope (84.29 °C) was found to be dominant amongst all the present azeotropes. This experiment showed encouraging results in terms of bottom purity and conversion. The conversion increased by 12% (from 79% without entrainer to 91% with entrainer) and loss of reactant through the bottom stream and loss of ester through the distillate reduced to almost negligible levels. Figure 5 compares profiles obtained with and without toluene. The shaded portion in the figure is the reactive zone filled with catalytic packing. Figure 5 clearly shows a drastic reduction in the loss of the reactants - acetic acid and MP, in the bottom product when toluene is used as entrainer. Figure 5c reveals a substantial increase in MPA concentration (approx. 60% by mole) in the bottom product with, negligible loss through the aqueous phase of the decanter in the presence of toluene. The presence of 20 ACS Paragon Plus Environment

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toluene as an entrainer thus had a positive effect on conversion and separation. However, some acetic acid was still found to escape through the aqueous phase from the decanter, in the presence of toluene. The loss of reactants can be further minimised to achieve close to complete conversion by varying the design and operating parameters like feed position of reactants, feed position of entrainer and reboiler duty. This is elaborated in the subsequent sections.

a)

b)

c)

Figure 5. Comparison of experimental profiles of the components along the length of the column with and without toluene a) Mole fraction Acetic acid b) Mole fraction MP c) Mole fraction MPA.



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It may be noted that in most of the runs, we observed small amounts of a high boiling side product propanediol diacetate (< 3 mol%) in the reboiler. Interestingly, this side product was not detected in the batch kinetics experiments under closed conditions without water removal. It is conjectured that the methoxy group is released from the MPA molecule in the form of methanol, which upon reaction with acetic acid, forms volatile methyl acetate that escapes through condenser. This hypothesis needs further substantiation by careful experiments allowing one to detect components at low levels. 4.3 Parametric study The experimental results obtained are compared with the predictions by the RADFRAC model from Aspen Plus, under identical conditions. The RADFRAC model uses equilibrium stage assumption with kinetically controlled reaction. Though the simulator does not consider column hardware effects, it gives an insight into the effect of several operating and design parameters and helps us in arriving at the suitable set of conditions to obtain best possible performance. The output results from the simulation are the steady state values of conversion, compositions of the product streams, flow rates both for distillate and final bottom product, composition and temperature profiles along the column. Table 4 lists the input specifications used in base case simulations. In the following section, we explain the effect of different parameters.


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Table 4. Specification used in Aspen Plus simulation

Parameter

Specification

Property method

UNIQ – HOC

Interaction parameters values

All experimental and UNIFAC (Table 3)

Reaction kinetics

All experimental and regressed vales (Table 2)

RADFRAC No of stages

19

Reboiler duty

300W to 500 W

Reactive Stages

9-12

Catalyst loading

65 gms per stage

Reboiler type

Kettle

Valid phase

Vapor- liquid- liquid

Liquid split consideration

16-19 stages

Feed temperature and pressure

30 °C and 1 atm.

Decanter Operating temperature and pressure

30 °C and 1 atm.

Key component for second liquid

0.5

phase Phase split determination by

Equating fugacities of toluene and water in the two phases



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Reboiler duty The effect of reboiler duty on the column performance was studied over the range 300-500 W. An increase in reboiler duty shifts the reaction in forward direction by virtue of increased water removal, thereby increasing the overall conversion. However, at a higher reboiler duty (~500 W), the concentration of high boiling methoxy propyl acetate increases in the reactive zone, as a result, the temperature in this zone approaches close to thermal stability of the catalyst (120 °C), which is not desired. Figure 6 exhibits the experimental and simulation composition profiles, along with a comparison of conversions obtained by experiments and simulation at different reboiler duties, under otherwise identical column conditions. Further analysis of the aqueous phase from the decanter showed that the loss of reactants is significant at a reboiler duty of 500 W, leading to a drop in conversion. The mole fractions of AA and MP in aqueous phase were observed to be 0.019 and 0.082, respectively, for a reboiler duty of 500 W. It was also observed that the proportion of side product, propane diol diacetate, in the bottoms also increases substantially at higher reboiler duty (7% at 500W). A reasonably high conversion was realised in the case of reboiler duties of 300 W and 400 W (Figure 6d). Hence, further parametric studies were performed with a fixed reboiler duty of 400 W and reactant mole ratio of 1:1 (Acetic acid: MP).


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a)

b)

c)

d)

Figure 6. Composition profiles of components along the length of column and conversion values for different reboiler duties a) acetic acid b) MP c) MPA d) Conversion of acetic acid by experiments and simulation.(Operating conditions: Pressure: 1 atm; Reboiler duty = 300W-500 W; feed rate = 8 ml/min; AA: MP mole ratio =1; feed temperature = 30 °C)



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Feed Position Since there is not much difference in the volatilities of MP and acetic acid, both the reactants were mixed along with toluene and fed at the bottom (stage 8), middle (stage 11) and top of the reactive section (stage 12) in three independent experiments under otherwise similar conditions. The experimental results show the conversion approaching 100% in all the three cases. Figure 7a exhibits the comparison of experimental and simulation conversions at the three distinct feed positions. Among these, top feed position showed relatively higher conversion. The catalyst can also be maintained below the thermal stability limit without the need of an external toluene make-up. Figure 7b compares the experimental and simulated temperature profiles for this case. We observed small amount of toluene leaving in the bottom product when the feed was introduced at the top of the RD column. Hence, this configuration is not recommended. On the other hand, when the feed is introduced at the bottom of the reactive section, we observed temperature rise in certain parts of the reactive zone. These problems of toluene escape from bottom and local temperature rise were not encountered when feed was introduced in the middle of the reactive zone. Hence, it is recommended that the feed be introduced at the middle of the reactive section.

26

ACS Paragon Plus Environment

Page 27 of 42

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b)

a)

Figure 7. Conversion and temperature profiles at different feed positions and fix reboiler duty 400 W a) Conversion values experimental and simulated at three feeding positions (stage- 8, stage-11, stage-12) b) Comparison of experimental and simulated temperature profiles for the feed at stage 12. Furthermore, the effect of adding entrainer at different locations along the column was also studied. In these experiments, the feed, comprising of both the reactants was introduced in the middle of the reactive section and entrainer feed location was varied from the bottom to the top of the reactive section. As expected, since only a small quantity of entrainer is added as make-up stream, the results indicated conversion to be insensitive to the entrainer positions. Noticeable 27



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reduction in the time that is required to achieve steady state (nearly 50%), was observed for the case when entrainer is introduced at the top position as compared to other entrainer positions. No significant change in bottom purity was observed. Detailed dynamic studies may be performed to explain this observation. The comparison of simulation and experimental results for one representative run is shown in the Figure 8. The agreement is reasonably good. The predicted concentrations of water are however, slightly lower than the experimental. Similar observations were also made in some of our earlier studies [e.g. 11]. In the upper section of the column (stage 17 and above), phase split was observed and exact overall composition was difficult to determine and hence not shown in Figure 8. Nevertheless, the compositions of MPA, AA and MP in the aqueous phase (experimental) are mentioned in the caption of Figure 8.

28


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a)

d)

b)

e)

c)

f)

Figure 8. Comparison of experimental and simulation results obtained for concentrations of a) acetic acid b) methoxy propanol c) methoxy propyl acetate d) toluene e) water and f) column temperature. (Compostion of aquesous phase in the decanter: AA = 0.014, MP, MPA~0) 29



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Experimentally validated simulator is used for parametric study over a wide range of parameters. This helps in achieving the conditions corresponding to quantitative conversion with the minimum energy consumption. Reflux Split Ratio Reflux split Ratio is defined as the ratio of reflux given to the top stage to the total reflux sent to the column. It means when all the reflux is sent to the top stage of the column then the split ratio is unity. Split ratio less than unity means part of the reflux is sent at some other location in the column. Such a reflux policy has been proved to be beneficial in case of some esterification reactions involving liquid-liquid split in the overhead stream.18 Part of the organic phase, if fed below the reactive zone, removes water efficiently from reactive section and enhances conversion especially in the case of high boiling alcohols/esters. In case of esterification of 2ethyl-hexanol, the conversion was found to be higher at low reflux split ratio at relatively lower reboiler duties (70% at a reflux split ratio of 0.2 against 50% at a reflux split ration of >0.6). This prompted us to study the effect of split ratio in the present case. The effect of reflux split ratio on conversion was studied at different reboiler duties. The reboiler duty was varied from 3 kW to 5 kW at identical column conditions. It is seen from Figure 9a that for all the reboiler duties, conversion increases with increasing reflux split ratio. Unlike the case of ethyl hexyl acetate, there is no specific advantage of splitting the reflux. This is mainly because toluene, sent to the bottom of the reactive zone, removes not only acetic acid but also methoxy propanol (MP) from the reactive zone, thereby adversely affecting the rate of reaction. In case of 2-ethyl-hexanol, the volatility of alcohol is much less than that of toluene and one realizes positive effect of split under certain conditions.

30


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a)

b)

Figure 9. a) Effect of split ratio on conversion at different reboiler duties b) Loss of acetic acid and MP in aqueous phase of decanter at different reflux split ratio studied for a fixed reboiler duty of 3 kW. Column conditions: number of theoretical stages = 35; feed rate of acetic acid and MP = 0.050 kmol/hr each; number of reactive stages = 7. Number of reactive stages and catalyst hold up per stage Effect of catalyst loading on conversion is studied by varying the amount of catalyst on a single stage (stage 14). As the amount of catalyst on a single stage was increased from 1 to 10 kg (Figure 10a), the conversion was found to increase significantly. However, addition of catalyst beyond this limit (10 kg) led to only a small increment in conversion value. Alternatively, the number of reactive stages was also increased while keeping a fixed catalyst loading (60 kg) on each stage. The conversion increased with an increase in the number of reactive stages, as depicted in Figure 10b. Simulation was also performed with a fixed catalyst loading of 60 kg, equally distributed over the reactive stages. Conversion was found to increase with increase in number of reactive stages from 1 to 12. However, beyond this point an increase in the number of 31



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reactive stages did not show a significant increase in the conversion. Thus the total catalyst requirement reduces if it is distributed by extending the length of the reactive zone. This indicates a synergism arising from interplay of reaction and distillation.

a)

b)

Figure 10. a) Effect of single stage catalyst loading on conversion for a feed flow rate of acetic acid and MP = 0.050 kmol/hr each; reboiler duty = 5 kW; number of theoretical stages = 35 and b) Effect of number of reactive stages for a fixed catalyst loading of i) stage ii)

60 kg catalyst distributed equally on all reactive stages iii)

60 kg catalyst on each 4 kg catalyst distributed

equally on all reactive stages. Effect of rectifying and stripping sections The effect of rectifying and stripping sections was studied under otherwise similar conditions, i.e. feed flow rate of acetic acid and MP = 0.05 kmol/hr each; total catalyst = 60 kg; number of reactive stages = 25, Reboiler duty of 5 kW.

32


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Rectifying section: On reducing the height of rectifying section, a higher proportion of both the reactants were observed in the aqueous phase of the decanter. This was due to a slight decrease in the overall reactant conversion. Although the drop in conversion is small, it decreases the purity of aqueous phase. The number of stages in the rectifying section has to be decided such that the loss of reactants is minimal and hence, the purity of aqueous phase is maintained. Hence, approximately 11 rectifying stages are sufficient for better purity in aqueous phase, as observed in Table 5. Table 5. Effect of number of rectifying stages on conversion and purity Rectifying section stages

Conversion (%) Purity of aqueous phase (%)

11

99

98.6

7

99

98.4

3

98.9

98.1

0

98.8

97.9

Stripping section The number of stages in the stripping section was also altered to observe its effect on conversion. As the numbers of stages in the stripping section were decreased, the conversion was observed to increase. A decrease in the number of stages meant that the reactive zone was now closer to the reboiler (at high temperature). As a result, the rate of reaction increased, leading to an increase in conversion. While optimising number of stages for stripping section, one should examine the temperature profile in the reactive section, as reducing stripping section height may also have an adverse effect on life of the catalyst in the reactive zone that get exposed to high temperature 33



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environment. Though the maximum conversion was observed for 5 stripping stages (Table 6), it is not recommended since the temperature profile (Figure 11) indicates high temperature zone at the bottom of the reactive section crossing the thermal stability limit of the catalyst (125 °C).

Figure 11. Temperature profile of reactive section for different number of stripping stages.

34


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Table 6. Effect of number of stripping stages on conversion and purity Stripping section stages

Conversion (%)

Purity of bottom product (%)

17

99

99

14

99.1

99

9

99.2

99.1

7

99.2

99.1

5

99.5

99.2

5. Recommended configuration The conventional reactor in which the conversion is limited by equilibrium may suffer from many downstream separation steps for the recovery and recycle of components. This is because of the homogeneous azeotropes involved in the reacting system. Such a process is not reported in literature. The commonly practised batch process for the production of alkyl esters with homogeneous catalyst is also associated with large reflux and higher energy consumption (see supporting information [S5] for the flowsheet and more details). Hence, based on the parametric studies, we recommend the best configuration that offers high conversion, minimum energy consumption and desired purity. It is shown in Figure 12. It has 43 stages altogether, with 28 reactive stages starting from stage 8 to stage 34, 67 kg of the total catalyst hold up in the reactive section and requires a reboiler duty of 6.8 kW. Feed (1:1 mole ratio of AA and MP) is introduced on stage 22, at a flow rate of 100 mol/hr, at temperature of 30 °C and atmospheric pressure along with very small toluene make-up (approximately, 0.0002 kmol/hr) to the column.

35



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Figure 12. Schematic diagram for proposed toluene based continuous RD process. Column conditions: Number of stages = 43; feed flow rate of acetic acid and MP = 0.050 kmol/hr each; Number of reactive stages, 8-34; Reboiler duty = 6.8 kW; Conversion= 99.3%.

6. Conclusion Entrainer based reactive distillation offers enhanced conversion and purity over the conventional RD process for the synthesis of methoxy propyl acetate. Unlike RD for other C4/C5 alcohols, the presence of close boiling azeotropes in the present case, does not allow us to design a column without entrainer, for high reactant conversion and product purity. Remarkable increase in conversion and purity were realized when entrainer-toluene was introduced in the process. Furthermore, parametric studies revealed that the quantitative conversion could be achieved 36


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without compromising on the thermal stability of the catalyst. The proposed alternative presents a more compact and cost-effective process for the production of MPA.

Nomenclature Mcat - weight of catalyst in kg. n0 - initial number of moles Xa - mole fraction of Acetic acid Xb - mole fraction of Methoxy propanol Xc - mole fraction of Methoxy propyl acetate Xd - mole fraction of Water kf - forward reaction rate constant kmol/(kg-sec) kb - backward reaction rate constants kmol/(kg-sec) kfo, - Arrhenius pre-exponential factor for forward rate constant, kmol/(kg-sec) kbo - Arrhenius pre-exponential factor for back-ward rate constant, kmol/(kg-sec) Ef - activation energy of forward reaction, J/mol Eb - activation energy of backward reaction, J/mol Abbreviations: AZEO – Azeotrope

37



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MPA - Methoxy Propyl Acetate MP - Methoxy Propanol RD - Reactive Distillation VLE - Vapour Liquid Equilibrium AA - Acetic Acid GC - Gas Chromatography RCM - Residue Curve Map NTSM - Number of Theoretical Stages per Metre Acknowledgement Authors acknowledge assistance and inputs provided by Dr. Ishan Sharma, Chemical Engineering Department IIT Bombay, India in simulation of the current work. Supporting Information Van’t Hoff plot for estimating heat of reaction; comparative experimental and simulation composition profiles for different parameter conditions; values of molecular volume and surface for components using UNIQUAC model; mass balance for experimental runs; conceptual conventional flowsheet for MPA synthesis. The supporting information is available free of charge via the Internet at http://pubs.acs.org

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For Table of contents/graphical abstract only

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Reactive Distillation for the Production of Methoxy Propyl Acetate

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