Design, Optimization and Retrofit of the Formic Acid Process II

Subscriber access provided by UNIV OF LOUISIANA

Process Systems Engineering

Design, Optimization and Retrofit of the Formic Acid Process II: Reactive Distillation and Reactive Dividing-Wall Column Retrofits Sergio Jose Rodrigues da Cunha, Gade Pandu Rangaiah, and Kus Hidajat Ind. Eng. Chem. Res., Just Accepted Manuscript • Publication Date (Web): 05 Oct 2018 Downloaded from http://pubs.acs.org on October 5, 2018

Just Accepted “Just Accepted” manuscripts have been peer-reviewed and accepted for publication. They are posted online prior to technical editing, formatting for publication and author proofing. The American Chemical Society provides “Just Accepted” as a service to the research community to expedite the dissemination of scientific material as soon as possible after acceptance. “Just Accepted” manuscripts appear in full in PDF format accompanied by an HTML abstract. “Just Accepted” manuscripts have been fully peer reviewed, but should not be considered the official version of record. They are citable by the Digital Object Identifier (DOI®). “Just Accepted” is an optional service offered to authors. Therefore, the “Just Accepted” Web site may not include all articles that will be published in the journal. After a manuscript is technically edited and formatted, it will be removed from the “Just Accepted” Web site and published as an ASAP article. Note that technical editing may introduce minor changes to the manuscript text and/or graphics which could affect content, and all legal disclaimers and ethical guidelines that apply to the journal pertain. ACS cannot be held responsible for errors or consequences arising from the use of information contained in these “Just Accepted” manuscripts.

is published by the American Chemical Society. 1155 Sixteenth Street N.W., Washington, DC 20036 Published by American Chemical Society. Copyright © American Chemical Society. However, no copyright claim is made to original U.S. Government works, or works produced by employees of any Commonwealth realm Crown government in the course of their duties.

Page 1 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Industrial & Engineering Chemistry Research

Design, Optimization and Retrofit of the Formic Acid Process II: Reactive Distillation and Reactive Dividing-Wall Column Retrofits Sergio da Cunhaa, G. P. Rangaiaha*, Kus Hidajata a

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

Singapore 117585 *

Corresponding author. Email address: [email protected]

Abstract Formic acid (FA) is an important chemical with many applications in the industry. Nevertheless, its manufacturing process has not received much attention in the open literature. In Part I of this work, FA process was simulated and optimized for the production of 98 wt% FA. Total annual cost (TAC) of the optimal base case process was 0.686 USD/kg. This process was then retrofitted via integration of two distillation columns into a dividing-wall column (DWC), but such retrofit was not profitable. Here, in Part II, FA process retrofit via reactive distillation (RD) and reactive dividing-wall column (RDWC) are investigated in detail. Results show that RD retrofit is not attractive as it increases TAC of the process. However, retrofit via RDWC is profitable, resulting in thermal energy savings of 20.4% and TAC savings of 0.020 to 0.042 USD/kg (by 3 to 6%). Keywords: Formic Acid; Process Retrofitting, Process Revamping; Reactive Distillation; Reactive DividingWall Column; Multi-Objective Optimization

1. Introduction Formic acid (FA) is an important acid with various uses in the industry. Its main application is in silage, to prevent growth of undesirable bacteria during fermentation of forage crops. FA is also used in leather manufacturing, to ensure acidic conditions prior to tanning. Furthermore, this acid can be used in pharmaceutical and food chemicals synthesis, in the textile industry as pHregulating agent, and in many other manufacturing processes.1 Also, recent studies suggest FA is a valuable fuel for fuel cells.2,3

1 ACS Paragon Plus Environment

Industrial & Engineering Chemistry Research 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Despite its importance, FA process design has not received much attention in the open literature. Hietala et al.1 provides an overview of FA applications, its chemical and physical properties, and the main manufacture routes. Methyl formate (MF) hydrolysis is the most common industrial practice for FA production, corresponding to 81% of the global production capacity, and three manufacture routes based on MF hydrolysis seem to dominate the market.1 In all of them, MF is first formed via methanol (methyl alcohol, MA) carbonylation. The so-called KemiraLeonard process uses two reactors operating at different conditions for MF hydrolysis. The preliminary reactor produces small amount of FA to catalyze the hydrolysis in the main reactor. Excess MF (2:1 to 4:1) is used in the main reactor, which operates at 90°C to 140°C.4 Even at high temperatures, the reaction mixture remains in the liquid phase due to the reactor operating pressure (5 to 18 atm). Excess MF is recovered and recycled via distillation,4 and aqueous FA is purified via pressure-swing distillation.1 In the USSR process, the hydrolysis reaction is conducted in a distillation column at atmospheric pressure and packed with an ion-exchange resin in the upper zone. The FA formed in this zone acts as catalyst for the hydrolysis in the lower part of the column. Both the upper and the lower zones are equipped with a thermostat jacket; the former is kept at 60-100°C, whereas the latter is kept at 100-107°C.5 Excess water is fed to the column (up to 3:1), and a mixture of water and FA with small amount of MA is obtained in the column bottoms stream. MA and unreacted MF is taken as distillate.5 FA in the column bottoms stream is recovered and purified up to 98 wt.% using a sequence of distillation columns.1 The last manufacture route described in Hietala et al.1 is the BASF process. In this process, MF hydrolysis happens with excess water (about 5:1). MF and MA are subsequently separated from aqueous FA using a single distillation column with side stream product. Aqueous FA is obtained in the column bottoms stream, and it is separated using liquid-liquid extraction followed by distillation.6 The extractant is preferentially a secondary amide, and 1 to 12 theoretical stages may be required in the liquid-liquid extraction column. Volume ratio of extractant to FA used in this column is from 2:1 to 5:1.6 Other manufacture routes for FA have been developed and patented recently. Luxi Chemicals patented a process for the production of 100,000 tons of FA per year.7 In this process, one preliminary and one main reactor are used for MF hydrolysis, similar to the Kemira-Leonard 2 ACS Paragon Plus Environment

Page 2 of 41

Page 3 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


process. The subsequent separation steps are not clearly described in the patent, though it seems FA is separated in a series of distillation columns. Another recent patent by Huang et al.8 discloses the production of FA via reactive distillation (RD) followed by distillation. The main difference between this process and the USSR process is that, in the latter, MF and MA are taken in the distillate of the RD column, whereas, in the former, only a very small amount of CO (gas purge) leaves as the distillate. Both MF and MA are taken in the RD bottoms stream, together with water and FA. Although detailed description of the process is given Huang et al.8, no experimental data validating the invention is reported. Finally, Kemira Chemicals patented a process in which the MF hydrolysis is conducted in a chromatographic reactor packed with ion-exchange resin.9 In this invention, not only the catalytic properties of the resin are used, but also its ability to separate different components of the reaction mixture. Some academic publications10–13 give more details on certain aspects of these manufacture routes, such as process development and simulation setup. Wang14 studied and reported kinetics of the MF hydrolysis using an ion-exchange resin as catalyst. Papers by Jogunola et al.15–17 provided kinetics for FA production via the auto-catalyzed MF hydrolysis reaction. Sahin et al.18 investigated FA separation from water via reactive extraction. Recently, we developed the complete FA process based on the BASF process, simulated it in Aspen Plus V9.0 and later optimized for two objectives: minimize fixed capital investment (FCI) and utility cost.19 This process design, referred to as the base case FA process, is illustrated in Figure 1. Methanol (methyl alcohol, MA) and carbon monoxide (CO) react in CSTR1, to form methyl formate (MF). Unreacted reactants and reaction products from flash (F) bottoms are sent to a distillation column (DC1). MA is recovered at the bottoms of DC1 and recycled back to CSTR1. MF and some MA are sent to CSTR2, where the former will react with water to form FA and MA. The outlet from CSTR2 is sent to column DC2, where MA and unreacted MF are separated from FA and unreacted water. The components in DC2 distillate are separated in column DC3, where the distillate (MF) is recycled to CSTR2 and the bottoms (MA) is recycled to CSTR1. Heavy components in DC2 bottoms are sent to an extraction column (EXT), wherein diisopropyl ether (DIPE) is used as the extractant/solvent. The organic outlet (FA + DIPE) from EXT is sent to another distillation column (DC4), where FA 98 wt% leaves at the bottoms, and DIPE with traces of water leaves at the top. The distillate of DC4 is then cooled and sent to a decanter, which separates DIPE from water. DIPE is recycled back to EXT, and the water streams from both 3 ACS Paragon Plus Environment


Page 4 of 41

decanter and EXT are mixed and sent to another distillation column (DC5). DC5 separates the small quantity of DIPE present in the aqueous mixture as the distillate for recycling to EXT, and the bottoms (water) is recycled to CSTR2. Economics for the base case process were evaluated using total annual cost (TAC), defined in Towler and Sinnot20: 𝑇𝐴𝐶 = 𝐶𝑂𝑀 + ACCR × FCI/PR

(1)

where COM refers to cost of manufacture without depreciation per unit of FA produced, ACCR is the annual capital charge ratio, FCI is the fixed capital investment for the entire process and PR is the annual production rate (=27,100,000 kg/y). Eq. 1 gives TAC per kg of FA product. COM and FCI are evaluated following the methodology in Turton et al.21 COM, which includes all manufacturing costs, is the sum of direct manufacturing costs, fixed manufacturing costs and general manufacturing expenses; it is given by:21 𝐶𝑂𝑀 =

0.18𝐹𝐶𝐼 + 2.73𝐶𝑂𝐿 + 1.23(𝐶𝑈𝑇 + 𝐶𝑊𝑇 + 𝐶𝑅𝑀) 𝑃𝑅

(2)

Here, COL is the cost of operating labor, CUT is the cost of utility, CWT is the cost of waste treatment, and CRM is the cost of raw material. All these are in USD/kg of FA.

Figure 1 – Base case process19


Page 5 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Eqs. 1 and 2 evaluate economics for a new project. To evaluate the effect of retrofit on the TAC of an existing plant, we assume that the plant retrofit affects only utility cost, raw material cost and production capacity. Other operating costs (e.g. cost of labor, taxes, maintenance) are assumed to remain constant. Therefore, TAC after retrofit is calculated using the following equation: 𝑇𝐴𝐶𝑟𝑡 =

PR𝑜𝑙𝑑 PR𝑟𝑡

× (𝑇𝐴𝐶𝑜𝑙𝑑 ― 𝐶𝑈𝑇,𝑜𝑙𝑑 ― 𝐶𝑅𝑀,𝑜𝑙𝑑) + 𝐶𝑈𝑇,𝑟𝑡 + 𝐶𝑅𝑀,𝑟𝑡 + FCI𝑟𝑡 ×

ACCR PR𝑟𝑡

(3)

where subscript old and rt of TAC refer to original and retrofitted processes, respectively. FCIrt refers to total module cost of new equipment added to an existing plant.21 Note that Eq. 3 accounts for small difference in the production rate before and after retrofitting. If this difference is neglected (i.e., PRold = PRrt), TACrt is simply TACold plus annualized FCI for retrofitting minus savings in utilities and raw material costs. Moreover, da Cunha et al.19 proposed three different retrofit solutions to improve economics of the optimal base case process. These solutions were obtained by: (i) replacing CSTR2 and DC2 by a reactive-distillation (RD) column; (ii) replacing DC2 and DC3 by a dividingwall column (DWC); and (iii) replacing CSTR2, DC2 and DC3 by a reactive dividing-wall column (RDWC). Among these, DWC retrofit has been simulated and optimized for minimizing FCI and utility cost. TAC for the FA base case process after optimization was 0.686 USD/kg.19 After DWC retrofit, TAC increased slightly to 0.701USD/kg, which shows that this retrofit is not profitable.19 In the present paper, we develop and evaluate economics of the two remaining retrofit solutions: RD and RDWC. RD is a relatively new technology that consists of a single column performing reaction and separation simultaneously. This technique improves yield of equilibrium-limited reactions by separating products from reactants. The reactants have higher concentration in the reactive zone of the column, which shifts the equilibrium towards the products. RD applications in the industry include synthesis of methyl tert-butyl ether (MTBE), methyl acetate and acetic acid.22 Among these, methyl acetate RD process developed by Eastman Kodak is particularly intensive and profitable. It replaces the traditional process with two reactors and eight distillation columns by a new process with only one RD column. Methanol (b.p. 65°C) is fed at the bottom of the reactive section, and acetic acid (b.p. 118°C) and catalyst (sulfuric acid) are fed at the top. Water (b.p. 100°C) leaves at the bottoms, whereas methyl acetate (b.p. 57.1°C) is taken as the distillate.22 5 ACS Paragon Plus Environment


Despite the benefits described earlier, there are some difficulties in RD implementation. One of them is when the reaction requires a long residence time. In this case, large column size/tray holdups are needed;23 this increases capital investment for the column. Process condition mismatch is another common problem in RD; optimal temperature and pressure for distillation may be different from those for reaction.23 A third drawback of RD is the occurrence of multiple steady-states.23 Internals in an RD column will be different from those in conventional distillation columns. For example, homogeneous RD processes require larger liquid hold-ups in order to achieve high reaction rates. Since tray columns have higher liquid hold-up than packed column24, the former is preferred for such applications. Weir height in homogeneous RD tray columns vary between 80 mm and 100 mm, whereas this height is 40 to 80 mm in conventional distillation.25 For heterogeneous catalyzed processes, tray or packed columns can be used. In tray columns, the catalyst can be loaded on the trays, inside the downcomers, or in packed layers between tray sections.23 In packed columns, catalyst particles are usually enveloped within wire gauzes, and these envelopes are packed together in the column to perform separation and reaction simultaneously. In such structures, catalyst loading corresponds to 20-25% of the bed volume.24 RDWC is also a cutting-edge technology for chemical processes. It is another integrated unit performing both reaction and separation. The main difference between RDWC and RD is that the former employs a dividing-wall column, whereas the latter makes use of a simple column without internal walls. Interest in RDWC has been increasing since 2014, as shown in Figure 2. It seems RDWC is still a concept under development, and it may take a few more years for RDWC to become accepted and used in the industry. Indeed, up to 2011, there were no reports on chemical facilities employing this technology.26 Drawbacks in RDWC implementation are expected to be those in RD and DWC implementation. These are: larger temperature difference between condenser and reboiler,27 requirement of hot utility at a higher temperature,27 increase in column height,27 controllability issues,26,28 development of vapor and liquid distribution systems inside the column,28 large column size,23 process condition mismatch for distillation and reaction,23 and multiple steady states.23 However, simulation results show that most of these issues do not apply for the FA process.


Page 6 of 41

Page 7 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Figure 2 – No. of journal publications (until 2017) in Scopus database, having the terms reactive dividing (divided) wall column, dividing (divided) wall reactive distillation column, RDWC or DWRDC in title, abstract or keywords The novelty and contribution of the present work are systematic simulation and optimization of RD and RDWC retrofit options, for the first time, for FA process using distillation and extraction for the downstream separation. There is no publication in the open literature on the retrofit of FA process with RD or RDWC. Hence, this paper is novel, significant and useful for retrofitting FA plants in operation in different parts of the world. The rest of this paper is organized as follows. Section 2 describes the RD retrofit development, simulation and optimization for two conflicting objectives: utility cost and FCI. Simulation and optimization of RDWC retrofit are described in Section 3. Section 4 presents the results of multi-objective optimization (MOO) studies mentioned in Sections 2 and 3. It also presents economics of RD and RDWC solutions selected from the respective Pareto-optimal front. Finally, Section 5 summarizes the main findings of this work.

2. FA process retrofit with an RD column 2.1 RD retrofit development and simulation The retrofitted FA process using an RD column is shown in Figure 3. In this, product purity was set to 98 wt% and the production capacity was set to 27,100 t/y, according to the base case FA process. Further, design parameters of all units, except RD and DC2, were taken from the base case simulation. RD unit was simulated in Aspen Plus V9.0 using a RADFRAC column and setting the reactive zone at the upper part of the column. Thus, the top of the RD column corresponds to 7 ACS Paragon Plus Environment


the reactive section. Distillate leaving the condenser (stream 21) is CO purge stream, with flowrate of only 0.20 kmol/h, which is much lower than the bottoms flowrate of 512.31 kmol/h. This configuration (reactive zone at the top, low distillate flowrate) is similar to that adopted by Wang14 and Wang et al.29, in which the RD column was operated at full reflux. In both systems, the lightest (more volatile) component was a reactant. Full reflux configuration is then used to prevent loss of light reactant in the distillate stream. The thermodynamic model chosen here (for both RD and RDWC simulations) was UNIQUAC with Hayden O’Connell equation of state for vapor phase. In addition, Henry law was used to calculate the amount of CO dissolved in liquid phase. Binary parameters and model validation are available in da Cunha et al.19

Figure 3 – The RD retrofitted process for FA production; stream data corresponding to the inlet/outlet streams of the RD column in the chosen optimal retrofit are given in Table A1 In the RD unit in Figure 3, the reactive section is packed with a strong ion-exchange resin, which is the catalyst for MF hydrolysis. Kinetics for this reaction are given by:14


Page 8 of 41

Page 9 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


𝑟=―

𝑐𝐶𝑐𝐷 1 𝑑𝑐𝐵 𝐸 = 𝑘0exp ― 𝑐𝐴𝑐𝐵 ― 𝐾 [𝑐𝑎𝑡] 𝑑𝑡 𝑅𝑇

)(

(

𝑘0 =

)

6.530 × 106

(5)

1 + 0.0869 × 𝑐2𝐴

(

𝐾 = 0.4492exp ―

251 𝑇

(4)

)

(6)

In these equations, ko is the pre-exponential factor (kmol×h-1×kg-1cat), E refers to the activation energy (63100 kJ.kmol-1), R is the universal gas constant (8.314 J×K-1×mol-1), T is the temperature (K), and cA, cB, cC and cD are respectively concentration of water, MF, FA and MA, in kmol.m-3. The water adsorbed by the resin limits the contact of other chemicals with the active surface. This phenomenon affects the performance of the catalyst,14 as seen by the denominator in eq. 5. The kinetics (eq. 4) for MF hydrolysis is not in the standard power-law form, and therefore it cannot be input into Aspen Plus via its user interface. So, it was included as a FORTRAN subroutine instead. This subroutine is compiled using the Customize Aspen Plus V9 tool. By using a subroutine, it is possible to write user-defined equation(s) to calculate the rate of reaction. Variables such as temperature and liquid concentrations in each stage are accessed from the process simulation and used to calculate r in eq. 4. Instead of using a FORTRAN subroutine to declare the kinetics, one could assume that the effect of water on catalyst performance is small (i.e., 1 + 0.0869 × 𝑐2𝐴  1). In this approximate case, the kinetics can be entered directly into the Aspen Plus kinetic sheet. The distillation column DC2 following the RD unit in Figure 3 separates methanol from FA and water. Because FA is present in the feed, the column shell and internals should be corrosion-resistant. In the base case design, FA was part of DC2 feed, but DC3 feed was free from acid. Therefore, DC2 was designed to resist FA corrosion (zirconium shell, nickel alloy trays), whereas DC3 (carbon steel for shell and trays) was not. Hence, DC2 is reused instead of DC3 in the retrofit configuration, and it has to operate between 1.62 bar and 2.33 bar to avoid entrainment. No other FA-resistant column is available from the base case process, and hence it is necessary to purchase a new column for the RD. Parameters for preliminary RD retrofit are given in Table 1. No. of stages refers to number of theoretical plates, numbered from top to bottom excluding condenser and reboiler. Besides, both 9 ACS Paragon Plus Environment


Page 10 of 41

MF and water feeds are in the liquid phase. No. of theoretical stages in DC2 was calculated based on the no. of real trays from the base case design (68) and DC2 efficiency (= 48%) estimated using the O’Connell equation in Wankat:30 𝐸 = 0.52782 ― 0.27511 × log10 (𝛼𝜇) + 0.044923 × [log10 (𝛼𝜇) ]2

(7)

Here, E is the column efficiency, and α and μ are respectively the relative volatility of the key components and the liquid viscosity of feed (in cP), both determined at average temperature and pressure of the column. MF conversion in the RD column, and MA recovery and purity in DC2 distillate stream were chosen as design specifications for RD and DC2 units, respectively. These values were set to ensure that stream compositions in the RD process are similar to those in the base case process. Heat integration for the RD retrofitted process was based on pinch analysis. Except for reboiler and condenser of the new RD column, all remaining heat exchangers (HEs) in the RD process were taken from the equipment used in the base case process. Note that heat integration is not shown in Figure 3 for clarity of process flowsheet, but the corresponding heat exchanger network (HEN) can be found in Appendix B.


Page 11 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Table 1 – Parameters for preliminary RD retrofit No. of reactive stages in RD

51

No. of separation stages in RD

8

RD feed stage for MF (stream 15)

1

RD feed stage for water (stream 18)

1

RD pressure (bar)

4.76

MF conversion

0.991

DC2 no. of stages

32

DC2 feed stage (stream 25)

25

DC2 pressure (bar)

2.03

MA recovery in DC2 distillate

0.995

MA purity in DC2 distillate

0.9946

2.2 RD column sizing, internals and cost estimation Sizing and costing of the RD column were made assuming structured packing. The reactive zone uses Katapak-S 250.Y, and the separation section uses Mellapak 250.Y. For both these packings, HETP is about 0.4 m.31 Considering height allowances from Branan32 for top section, liquid distributors, liquid redistributors, liquid collectors and bottom section, the total height of the column is ~34 m for the preliminary design. Column diameter was calculated via Aspen Plus, using built-in hydraulics correlations for Mellapak 250.Y. The diameter obtained was 0.84 m, which results in column height over diameter ratio of 35.7. Heuristics suggest that this ratio should not exceed 30,33 and therefore diameter was increased to 1.08 m. The volume of one theoretical stage with diameter of 1.08 m and HETP of 0.40 m is ~0.366 m3. Assuming catalyst occupancy of 24.8% (within the range given in the Introduction), the volume of catalyst per stage is ~0.091 m3. Assuming a resin with density of 1,200 kg/m3 (value taken from Sigma-Aldrich’s catalogue34 for Amberlyst 36), mass of catalyst per stage is ~109 kg. The amount of catalyst in each stage is important for reaction rate calculations because the basis for kinetics in eq. 4 is mass of catalyst. To evaluate the capital cost of the RD column, unit costs for structured packing and for liquid distributors were taken from Seider et al.35 These costs were inflated to CEPCI of 600, which gives $305/ft3 for structured packing and $152/ft3 for liquid distributors. Packing cost is given for 11 ACS Paragon Plus Environment


structured packing of corrugated-sheet type in stainless steel, such as Mellapak and Katapak. The authors confirmed with their industrial contact that stainless steel 904L and above can be used in the FA process. Cost of column shell was estimated according to Turton et al.21 It has been reported that Zirconium alloy is used as material of construction for column shell in the Kemira-Leonard FA process.36 However, material factor for zirconium shell is not available in the literature. Data from Couper et al.37 suggest that titanium and zirconium storage tanks have the same cost, and we assume that this remains valid for vessels. Therefore, material factor of 9.4 reported for titanium vessels21 is assumed for the RD shell of zirconium alloy. To estimate the cost of catalyst, prices of three different types of ion-exchange resins used in Jogunola et al.15 for MF hydrolysis reaction were compared. The resins used in their work were Amberlite IR-120 and Dowex 50Wx8-50/100/400. The authors could not find cost data for Dowex 50Wx8-50. Among the remaining resins, Dowex 50Wx8-100 is the most expensive, with unit cost of 350.4 USD/kg.38 This value is used to estimate the total cost of catalyst in the FA process developed. Finally, catalyst lifetime is assumed as 4 years; this lifetime was provided by our industrial contact for an ion-exchange resin used in the MTBE RD process. Therefore, cost of catalyst is divided by 4 years and included in the annual cost of catalyst/utilities, together with the cost of refilling estimated as 50% of the packing price for Katapak.39 Other utility costs include cost of steam, electricity, cooling water and fuel oil. Unit cost of these utilities is summarized in Table 2. Table 2 – Unit cost of utilities21 Utility

Cost (USD/GJ)

Cooling water

0.244*

Chilled water

4.43

Low pressure steam

13.28

Medium pressure steam

14.19

Fuel oil

14.2

Electricity

16.8

* Reference21 gives cost of cooling water for temperature change from 30oC to 40oC; here, we assumed this range to be from 30oC to 45oC, and proportionately updated the cost to $0.244/GJ.


Page 12 of 41

Page 13 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


The cost breakdown for the preliminary RD retrofit is summarized in Table S1 in the Supporting Information. This table shows that the preliminary configuration adopted in the RD retrofitted process is not profitable. Indeed, TAC after retrofit increases slightly from 0.686 USD/kg19 to 0.695 USD/kg. However, preliminary design can be improved by optimization, and so RD retrofit may be profitable for a different choice of design parameters. In Section 2.3, we describe the objective functions, decision variables and optimization program for MOO of the RD process. 2.3 Multi-Objective Optimization of RD retrofit The goal of retrofit in chemical facilities is often to reduce operating cost with a minimum capital investment. However, these two objectives are usually conflicting, and thus one cannot conceive a design minimizing both operating and capital costs. Optimization of two or more conflicting objectives will generate a set of non-dominated solutions called Pareto-optimal front. A specific solution from this set can be chosen based on other relevant factors. The RD retrofitted process has been optimized for two objectives: minimize FCI and cost of utilities and catalyst, simultaneously. Design parameters of downstream and upstream units of RD and DC2 (except for feed stage and pressure of DC2) were kept unchanged in order to avoid further equipment replacement. Therefore, decision variables (DVs) for optimization can only include RD and DC2 parameters. Relevant DVs, chosen after sensitivity analysis, and their ranges are summarized in Table 3. Table 3 – DVs and their ranges for the RD retrofitted process DV

Lower bound

Upper bound

Chosen optimal solution

No. of reactive stages in RD

27

51

34

No. of separation stages in RD

1

10

7

RD pressure (bar)

4.76

5.37

5.37

DC2 feed stage (stream 25)

25

27

25

DC2 pressure (bar)

1.62

2.33

1.68

To solve the MOO problem described in this section, the Integrated Multi-Objective Differential Evolution (IMODE) program40 was used. This program was integrated with Aspen 13 ACS Paragon Plus Environment


Plus simulation via Happ library, available in Aspen software package. Initial values of crossover and mutation probabilities in IMODE algorithm were chosen as 0.5; population size was set to 50 and maximum number of generations to 100. Maximum number of iterations to achieve convergence of tear streams in Aspen Plus simulation was set to 4000, which was found necessary due to the number of recycle loops in FA process (Figure 3). For some trial solutions, Aspen Plus simulation did not converge within this maximum number of iterations. In such non-converged situations, objective function values for trial solutions generated by the optimizer were set to a large pre-fixed number (1×1010). Since the objectives have to be minimized, the trial solutions resulting in non-converged simulations were later discarded from subsequent generations of optimization.

3. FA process retrofit with an RDWC 3.1 RDWC retrofit development and simulation Figure 4 shows the retrofitted FA process using RDWC, which replaces CSTR2, DC2 and DC3 in the base case FA process in Figure 1. Once again, purity, production capacity and the design parameters of all units except RDWC are the same as those in the base case process.19 RDWC unit was simulated in Aspen Plus V9.0 using two RADFRAC columns and setting the reactive zone at the upper part of the first column, as shown in Figures 5a and 5b. The top-wall RDWC configuration in Figure 5a was also employed in recent studies on intensified FA processes.11,13 Also, the same configuration was used by Wang et al.29 for methyl acetate hydrolysis. By separating the reactive section from section 2 (Figure 5) and by setting a low distillate rate for the CO purge stream, loss of MF in the distillate streams is avoided. Since there is no gas to be purged from the methyl acetate hydrolysis reaction, the condenser above the reactive section operates at full reflux in that process.29 In Figure 5b, a liquid stream flows from the bottom of section 1 to a certain stage of the main column (having sections 2 and 3), and a vapor stream is withdrawn from this same stage and sent to the bottom of section 1. This configuration creates a recycle loop in RDWC simulation, and this vapor stream (labelled TS in Figure 5b) is set as the tear stream to be converged in this loop. Specifications for the columns in Figure 5b are: (i) CO purge flow rate (= 0.20 kmol/hr); (ii) MA+MF mole fraction in the distillate stream of right column (= 0.9999); (iii) distillate rate of right column; and (iv) MF recovery in distillate and bottoms of right column. Specification (i) is 14 ACS Paragon Plus Environment

Page 14 of 41

Page 15 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


an operating specification, and has been set to a fixed value. Design specification (ii) was set to a fixed value, to be achieved by varying the reboiler duty. Operating specification (iii) is calculated via a calculator block, based on flow rate and composition of the two feeds.

Figure 4 - The RDWC retrofitted process for FA production; stream data corresponding to the inlet/outlet streams of the RDWC column in the chosen optimal retrofit are given in Table A2

(a)

(b)

Figure 5 – (a) RDWC scheme and (b) the equivalent two-column configuration for simulation 15 ACS Paragon Plus Environment


Page 16 of 41

As can be seen in Figure 4, RDWC retrofitted process has several recycle loops (besides that in RDWC), and so the composition and flows of streams 15 and 18 (feeds to RDWC) may change slightly during each iteration for converging on these recycles. The distillate rate on the main column is calculated based on the feeds in the current iteration, considering MA+MF recovery of 99.4% and mole purity of 99.99%. Finally, design specification (iv) on MF recovery in distillate and bottoms of right column is set with the help of design specs block in Aspen Plus. MF recovery in streams 23 and 29 with respect to inlet streams 15 and 18 (Figure 4) is set to a low value of 0.0071, which is equivalent to MF conversion of 0.993. This conversion was calculated based on results for the base case process. Specification (iv) is achieved by varying flowrate of vapor stream TS in Figure 5b. Eqs. 4 to 6 are the kinetics for the hydrolysis reaction. Design parameters used in the preliminary RDWC retrofit are shown in Table 4. Both water and MF feeds are in the liquid phase. Specifications related to units other than RDWC were kept unchanged from the base case design, in order to avoid further equipment replacement. Indeed, apart from RDWC itself, only 3 new HEs are required. Rest of the units from the base case process can be re-used. Table 4 – Parameters for the preliminary RDWC retrofit No. of reactive stages (reactive section)

51

No. of stages in section 1

9


60


12

RDWC pressure (bar)

4.76

Water feed sage (stream 18)

1

MF feed stage (stream 15)

1

MF conversion

0.993

Heat integration for the RDWC retrofitted process was based on pinch analysis. Except for the RDWC reboiler and two extra HEs, all remaining HEs in the RDWC process were taken from the equipment used in the base case process. The heat integration is not shown in Figure 4 for clarity, but the corresponding HEN can be found in Appendix B.


Page 17 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


3.2 RDWC column sizing, internals and cost estimation In any retrofit, equipment available in the original plant should be reused as much as possible, in order to minimize capital investment. Therefore, we assess whether one of the replaced columns (i.e. DC2 and DC3) could be reused as an RDWC. Column DC3 is not made of corrosionresistant material because it does not perform FA separation in the base case process. Therefore, this column shell cannot be employed for the new RDWC unit. DC2 was designed to resist FA corrosion, but its diameter (~ 1.1 m) is smaller than the diameter required for the RDWC to operate at 80% capacity (1.3 m according to preliminary and optimal designs). Therefore, neither DC2 nor DC3 can be reused for the RDWC shell, and a new unit (shell + internals) needs to be purchased. Sections 1, 2 and 3 of the columns in Figure 5b were sized according to their internal flows, using Aspen Plus column internals feature. Here again, we have considered Mellapak 250.Y structured packing for sizing. Similar to the RD column, HETP was set to 0.40 m. Diameter of the reactive zone was fixed to 1.08 m in order to accommodate 0.091 m3 of catalyst within each theoretical stage. A stage with diameter of 1.08 m and HETP = 0.4 m corresponds to a volume of 0.37 m3, and thus catalyst volume corresponds to 25% of the stage. The diameter of the RDWC was calculated as follows. First, we computed the total cross-section area of three different parts of the column: section 3, sections 1 + 2, and reactive section + section 2. We chose the largest area among them and then calculated the equivalent diameter. Knowing the overall diameter, one can estimate the capital cost of the column and its internals. Costing methodology for the RDWC retrofit is the same as that described in Section 3.2. Table S2 in the Supporting Information shows the cost breakdown for the RDWC preliminary retrofit. Table S2 shows that the preliminary configuration adopted in the RDWC retrofitted process is promising, as TAC after retrofit decreases from 0.686 USD/kg19 to 0.676 USD/kg. Besides, preliminary design can be improved by optimization, and so the RDWC retrofit may increase savings even further for a different choice of design parameters. In the next section, objective functions, decision variables and optimization program for MOO of the RDWC process are described. 3.3 Multi-Objective Optimization of RDWC retrofit Similar to the RD retrofit optimization, RDWC process has been simultaneously optimized for minimization of FCI and cost of utilities and catalyst. Design parameters of downstream and 17 ACS Paragon Plus Environment


Page 18 of 41

upstream units of RDWC were kept unchanged in order to avoid further equipment replacement. Relevant DVs were chosen after sensitivity analysis, and their ranges are summarized in Table 5. Similar to the MOO problem solution described in Section 2.3, IMODE program40 was used to optimize the RDWC process. Initial values of crossover and mutation probabilities in IMODE algorithm were chosen as 0.5, and the population size was set to 50. The program was terminated after 30 days, resulting in a total of 28 generations. The very low speed observed in this MOO run (~ 1 generation/day) is due to the simulation setup in Aspen Plus. The use of two-column configuration for RDWC simulation (Figure 5b) creates an internal convergence loop, where TS is the tear stream targeted for convergence. Furthermore, the design specs block described in Section 3.1 creates another loop. For particular RDWC feed compositions (streams 15 and 18), the total flowrate of stream TS will be set to an initial value, and a number of iterations will be computed until TS composition and temperature converge. The MF conversion is then calculated, and if this value is close to 0.993 no more iterations are required for the RDWC column. If the conversion is significantly different from 0.993, TS flowrate is adjusted and a new set of iterations is computed until TS composition and temperature converge. This procedure is repeated until MF conversion in the column equals 0.993, within certain tolerance. Table 5 – DVs and their ranges for the RDWC retrofitted process Chosen optimal

DV

Lower bound

Upper bound

No. of reactive stages

21

51

29


2

11

9


15

39

38


6

14

6

RD pressure (bar)

4.76

5.67

5.67

solution

Therefore, the flowsheet in Figure 4 has 3 levels of iterations, to account for overall flowsheet convergence (outer level), MF conversion in the design specs block (middle level) and convergence of tear stream TS in Figure 5b (inner level). Maximum number of iterations set for outer, middle and inner levels were 1200, 500 and 9999, respectively. These levels and iterations 18 ACS Paragon Plus Environment

Page 19 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


result in very slow convergence of RDWC process simulation. Similar to the approach for the RD retrofit optimization, objective function values of non-converged trial solutions were set to 1×1010 so that these solutions will be discarded in the subsequent generations.

4. Results and discussion 4.1 Sensitivity analysis Sensitivity analysis gives insights on how each of the decision variables affects the objectives. To perform this analysis, all the decision variables are set to predefined values, and then they are changed one at a time in order to observe the trends in each objective function. Figures 6 and 7 depict the sensitivity analysis for the RD and RDWC retrofits. In both these figures, Utility cost axis refers to the cost of utilities and catalyst.



Figure 6 – Sensitivity analysis for the RD retrofit As depicted in Figures 6a and 6b, no. of separation stages and no. of reactive stages in the RD column have significant influence on the objectives. Conflicting behavior is observed for no. of separation stages higher than 5, and for no. of reactive stages 40 and above. As expected, higher no. of stages generally increases FCI and decreases the utility cost. The high FCI value for 2 separation stages (Figure 6a) is due to the large increase in RD reboiler duty, which significantly increases the cost of RD reboiler. Fluctuation of FCI in Figure 6b is due to the following reason. 20 ACS Paragon Plus Environment

Page 20 of 41

Page 21 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


HE 5 in the base case (see Figure A1 in da Cunha et al.19) is reused here as the condenser of RD column. For 27 and 34 reactive stages, the area of this condenser is much higher than the area of HE 5, and therefore a new condenser needs to be purchased. Beyond these values, condenser area is sufficiently small for reusing HE 5 from the base case, which explains the drop in FCI at 39 reactive stages. Finally, the increase in utility cost at 51 reactive stages (Figure 6b) happens because the extra annual cost for catalyst and its refilling negates the savings in low pressure steam.

Figure 7 – Sensitivity analysis for the RDWC retrofit 21 ACS Paragon Plus Environment


The conflicting behavior of objective functions is clearly observed by varying the RD column pressure (Figures 6c). FCI increases with pressure for two reasons: increase in the shell thickness and in RD reboiler area. Reboiler area increases because the bottoms of RD column will be at a higher temperature, and consequently the temperature difference between this stream and low pressure steam will be smaller, resulting in larger area requirement. The decrease in utility cost as pressure increases comes from the decrease in reboiler duty for columns RD and DC2. The increase in temperature profile due to higher pressure accelerates the reaction in the RD column, requiring less MF reflux for the reaction to take place and thus decreasing the reboiler duty of RD column. Besides, an increase in vapor fraction of stream 25 results in slightly lower DC2 reboiler duty at higher RD pressure. Further, Figures 6d and 6e show that DC2 feed stage and pressure affect only the utility cost. Sensitivity analysis for the RDWC retrofit also provides valuable insights on the effect of each decision variable on the economics of the process. The unexpected decrease in FCI from 21 reactive stages to 27 reactive stages (Figure 7a) happens because the column height remains unchanged, whereas the diameter decreases. Height remains the same because section 2 of the column (see Figure 5) has a total of 39 stages, higher than the no. of stages of reactive section (21 or 27) + section 1 (9). The diameter decreases because the reboiler duty of the column is slightly lower. However, further increase in no. of reactive stages to 33 will affect the column height, since the sum of no. of reactive stages and stages in section 1 is more than 39. As a result, FCI for the column increases. The subsequent decrease in FCI at 39 reactive stages is due to savings in heat exchangers. Indeed, only for 39 or more reactive stages, the area required for the RDWC condenser at the pre-fractionator side (see Figure 5) is small enough so as to reuse HE 7 from the base case process (see Figure A1 in da Cunha et al.19). Finally, unexpected high FCI for 2 stages in section 1 (Figure 7b) is due to the high reboiler duty in the RDWC. This high duty increases the cost of reboiler and the column diameter, negating the expected savings from reduction in column height. The variation of objective functions observed in Figures 7c and 7d is as expected. However, it is interesting that the no. of stages in section 3 has very little impact on the utility cost of the process. Finally, RDWC pressure is the decision variable with less effect on FCI. The change in utility cost may be related to a change in the relative volatility of components, which affects the reboiler duty of RDWC.


Page 22 of 41

Page 23 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60




4.2 MOO results and economics of chosen solutions 4.2.1 RD retrofit The evolution of the Pareto-optimal solutions for the RD retrofit optimization, as well as its final results, are shown in Figure 8. In this figure, Total Module Cost refers to the cost of the new RD unit and its reboiler. As depicted in the insert in Figure 8, there is no significant improvement of the solutions after 20 generations. The MOO converged according to the Chisquared termination criterion40 after 55 generations, which further confirms the accuracy of the Pareto-optimal front obtained. For clarity, the final set of non-dominated solutions is shown separately in Figure 8 itself.

Figure 8 – Pareto-optimal front (main plot) and intermediate solutions (insert) for the RD retrofit Utility cost decreases from 0.2137 USD/kg to 0.2116 USD/kg for an increase of $46,000 in FCI. Further increase in FCI by $33,000 results in savings of only 0.0002 USD/kg in utility cost. Finally, utility cost decreases from 0.2114 USD/kg to 0.2105 USD/kg for a corresponding increase in FCI from $1,251,000 to $1,356,000. The chosen optimal solution is the solution in the Paretooptimal front which minimizes TAC after retrofit (TACrt in Eq. 3). This solution, hereafter referred simply as optimal RD solution, is identified with a filled black bullet (•) in Figure 8. The DV values corresponding to this solution are shown in Table 3. In order to calculate TACrt using eq. 3, cost of new equipment has been amortized considering 8-years period. This period comes from two assumptions: plant lifetime of 15 years41 and retrofit implementation after 7 years of operation. 24 ACS Paragon Plus Environment

Page 24 of 41

Page 25 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Considering an annual interest rate of 0.15 as before gives ACCR = 0.223. Cost analysis for the optimal RD solution is given in Table 6. Table 6 – Cost breakdown for the optimal RD retrofit FCI (Total Module Cost) RD shell

$670,000

Packing + liquid distributors/redistributors

$200,000

RD reboiler

$430,000

Total FCI

$1,300,000

Utility and raw material cost (USD/kg of FA) Steam

0.187

Cooling water

0.011

Electricity

0.000

Fuel oil

0.000

Catalyst cost

0.013

Total utility cost

0.211

Total raw material cost

0.085

TAC after retrofit (using eq. 3)

0.688 USD/kg

TAC for optimal base case19

0.686 USD/kg

Savings in utility cost

0.0075 USD/kg

Savings in raw material cost

0.0004 USD/kg

Results in Table 6 show that RD is not profitable from the retrofit perspective. Very small savings from raw material come from lower MA make-up rate, achievable since the RD unit and column DC2 together (Figure 3) purge less material than columns DC2 and DC3 combined in the original process (Figure 1). Savings in utilities are small, and reasons for this are as follows, similar to those given for the DWC retrofit19. First, DC2 and DC3 in the base case were operating in a double-effect configuration. This configuration allows integration of DC2 condenser (at 3.78 bar) with DC3 reboiler (at 1.01 bar). When CSTR2 and DC3 are replaced by the RD column, doubleeffect operation is no longer possible. Inlet temperature of RD condenser (85.6°C) is not sufficiently high to provide heat to DC2 reboiler (at 117°C). Second, stream 23 from DC2 bottoms 25 ACS Paragon Plus Environment


in the base case process (at 147.3°C) was used to provide energy for the reboiler of DC5. After RD retrofit, stream 23 is at a lower temperature (118.8°C), and so the heat available for use in DC5 reboiler is lower. Therefore, more low pressure steam is required to boil up the bottoms of DC5. The liquid composition profile of the optimal RD unit is shown in Figure 9. In this figure, stages 1 and 43 are respectively the condenser and reboiler of the column. The reactive section is between (and including) stages 2 and 35, and both inlet streams 15 and 18 in Figure 3 (i.e., MF and water in liquid phase) are fed to the top of the column (stage 2). The reactants and products show an almost linear trend within the reactive section (Figure 9). The temperature also increases linearly within the reactive zone. A similar linear profile was observed for the methyl acetate hydrolysis in a total reflux RD column.42 The separation section (stages 36 to 42 inclusive) is mainly responsible for recovering unreacted MF and returning it back to the reactive zone.

Figure 9 – Temperature and liquid composition profile of the optimal RD column 4.2.2 RDWC retrofit The evolution of the Pareto-optimal solutions for the RDWC retrofit optimization, as well as its final results, are shown in Figure 10. In this figure, Total Module Cost refers to the cost of the new RDWC unit and 3 extra HEs: RDWC reboiler, HE 2 and HE 5 (Figure B2). As depicted in the insert in Figure 10, there is no significant improvement of solutions after 20 generations. For clarity, non-dominated solutions are shown separately within Figure 10. 26 ACS Paragon Plus Environment

Page 26 of 41

Page 27 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Figure 10 - Pareto-optimal front (main plot) and intermediate solutions (insert) for the RDWC retrofit Utility cost decreases from 0.1813 USD/kg to 0.1764 USD/kg for an increase of $28,000 in FCI for RDWC retrofit. Further increase in FCI by $180,000 results in savings of only 0.001 USD/kg in utility cost. Similar to the previous section, the chosen optimal solution is the solution in the Pareto-optimal front, which minimizes TACrt, and it is identified with a filled black bullet (•) in Figure 10. The DV values corresponding to this solution are shown in Table 5. As before, amortization period for capital investment was taken as 8 years. Cost analysis for the optimal RDWC solution is given in Table 7. The maximum temperature difference across the wall for the chosen optimal solution is around 28°C, well below the limit of 40°C when mechanical stress is observed.27 Results in Table 7 show that RDWC is profitable from the retrofit perspective. Similar to the RD retrofit, small savings for raw material come from lower MA make-up rate, achievable since the RDWC unit (Figure 4) purges less material than columns DC2 and DC3 combined in the original process (Figure 1). The relatively significant savings in utility cost is due to a significant reduction (by 20.4%) in thermal energy consumption, which comes mainly from the heat integration between the RDWC condenser at the top of section 2 (Figure 5) and the reboiler of DC4. This integration corresponds to HE5 in Figure B2, resulting in lower steam consumption and savings of 0.033 USD/kg in utility cost. Remaining savings come from lower steam consumption 27 ACS Paragon Plus Environment


Page 28 of 41

in DC5 reboiler (0.003 USD/kg) and lower utility consumption in RDWC compared with CSTR2 + DC2 + DC3 in the optimal base case (0.016 USD/kg). Some of these savings are offset by the cost of catalyst in RDWC, resulting in the total savings in utility cost given in Table 7. The payback period necessary for the savings in utility and raw material cost to offset the capital investment is 2.5 years. Table 7 – Cost breakdown for the optimal RDWC retrofit FCI (Total Module Cost) RD shell

$1,140,000

Packing + liquid distributors/redistributors

$320,000

Extra heat exchangers

$1,470,000

Total FCI

$2,930,000

Utility and raw material cost (USD/kg of FA) Steam

0.155

Cooling water

0.010

Electricity

0.000

Fuel oil

0.000

Catalyst cost

0.011

Total utility cost

0.176

Total raw material cost

0.084

TAC after retrofit (using eq. 3)

0.666 USD/kg

TAC from optimal base case19

0.686 USD/kg

Savings in utility cost

0.042 USD/kg

Savings in raw material cost

0.001 USD/kg

Note that most of the savings in utility cost comes from the new heat integration between one of RDWC condenser and the reboiler of DC5. Without that, savings in utility cost would be reduced to 0.009 USD/kg, close to the savings found for the RD process, but with higher capital investment. This shows that the double effect configuration of columns DC2 and DC3 in the base case process is very energy-efficient, and retrofits affecting this configuration may not be profitable. 28 ACS Paragon Plus Environment

Page 29 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


In the liquid composition profiles of the optimal RDWC unit (shown in Figure 11), there are two different profiles between stages 1 and 39. These profiles correspond to the left (prefractionator) and right of the wall (see Figure 5). The reactive section is placed between (and including) stages 2 and 30 in the pre-fractionator, and both inlet streams 15 and 18 in Figure 4 (i.e., MF and water in liquid phase) are fed to the top of the reactive zone (stage 2). The trend of composition profiles in the RDWC reactive section is very similar to that in the reactive section of the RD column, with reactants and products varying linearly. Water mole fraction decreases from 0.66 to 0.54 towards the end of the reactive zone, and MF mole fraction decreases from 0.29 to 0.19. Further, MA mole fraction increases from 0.05 to 0.17, and FA mole fraction increases from 0.00 to 0.10 towards the end of the reactive section (stage 30). These ranges are close to those observed in Figure 9 for RD column. Just below the reactive section, section 1 (see Figure 5) recycles unreacted MF back to the reactive zone.

Figure 11 – (a) Liquid composition profile and (b) temperature profile of the optimal RDWC column The right side of the wall (section 2 in Figure 5) separates MA from water, and pure MA is obtained as the distillate at the top of this section. The bottom section of the RDWC column (section 3 in Figure 5, stages 40 to 45 in Figure 11) separates MA from aqueous FA. MF concentration is not shown between stages 40 to 46 because it is present in only trace amounts in this region. Finally, as shown in Figure 11, temperature in section 2 of the RDWC (Figure 5) is mostly higher than the temperature in the pre-fractionator side. This is because the light key in the pre-fractionator side is MF, whereas light key in section 2 is MA (which is less volatile than MF). Economics of RDWC retrofit can be improved even further if we consider alternative scenarios. For example, if we assume that the columns and HEs replaced in this retrofit can be 29 ACS Paragon Plus Environment


sold, it is possible to recover $1,090,000; this is assuming an annual depreciation of 0.067 (1/15) over a usage period of 7 years. In this scenario, TACrt drops to 0.657 USD/kg, and the payback period of investment for retrofitting decreases to 1.6 years. Another possibility is to use zirconium cladding instead of zirconium for RDWC shell. The material factor for zirconium cladding is assumed to be the same as the material factor for titanium cladding (= 4.7).21 Using this alternative material of construction brings FCI down to $2,410,000, resulting in TACrt of 0.662 USD/kg and payback period of 2.1 years. We can also evaluate a third optimistic scenario with a lower catalyst cost. Among the options mentioned in Section 2.2, Amberlite IR-120 is the cheapest with unit cost of 48.4 USD/kg.43 In this scenario, TACrt is 0.657 USD/kg, with payback period of 2.1 years. If all these options (i.e., selling of replaced equipment, zirconium cladding and lower catalyst cost) are considered together, TACrt reduces to 0.644 USD/kg, with payback period of only 0.9 years for the investment for retrofitting. It is not possible to compare our process configuration and results with others in the literature. Two recent studies11,13 on FA process simulation propose alternative configurations for the production of FA, but with lower purity requirement of 85 wt%. Another article12 discusses process alternatives for the production of 99 wt% FA. Upstream process configuration in this process is similar to that shown in Figure 1: CSTR1, F and DC1 are common units for both configurations. However, hydrolysis and subsequent separation steps differ significantly. The downstream process configuration minimizing energy consumption in Novita et al.12 consists of an RD column followed by a DWC. In this paper, FCI is not reported, and not enough information is available to estimate the capital investment for the proposed plant. Further, amount of catalyst in the RD column and electrical energy consumption are not given. Also, there are significant mass balance errors in the process diagram presented for RD and DWC units (inlet/outlet flowrates are 660.4/625.3 kmol/h for RD and 624.5/659.5 kmol/h for DWC)12, which may be due to typographical and/or simulation errors. Finally, Novita et al.12 assumed catalyst (an ion-exchange resin) density of 3,600 kg/m3, much higher than that used in the present work (1,200 kg/m3) taken from Sigma-Aldrich Handbook.34 Assuming catalyst density of 3,600 instead of 1,200 kg/m3 triples the rate of reaction in each stage.


Page 30 of 41

Page 31 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


5. Conclusions A two-part study has been conducted to investigate several retrofit alternatives for FA process based on an industrial configuration. In the first part,19 the base case process was simulated and optimized; then, retrofit via DWC was assessed and found to be uneconomical. In the second part (present paper), RD and RDWC retrofits for FA process were simulated and optimized to assess their techno-economic feasibility. RD was found to be a better option than DWC, though it is still inadequate to improve economics of the optimal base case. These unusual results for RD and DWC retrofits are due to the double-effect configuration of DC2 and DC3 in the base case (Figure 1), which brings large energy savings. Only by adding an even more intensified unit (RDWC), sufficient benefits are possible from retrofit. The savings in utility and raw material costs from the RDWC retrofit (0.043 USD/kg) overcome the additional capital investment of $2,930,000, resulting in TAC after retrofit of 0.666 USD/kg. Table 8 summarizes the economics of base case and all three retrofit options. Of these, RDWC retrofit is economically attractive, and it also reduces utilities (mostly steam energy by 20.4%) consumption. Economics of RDWC retrofit can be further improved to 0.644 USD/kg by selling replaced equipment, using zirconium cladding as material of construction and lower catalyst cost. Table 8 – Cost analysis for the base case FA process and its retrofit options DWC

RD

RDWC

Retrofit

Retrofit

Retrofit

16,280,000

NA

NA

NA

Addition capital investment (USD)

NA

2,000,000

1,300,000

2,930,000

Utility cost (USD/kg)

0.218

0.218

0.211

0.176

Raw material cost (USD/kg)

0.085

0.084

0.085

0.084

Production capacity (t/y)

27,100

27,100

27,100

27,100

TAC (USD/kg)

0.686

0.701

0.688

0.666

Factor

Base case

Grassroots investment (USD)

NA – Not applicable The retrofit options studied here were derived in Part I of this work,19 using the retrofit methodology developed by Niu and Rangaiah.44 However, other plant-wide retrofit methodologies (such as the one developed by Lorenz et al.45) can be applied to screen for alternative retrofit solutions. Furthermore, RDWC simulation results should be validated, for example, with 31 ACS Paragon Plus Environment


experiments before implementing this retrofit. Lab- and pilot-scale arrangements for RDWC experiments described in the literature29,46 can be used for this purpose.

Supporting Information Cost breakdown for the preliminary RD and RDWC retrofits.


Page 32 of 41

Page 33 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Nomenclature ACCR

Annual Capital Charge Ratio

CO

Carbon Monoxide

COM

Cost of Manufacture (USD/kg)

CSTR

Continuous Stirred Tank Reactor

DC

Distillation Column

DIPE

Diisopropyl Ether

DV

Decision Variable

DWC

Dividing-Wall Column

FA

Formic Acid

FCI

Fixed Capital Investment ($)

HE

Heat Exchanger

HEN

Heat Exchanger Network

HETP

Height Equivalent to a Theoretical Plate

IMODE

Integrated Multi-Objective Differential Evolution

MA

Methanol (Methyl Alcohol)

MOO

Multi-Objective Optimization

MTBE

Methyl Tert-Butyl Ether

MF

Methyl Formate

PR

Production Rate (kg/y)

RD

Reactive Distillation

RDWC

Reactive Dividing-Wall Column

TAC

Total Annual Cost (USD/kg)

TACrt

Total Annual Cost After Retrofit (USD/kg)

References (1)

Hietala, J.; Vuori, A.; Johnsson, P.; Pollari, I.; Reutemann, W.; Kieczka, H. Formic Acid. In Ullmann’s Encyclopedia of Industrial Chemistry; Wiley: Weinheim, 2016.

(2)

Rice, C.; Ha, S.; Masel, R. I.; Waszczuk, P.; Wieckowski, A.; Barnard, T. Direct Formic 33 ACS Paragon Plus Environment


Acid Fuel Cells. J. Power Sources 2002, 111, 83–89. (3)

Chang, J.; Feng, L.; Liu, C.; Xing, W.; Hu, X. An Effective Pd-Ni2P/C Anode Catalyst for Direct Formic Acid Fuel Cells. Angew. Chemie - Int. Ed. 2014, 53, 122–126.

(4)

Leonard, J. D. Preparation of Formic Acid by Hydrolysis of Methyl Formate. 4,299,981, 1981.

(5)

Moiseev, I. I.; Tagaev, O. A.; Zavoronkov, N. M.; Pazderskij, J. A.; Kaleciz, I. V.; Mokrij, E. N.; Skacko, V. P.; Lezjuk, V. V. No Title. DE Patent 3220555 C2, 1982.

(6)

Wolf, D.; Schmidt, R.; Block, U.; Schoenmakers, H.; Bott, K.; Kaibel, G. Production of Anhydrous or Substantially Anhydrous Formic Acid. U.S. Patent 4,326,073, 1982.

(7)

Zhang, J.; Zheng, H.; Li, M. X.; Zhou, M. Y.; Wu, M. J.; Feng, C. X.; Zhao, X. Y.; Wang, T.; Liu, L. Production Device with Yearly Output of 100,000 Tons of Formic Acid. 202968428U, 2012.

(8)

Huang, H. P.; Lee, M. J.; Lee, H. Y.; Chen, J. H. Formic Acid Manufacturing System and Method of Manufacturing Formic Acid. U.S. Patent 0123157 A1, 2012.

(9)

Saari, K.; Tirronen, E.; Vuori, A.; Lahtinen, M. Method for Preparing Formic Acid. 6429333 B1, 2002.

(10)

Kashani, S. Y.; Farhadi, F.; Baghalha, M. Simulation of Leonard-Kemira’s Formic Acid Production Process from CO. In 3rd Conference on Process Engineering Conference (oil, gas, refining and petrochemicals); 2014.

(11)

Novita, F. J.; Lee, H. Y.; Lee, M. Self-Heat Recuperative Dividing Wall Column for Enhancing the Energy Efficiency of the Reactive Distillation Process in the Formic Acid Production Process. Chem. Eng. Process. Process Intensif. 2015, 97, 144–152.

(12)

Novita, F. J.; Lee, H. Y.; Lee, M. Plantwide Design for High-Purity Formic Acid Reactive Distillation Process with Dividing Wall Column and External Heat Integration Arrangements. Korean J. Chem. Eng. 2018, 35, 1–15.

(13)

Sharma, S.; Patle, D. S.; Gadhamsetti, A. P.; Pandit, S.; Manca, D.; Nirmala, G. S. Intensification and Performance Assessment of the Formic Acid Production Process 34 ACS Paragon Plus Environment

Page 34 of 41

Page 35 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


through a Dividing Wall Reactive Distillation Column with Vapor Recompression. Chem. Eng. Process. Process Intensif. 2018, 123, 204–213. (14)

Wang, C. Study on Hydrolysis of Methyl Formate into Formic Acid in a Catalytic Distillation Column. J. Chem. Eng. Chinese Univ. 2006, 20 (6), 898–903.

(15)

Jogunola, O.; Salmi, T.; Eränen, K.; Mikkola, J. P. Qualitative Treatment of Catalytic Hydrolysis of Alkyl Formates. Appl. Catal. A Gen. 2010, 384, 36–44.

(16)

Jogunola, O.; Salmi, T.; Wärnå, J.; Mikkola, J. P. Kinetic Studies of Alkyl Formate Hydrolysis Using Formic Acid as a Catalyst. J. Chem. Technol. Biotechnol. 2012, 87, 286– 293.

(17)

Jogunola, O.; Salmi, T.; Mikkola, J. P. Screening of Potential Complexing Agents in Methyl Formate Hydrolysis. J. Mol. Liq. 2014, 196, 334–339.

(18)

Sahin, S.; Bayazit, S. S.; Bilgin, M.; Inci, I. Investigation of Formic Acid Separation from Aqueous Solution by Reactive Extraction : Effects of Extractant and Diluent. J. Chem. Eng. Data 2010, 55 (4), 1519–1522.

(19)

da Cunha, S.; Rangaiah, G. P.; Hidajat, K. Design, Optimization and Retrofit of the Formic Acid Process I: Base Case Design and Dividing-Wall Column Retrofit. Ind. Eng. Chem. Res. 2018, 57, 9554–9570.

(20)

Towler, G.; Sinnot, R. Economic Evaluation of Projects; Butterworth-Heinemann: Oxford, 2013.

(21)

Turton, R.; Bailie, R. C.; Whiting, W. B.; Shaeiwitz, J. A.; Bhattacharyya, D. Analysis, Synthesis, and Design of Chemical Processes, 4th ed.; Pearson: Upper Sadle River, 2013.

(22)

Sharma, M. M.; Mahajani, S. M. Industrial Applications of Reactive Distillation. In Reactive Distillation: Status and Future Directions; Sundmacher, K., Kienle, A., Eds.; Wiley: Morlenbach, 2002.

(23)

Taylor, R.; Krishna, R. Modelling Reactive Distillation. Chem. Eng. Sci. 2000, 55, 5183– 5229.

(24)

Krishna, R. Reactive Separations: More Ways to Skin a Cat. Chem. Eng. Sci. 2002, 57, 35 ACS Paragon Plus Environment


1491–1504. (25)

Kister, H. Z.; Mathias, P. M.; Steinmeyer, D. E.; Penney, W. R.; Crocker, B. B.; Fair, J. R. Equipment for Distillation, Gas Absorption, Phase Dispertion, and Phase Separation. In Perry’s Chemical Engineer’s Handbook; Green, D. W., Perry, R. H., Eds.; McGraw-Hill: New York, 2008.

(26)

Yildirim, Ö.; Kiss, A. A.; Kenig, E. Y. Dividing Wall Columns in Chemical Process Industry: A Review on Current Activities. Sep. Purif. Technol. 2011, 80, 403–417.

(27)

Asprion, N.; Kaibel, G. Dividing Wall Columns: Fundamentals and Recent Advances. Chem. Eng. Process. Process Intensif. 2010, 49, 139–146.

(28)

Staak, D.; Grützner, T.; Schwegler, B.; Roederer, D. Dividing Wall Column for Industrial Multi Purpose Use. Chem. Eng. Process. Process Intensif. 2014, 75, 48–57.

(29)

Wang, X.; Wang, H.; Chen, J.; Zheng, W.; Qiu, T. High Conversion of Methyl Acetate Hydrolysis in a Reactive Dividing Wall Column by Weakening the Self-Catalyzed Esterification Reaction. Ind. Eng. Chem. Res. 2017, 56, 9177–9187.

(30)

Wankat, P. C. Separation Process Engineering, 2nd ed.; Pearson: Westford, 2007.

(31)

Sulzer Chemtech. Structured Packings for Distillation, Absorption and Reactive Distillation; Winterthur.

(32)

Branan, C. Rules of Thumb for Chemical Engineers, 4th ed.; GPP: Burlington, 2005.

(33)

Maroulis, Z. B.; Saravacos, G. D. Food Process Design, 1st ed.; CRC Press: New York, 2003.

(34)

Sigma-Aldrich. Handbook of Fine Chemicals; 2011.

(35)

Seider, W. D.; Seader, J. D.; Lewin, D. R. Cost Accounting and Capital Cost Estimation, 2nd ed.; Wiley: New York, 2004.

(36)

Yau, T. L. Corrosion of Zirconium and Its Alloys. In Shreir’s Corrosion; Richardson, T. J. A., Ed.; Elsevier B.V., 2010.

(37)

Couper, J. R.; Penney, W. R.; Fair, J. R.; Walas, S. M. Chemical Process Equipment, 3rd 36 ACS Paragon Plus Environment

Page 36 of 41

Page 37 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


ed.; Butterworth-Heinemann: Waltham, 2012. (38)

Fisher Scientific. Dowex 50WX8 50-100 (H) www.fishersci.com (accessed Sep 17, 2017).

(39)

Götze, L.; Bailer, O. Reactive Distillation with KATAPAK-S. Sulzer Technical Review. 1999, pp 29–31.

(40)

Sharma, S.; Rangaiah, G. P.; Marechal, F. Integrated Multi-Objective Differential Evolution and Its Application to Amine Absorption Process for Natural Gas Sweetening. In Differential Evolution in Chemical Engineering; Rangaiah, G. P., Sharma, S., Eds.; World Scientific: Singapore, 2017.

(41)

Brennan, D. Life-Cycle Evaluation of Chemical Processing Plants. In Environmentally Conscious Materials and Chemicals Processing; Kutz, M., Ed.; Wiley, 2007.

(42)

Lin, Y.-D.; Chen, J.-H.; Cheng, J.-K.; Huang, H. P.; Yu, C.-C. Process Alternatives for Methyl Acetate Conversion Using Reactive Distllation. 1. Hydrolysis. Chem. Eng. Sci. 2008, 63, 1668–1682.

(43)

Fisher Scientific. Amberlite IR-120 H www.fishersci.com (accessed Sep 17, 2017).

(44)

Niu, M. W.; Rangaiah, G. P. Process Retrofitting via Intensification: A Heuristic Methodology and Its Application to Isopropyl Alcohol Process. Ind. Eng. Chem. Res. 2016, 55, 3614–3629.

(45)

Lorenz, H. M.; Otte, D.; Staak, D.; Repke, J. U. Increasing the Efficiency of Existing Complex Downstream Processes by Plant-Wide Retrofit. Chem. Eng. Res. Des. 2015, 99, 199–207.

(46)

Sander, S.; Flisch, C.; Geissler, E.; Schoenmakers, H.; Ryll, O.; Hasse, H. Methyl Acetate Hydrolysis in a Reactive Divided Wall Column. Chem. Eng. Res. Des. 2007, 85, 149–154.



Page 38 of 41

Appendix A Tables A1 and A2 give the stream data for inlet and outlet streams of RD and RDWC columns in Figures 3 and 4, respectively. Flowrate, composition and temperature of other process streams are almost the same as those reported for the optimal base case process.19 Table A1 – Data for inlet and outlet streams of RD column in Figure 3 Vapor

Flow rate

Mole fractions of CO, MF, MA, Water

fraction

(kmol/h)

and FA

5.37

0.000

93.88

0.001

0.746

0.253

0.000

0.000

88.8

5.37

0.000

418.23

0.000

0.000

0.001

0.998

0.001

21

63.0

5.37

1.000

0.20

0.453

0.508

0.008

0.031

0.000

22

138.1

5.37

0.000

511.9

0.000

0.001

0.183

0.680

0.136

Stream

T (°C)

P (bar)

15

17.0

18

Table A2 – Data for inlet and outlet streams of RDWC column in Figure 4 Vapor

Flow rate

Mole fractions of CO, MF, MA, Water

fraction

(kmol/h)

and FA

5.67

0.000

93.85

0.001

0.746

0.253

0.000

0.000

88.8

5.67

0.000

419.81

0.000

0.000

0.001

0.998

0.001

21

64.8

5.67

1.000

0.20

0.453

0.506

0.008

0.032

0.000

23

165.3

5.67

0.000

419.80

0.000

0.000

0.001

0.833

0.166

29

115.5

5.67

0.000

93.66

0.000

0.005

0.995

0.000

0.000

Stream

T (°C)

P (bar)

15

55.0

18

Appendix B Figures B1 and B2 present HEN for the RD and RDWC retrofitted processes in Figures 3 and 4, respectively. They show the configuration, which is the same for both preliminary and optimal designs. In Figures B1 and B2, stream names are on the left side, and the values above each arrow are the stream temperatures (in °C). When two numbers appear next to each other (for example, 125.2/118.1 in stream 23 in Figure B1), the first number refers to the value for the 38 ACS Paragon Plus Environment

Page 39 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


preliminary design and the second refers to that for the optimal solution. REBi refers to the bottom stream going to the reboiler in column ‘i’, whereas CCi refers to the top stream going to the condenser in column ‘i'. CRDWC2 refers to the condenser of the RDWC unit at the top of section 2 (in Figure 5). Only streams participating in heat integration are displayed in Figures B1 and B2, and streams being heated/cooled entirely by utilities are omitted. The minimum temperature approach considered for the HEN was 10°C. Besides, streams being integrated are separated from each other by at most 4 units. This ensures that the heat integration is feasible since the streams will be close to each other in the plant.

Figure B1 – HEN for the RD retrofit (preliminary and optimal solutions)



Figure B2 – HEN for the RDWC retrofit (preliminary and optimal solutions)

Table of Contents Graphic


Page 40 of 41

Page 41 of 41 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


For Table of Contents Only


Design, Optimization and Retrofit of the Formic Acid Process II

Recommend Documents