Acetic Acid

13 ACETIC ACID CHAPTER OUTLINE 13.1 Introduction 483 13.2 Process Chemistry 485 13.2.1 Rhodium Catalyzed Carbonylation 485 13.2.2 Iridium Catalyzed Carbonylation 488 13.2.3 Heterogeneous Catalyst 489 13.3 Chemical Equilibrium 490 13.4 Kinetic Aspects 491 13.5 Physical Properties 493 13.5.1 Acetic Acid 493 13.5.2 Components 495 13.6 Health, Safety and Environment 495 13.7 Input/Output Analysis 496 13.8 Reaction Section 496 13.8.1 Types of Reactors 496 13.8.2 Simulation of the Reaction System 497 13.8.3 Reactor Sizing 500 13.9 Separation Section 501 13.10 Process Integration 506 13.10.1 Integration of Reaction and Separation Sections 13.10.2 Heat Pump Assisted Distillation 507 13.10.3 Heterogeneous Catalyst Process 509 13.11 Economic Evaluation 509 13.12 Sustainability Analysis 513 13.13 Conclusions 518 References 519

506

13.1 Introduction Acetic acid is a major chemical product with over 16 million tons as global consumption in 2018, BP Chemicals being the largest supplier. The overall acetic acid production is shared between the manufacturing of vinyl acetate monomer (VAM) 33%, acetic anhydride 18%, monochloroacetic acid (MCA) 17%, acetate esters Applications in Design and Simulation of Sustainable Chemical Processes. https://doi.org/10.1016/B978-0-444-63876-2.00013-9 Copyright © 2019 Elsevier B.V. All rights reserved.

483

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Chapter 13 ACETIC ACID

17%, and terephthalic acid 17% (Le Berre et al., 2012). The development is driven by the growth of various consumer products as coatings, adhesives, resins, varnishes, vinyl alcohols, polymers for packaging industry, cellulosic fibers, agrochemicals, etc. In terms of value and volume, AsiaePacific has today the highest share, followed by North-America, Europe, Middle East and Africa. The biggest importing countries of acetic acid are Belgium, India, Mexico, Germany and Brazil. The manufacturing of acetic acid by the carbonylation of methanol takes the largest part of the market, about 75%. This share is in continuous growth. The trend is favourable from sustainability view point, as it will be demonstrated in this case study. Beside methanol, the synthesis may start from di-methyl ether (DME), and in the end from materials leading to syngas, as biogas and biomass. For this reason, the acetic acid may be a 100% green product, and on this basis promoting other bio-chemicals. For example, if both acetic acid and ethylene are manufactured from methanol then the vinyl acetate is a bio-product, as well as the valuable acrylic paints produced from it. In this way a variety of products issued from a syngas bio-refinery can replace traditional petrochemicals. Another process based on renewable materials, known from old times, is the fermentation of sugar residues solutions, as spoiled wine, driven by suitable bacteria (Clostridium thermoaceticum) at pH below 5. Despite the interest, the economic efficiency is penalized by high energy consumption in the separation (concentration) and purification operations. The result is that the fermentation process is limited today at only 10% of the market. Other traditional processes are based on petrochemical feedstock, as liquid-phase oxidation (n-butane and naphta), or vapor phase-oxidation (ethane and ethene), but these are going to be replaced by methanol carbonylation (Le Berre et al., 2012). A novel route to produce acetic acid reported recently involves the hydrocarboxylation of methanol with CO2 and H2 (Qian et al., 2016). The reaction can be efficiently catalyzed by RueRh bimetallic catalyst using imidazole as ligand and LiI as promoter in 1,3-dimethyl-2-imidazolidinone (DMI) solvent with outstanding reaction results. This method opens a road to fix CO2 into bulk chemicals using easily available and cheap feedstocks. In the following section we will tackle the conceptual design of a methanol carbonylation process.

Chapter 13 ACETIC ACID

The process can be described by the following overall stoichiometric equation: CH3eOH þ CO / CH3eCOOH with DHr,298 ¼ 135.6 kJ/mol

(13.1)

The remarkable feature of the above reaction is a theoretical carbon yield of 100%. Indeed, the rhodium-based catalysts used today are highly selective with a carbon yield close to 100%. Methanol carbonylation reaction is moderately exothermic and thus thermodynamically favored by lower temperature. Since the reaction takes place in liquid phase, the pressure should be high enough to ensure suitable CO concentration. The first commercial methanol carbonylation plant was builtup by BASF in Germany in the 1960s based on a cobalt/hydrogen iodide homogeneous catalyst. The reaction conditions were difficult, at high pressure over 500 bar and temperature between 150 and 200 C. Because of highly corrosive chemicals, the equipment was built from very expensive materials, as titanium and Hasteloy. The overall yield was about 90%, with the formation of numerous by-products. In 1970-decade the process was significantly improved by Monsanto in USA, which discovered a highly active and selective Rhodium-based catalyst. The temperature range was similar, but the operating pressure dropped from 30 to 60 bar, while yield increased to 99%. Further improvement took place in 1996 when BP Chemicals from UK overtook the Monsanto licence. BP invented a new catalytic system based on Iridium (Cativa process), still expensive but cheaper than the one based on rhodium. The optimal temperature remains in the range of 180e220 C, but the pressure drops further to 20e40 bar. The yield is over 99% with less impurities and simplified separation.

13.2 Process Chemistry Although the overall stoichiometry is simple, the chemistry is complex when examining the catalytic reaction mechanism. This aspect has profound implications on conceptual process design.

13.2.1 Rhodium Catalyzed Carbonylation The first presentation regards the reaction mechanism of the original Monsanto process, as depicted in Fig. 13.1. The reaction takes place in liquid phase, methyl acetate being used as solvent,

485

486

Chapter 13 ACETIC ACID

I

I

–

CO

A

Rh

CO

I MeCOI Me

I

H2O I

RhIII I

CO I

RhIII

I HI

MeCO2H

MeOH Me

Scheme I The reaction cycle for the Monsanto rhodium-catalysed carbonylation of methanol to acetic acid

–

Me

CO

OC D

Mel –

CO B CO

I

–

CO

I RhIII

CO

I CO

I

C

Figure 13.1 Reaction mechanism by the original Monsanto methanol carbonylation process From Jones, J.H., 2000. The Cativa process for the manufacture of acetic acid. Platinum Metals Review, 44, 94e105 with permission.

in which the catalytic complex is soluble. The reaction mechanism consists of six elementary steps. A prerequisite is the formation of CH3I and H2O species, further implied in the catalytic cycle but recovered and recycled at the end. The initial metallic complex is [Rh(CO)2I2] reacts in the first step (A) with CH3I giving the complex [Rh(CO)2MeI3]. This undergoes in the second step (B) a molecular rearrangement leading to the formation of a complex that includes now an acyl group [Rh(CO) (COeMe)I3]. In the last third step (C) the addition of a CO molecule gives the unstable intermediate [Rh(CO)2(COeMe)I3]. In the fourth step (D) the expulsion of the acyl iodide CH3COI takes place while rebuilding the initial catalytic complex. The last step is the hydrolysis with water producing acetic acid and HI that re-joins the cycle. The elementary reactions may be outlined as follows (Moulijn et al., 2013): CH3eOH þ HI 4 CH3I þ H2O CH3I þ CO / CH3COI

DHr,298 ¼ 53.1 kJ/mol (13.2) DHr,298 ¼ 66.9 kJ/mol (13.3)

CH3COI þ H2O / CH3COOH þ HI DHr,298 ¼ 15.6 kJ/mol

(13.4)

Chapter 13 ACETIC ACID

The presence of a controlled amount of water is essential in the reaction. Water forms directly by the esterification reaction between methanol and acetic acid, but its presence comes mainly by recycling from the separation section. When the water content is higher than 8 wt.%., the rate determining step in the first stage (A) is the addition of methyl iodide to the rhodium center. RCH3COOH ¼ k  [Rh]  [CH3I]

(13.5)

The reaction rate is then essentially first order in both catalyst and methyl iodide concentrations, but independent of the reactant concentrations. Therefore, high conversions can be obtained even in a continuous stirred-tank reactor of relatively small volume (Moulijn et al., 2013). If the water content is less than 8 wt.%, the rate determining step is (D), the reductive elimination of the acyl species. Keeping the integrity of the catalyst is essential. The principal cause of catalyst loss by precipitation is the reaction of rhodiumacyl species D with hydrogen iodide that leads to acetaldehyde and the complex [Rh(CO)I4]-. The latter may cause a catalyst loss by precipitation of the inactive and insoluble rhodium triiodide, when the CO concentration is insufficient. [Rh(CO) (COeCH3)I3] þ HI / CH3eCHO þ [Rh(CO)I4] / RhI3 þI þ CO (13.6) Thus, a minimum partial pressure of CO is required for preserving the catalyst activity. The Monsanto process has been improved by Celanese by means of acetic acid optimization (AO) technology which incorporates catalyst promoters based on alkali iodides. This technology increases significantly the reactants’ efficiencies at lower water concentrations. However, the presence of iodide species may rise problems for their removal from the final product, when the purity specifications of acetic acid are in the range of ppb’s, as for example required by the vinyl acetate monomer (VAM) production process. Although rhodium-catalysed carbonylation of methanol is highly selective, it suffers from some side reactions. For example, rhodium also catalyses the water gas shift reaction: CO þ H2O / CO2 þ H2

(13.7)

Further methane formation may occur by reaction of hydrogen with methanol: CH3OH þ H2 / CH4 þ H2O

(13.8)

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Chapter 13 ACETIC ACID

These side reactions may reduce the amount CO available for reaction down to a point where the catalyst integrity is in danger. Another unwanted reaction is the formation of acetaldehyde. The reaction with hydrogen gives ethanol, and further by carbonylation the propionic acid, which is the main heavy impurity. H2

CO

CH3eCHO ƒ! CH3eCH2eOH ƒ! CH3eCH2eCOOH (13.9) High-boilers can be formed by aldol condensation reactions. Therefore, in the Monsanto technology the separation system needs typically three columns and substantial energy consumption.

13.2.2 Iridium Catalyzed Carbonylation The use of Iridium as coordinating metal is justified by its price which is 17 times lower than rhodium (Jones, 2000). But as catalyst, Iridium brings also significant modification in the reaction mechanism (see Fig. 13.2) with strong impact on the process performance (see Table 13.1). The oxidative addition of methyl – I I

CH3COI

I–

Ir

CO CO

Fast CH3I

H2O

COCH3 I

I

Ir I

–

CH3

CO CO

I HI

CH3CO2H

CO

Ir

CO

CH3OH Rate-determining step I–

CH3 I

Ir

I

CO CO

CO

CO

Figure 13.2 Reaction mechanism by iridium catalyzed carbonylation in the Cativa process From Jones, J.H., 2000. The Cativa process for the manufacture of acetic acid. Platinum Metals Review, 44, 94e105. with permission.

Chapter 13 ACETIC ACID

489

Table 13.1 Comparison of Methanol Carbonylation by Rhodium and Iridium Catalysts (Jones, 2000). Factor

Rhodium Catalyst

Iridium Catalyst

Water

1-st order below 8 wt.% Independent above 8 wt.% Independent above 1 wt.% 1st order

Increases with water up to about 5 wt.%, then decreases Increases with increasing methyl acetate Increases with methyl iodide up to 6 wt.%, then independent Increases with increasing CO partial pressure. As the CO partial pressure falls below 8 bar the rate decreases more rapidly 1 s t order, effect tails off at high catalyst concentrations Increases with increasing promoter, effect tails off at higher concentrations As the corrosion metals increase in concentration, the rate decreases

Methyl acetate Methyl iodide CO partial pressure

A minimum CO partial pressure is required

Coordination metal Promoter

1st order Non applicable

Corrosion metals

Independent

iodide to the iridium center is about 150 times faster compared to rhodium, and thus does not determine the reaction rate. The slowest step becomes the methyl migration to form the acyl complex. This step involves substitution of iodide center with carbon monoxide. Thus, the reaction rate dependence is totally different from that for the rhodium-catalysed process: RCH3COOH ¼ k  [Ir][CO] /[I]

(13.10)

The inverse dependence on the ionic iodide concentration implies that removing it increases the reaction rate. This operation may be done by “promoters”. These may be simple iodide complexes of zinc, cadmium, mercury, gallium and indium, as well as carbonyl-iodide complexes of tungsten, rhenium, ruthenium and osmium. The use of promoters can increase greatly the reaction rate, up to 20 times by molar ratio Ir/Ru of 5. An additional effect is protection of catalyst against deactivation.

13.2.3 Heterogeneous Catalyst A significant progress was achieved by the Japanese company Chiyoda that developed a process using a heterogeneous rhodium catalyst under the commercial name CT-ACETICA

490

Chapter 13 ACETIC ACID

I

–

+

CH3 N

[Rh(CO)2I2]

–

–

+

CH3 N + I

+ [Rh(CO)2I2] Active Form of Rhodium at High CO Partial Pressure

Polyvinylpyridine

–

Rh-Complexed Resin

Figure 13.3 Rhodium immobilisation on vynilpyridine resin. Yoneda, N., Takeshi M., Weiszmann, J., Spehlmann, B., 1999. The Chiyoda Acetica process: a novel acetic acid technologies, Surface Science and Catalysis, 121, 93e98, with permission.

process. As shown in Fig. 13.3, the rhodium complex catalyst is immobilised on a vinyl-pyridine resin. The catalyst remains in the reaction space and its integrity is preserved. The amount of catalyst in the reactor can be raised to enhance the space-time yield (STY) of the carbonylation reaction. The loss of rhodium catalyst in the distillation section can be reduced to zero, as this remains in the reactor. Since employing lower amount of water (under 2%), the energy requirement for separation of water from the acetic acid product is drastically reduced. From reaction engineering viewpoint the process adopts a bubble-column loop type reactor.

13.3 Chemical Equilibrium The thermodynamic analysis gives useful insights on process feasibility. Table 13.2 presents the results of chemical equilibrium calculations using Gibbs free minimisation reactor model Aspen Plus 9.0. The reaction takes place at 20 bar and 200 C starting with different composition. In the first case the reactants are near stoichiometry with 10% excess of CO. The result is full conversion of methanol, the CO excess leaving unreacted. The full conversion really happens, the catalyst being very active in these conditions. If methanol is taken in excess then the CO conversion is complete, but beside acetic acid, smaller amounts of methyl acetate and water are formed. In the third case an amount of methyl acetate is initially present, but this does not affect the carbonylation reaction. But when ester, acetic acid and water are present (due to recycles), then the mixture’s composition

Chapter 13 ACETIC ACID

491

Table 13.2 Equilibrium Composition of a Synthesis Mixture at 200C and 20 bar. Excess CO

Excess MeOH

Methyl Acetate

Methyl Acetate D water

Initial Equilibrium Initial Equilibrium Initial Equilibrium Initial Equilibrium Methanol CO Acetic acid Me acetate Water

1 1.2 0 0 0

0.000 0.201 0.998 0.001 0.001

1.2 1 0 0 0

0.010 0.000 0.810 0.190 0.190

1 1.2 0 0.5 0

0.000 0.200 1.000 0.500 0.000

is affected by the equilibrium of esterification and hydrolysis reactions. When superposing the effect of catalyst, it comes out that the prediction of the composition by thermodynamic and kinetic modeling is a complex matter (Reza et al., 2012).

13.4 Kinetic Aspects In this section we review briefly key kinetic aspects of methanol carbonylation with impact on process design. The reaction rate by methanol carbonylation exhibits a complex dependence on various factors, as concentrations of reactants, catalytic complex, co-catalyst and promotors, but most of all on the water amount present in the reaction mixture. Note that the concentrations of species are also constrained by phase and chemical equilibria. Fig. 13.4 displays the reaction rate versus water concentration for the original Monsanto (Rh catalyst) process compared with Cativa process (Ir/Rh catalyst). In the first case there is a strong rise of the reaction rate up to 8 %wt., where this becomes constant and independent of both CO and methanol (zero order reaction) as shown by Eq. (13.5). Note that a large amount of water implies higher energy consumption for water removal and recycling by distillation, since the separation acetic acid/water is very difficult. In the case of an Iridium catalyst using Ruthenium as promotor (1:2 M), one may note a strong process enhancement in general. A rising path is noted up with a peak of 40 kmol/m3/h at 4 %wt. water, which is followed by a decline.

1 1.1 0.2 1 0.5

0.046 0.000 1.446 0.854 0.354

492

Chapter 13 ACETIC ACID

Figure 13.4 Reaction rate versus water concentration for Monsanto and Cativa processes. With permission, Jones, J.H., 2000. The CativaTM process for the manufacture of acetic acid. Platinum Metals Review, 44, 94e105.

The literature search retrieves some kinetic models for determining the reaction rate as function of other chemical species imbedding phase and chemical equilibrium, as a simplified model (Nowicki et al., 1992), or a more complex one (Reza et al., 2012). However, the applicability of these models for the industrial reactor design was not tested. It may be concluded that the Iridium based catalyst bring notable advantages, namely higher stability over a wide range of process conditions. The process rate is less dependent on methyl-iodide concentration, but it depends quite significantly on CO partial pressure, the concentration of catalyst metal complex and of the promoter amount, as well as on methyl acetate amount used as solvent. However, the formation of impurities is reduced. The result is a global selectivity over 99% with respect to methanol. This complex dependence may be exploited in order to achieve the best productivity with low energetic consumption. Table 13.3 present results obtained in industrial conditions and continous operation (Sunley and Watson, 2000). Increased amount of Iridium and Ruthenium lead to a notable enhancement effect (experiments 1 to 3). Lower water amount gives higher rate, but higher impurity content too. Lower Iridium level may be compensated by more CH3I co-catalyst (experiments 5e6). In addition, the lowest amount of propionic acid (PrAc) is obtained. The highest rate is obtained by a simutanous optimisation of variables (experiment 7). The use of heterogeneous catalyst removes the limitations of homogeneous catalysts. Much higher catalyst concentrations may be realized, and higher process rate can be achieved without

Chapter 13 ACETIC ACID

493

Table 13.3 Reaction Rate and Product Distribution in Industrial Conditions by Cativa Process. No.

Ir, ppm

Ru, ppm

T, C

CH3I %wt.

MeAcI %wt.

H 2O %wt.

PCO bar

Rate mol/Lh

PrAc ppm

1 2 3 4 5 6 7

1600 3240 1780 1130 840 870 1170

e e 2300 1640 2330 2870 1720

190 190.6 190.0 192.5 189.0 189.0 192.6

5.3 3.8 5.1 7.0 10.00 10.3 6.7

15.8 15.7 15.1 16.2 21.9 22.3 16.4

7.0 7.1 7.2 2.8 4.6 5.7 4.1

11 10 8.9 8.4 8.3 8.3 8.0

11 17.0 17.1 19.2 19.8 19.8 22.5

530 930 600 950 360 290 670

using promoters (Yoneda and Takeshi, 1999). Immobilization reduces the loss of expensive Rhodium metal, while low water concentration has as result lower energetic consumptions.

13.5 Physical Properties 13.5.1 Acetic Acid The acetic acid is a clear colourless liquid with an irritating vinegar odor detectable from 1 ppm, toxic and aggressive. The melting point is at 16.73 C and normal boiling point at 117.9 C. Table 13.4 presents the vapor pressure of acetic acid as function of temperature. Acetic acid is sold commercially as glacial acetic acid with less than 1% water and over 98% concentration, other impurities resulting from the manufacturing process. The density of pure acetic acid is 1.0491 at 20 C and 0.9599 at 100 C. The acetic acid is highly hygroscopic. The purity of water solutions can be related with their freezing point.

Table 13.4 Vapor Pressure of Acetic Acid as Function of Temperature. T/ C Pv/kPa

20 1.57

60 4.53

100 55.33

140 184.1

180 509

220 1170

280 3280

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Chapter 13 ACETIC ACID

The mixture of acetic acid with water has very peculiar properties. Higher density is observed between 67% and 87%, which proves that the acetic acid develops molecular associations. Measuring the vapor density of pure acetic acid demonstrates the formation of dimers. The liquid phase shows also the existence of monomer, dimers, and even higher oligomers. The increase in concentration favors the formation of dimers, while the increase of temperature the formation of monomers. As shown in Fig. 13.5, the T-xy diagram of the binary acetic acid/water at normal pressure exhibits a highly non-ideal behavior. There is a large variation in the relative volatility, which converge to one at water molar fractions over 90%. The existence of a “tangent-pinch” in a y-x diagram indicates that a large number of theoretical stages is necessary to perform the separation. The alternative at conventional distillation is heterogeneous azeotropic distillation with entrainer. The modeling of vapour-liquid equilibrium (VLE) can be done by using the chemical association theory, namely the model Hayden-O’Connell for vapor phase combined with activity description for liquid phase. This model describes accurately the experimental data, as shown in Fig. 13.5. In any process involving acetic acid, the selected property method must consider the dimerization of acetic acid in the vapor phase. Hence either the Hayden-O’Connell, Nothnagel or Marek-Standart correction should be used to account for the non-ideality of the vapor phase. Hayden-O’Connell is generally used for non-ideal vapor containing polar, solvating or associating components, at low to moderate pressures.

Figure 13.5 VLE for the system water-acetic acid at 1.013 bar.

Chapter 13 ACETIC ACID

13.5.2 Components The key components in separations are: acetic acid, methanol, carbon oxide, water, methyl acetate, methyl iodide. Fig. 13.6 displays the variation of vapor pressure with the temperature. Methyl iodide is the most volatile (nbp 42.4 C), followed by methyl acetate (56.9 C), methanol (64.7 C), water (100.0 C), acetic acid (117.9 C) and propionic acid (141.2 C). The first three components may be considered as condensable lights. The separation of water and acetic acid is difficult due to a tangent pinch in the T-x,y diagram. The final purification of acid acetic from propionic acid should also be feasible.

13.6 Health, Safety and Environment The conceptual design of an acetic acid plant by methanol carbonylation must consider serious hazards caused by the toxicity and corrosion action of different chemicals. Acetic acid itself is toxic by inhalation, ingestion and skin contact, flammable and harmful to aquatic life. The personnel should avoid inhalation exposure above 25 mg/m3. Environmental problem is caused by the waste water treatment of streams containing acetic acid, formic acid and other organics. These chemicals must be neutralized and then degraded biologically. The acetic acid is highly reactive and incompatible with many materials. There are strict regulations regarding storage and transportation. Suitable materials are containers lined with stainless steel, glass, or polyethylene. Aluminum forms a protective coating when exposed to acetic acid. 9 8

P-vapour /[bar]

7

Methanol Acetic acid Methyl acetate Water Methyl iodide Propionic acid

6 5 4 3 2 1 0

0

20

40

60

80

100

120

Temperature / [°C]

Figure 13.6 Vapor pressure of key components versus temperature.

495

496

Chapter 13 ACETIC ACID

The chemical equipment makes use of special materials of construction, such as Hastelloy, modified AISI 316 and 321 stainless steels. Temperature higher than 200 C and the presence of some contaminants such as halides can increase the corrosion rate. Other components encountered in the acetic acid manufacturing are also harmful for humans and environment, namely methyl iodide (see sustainability metrics). More information about toxicity and hazards caused by chemicals can be found in the public available MSDS datasheets, as well as on the website of the PubChem organisation.

13.7 Input/Output Analysis The plant capacity considered here is 200,000 metric tons per year acetic acid corresponding to a production of 25,000 kg/h or 416.3 kmol/h. The stoichiometric amounts of raw materials are 416.3 kmol/h of each methanol and CO, or 13,339 and 11,661 kg/h, respectively. A preliminary plant material balance may be performed based on yield data, which are estimated as 98% for methanol, and 95% for CO. Lower yield for CO may be explained by the secondary reactions presented previously leading to CO2, H2 and CH4.

13.8 Reaction Section 13.8.1 Types of Reactors Fig. 13.7 depicts the type of industrial reactors employed by methanol carbonylation. Processes based on homogeneous catalyst make use of CSTR-type reactors. In the Monsanto-Celanese technology the reactor is a mechanically agitated vessel capable of working at pressure in the range of 20e40 bar and temperature 180e250 C. The reactor is provided with external heat exchanger for steam generation (SG). In the Cativa process, the mixing is achieved by gas injection under pressure that makes mechanical agitation un-necessary. The reactor design in Chiyoda process using heterogeneous catalyst is different. A slurry gas-liquid tower reactor was developed. The gas is injected at bottom through an efficient distribution device as very small bubbles such to saturate the liquid phase in CO. An external loop for liquid-phase circulation is created by the density difference caused by the CO consumption. The reaction takes place in the riser, while the downcomer ensures catalyst recycle and heat transfer through an external exchanger.

Chapter 13 ACETIC ACID

Offgas Product

Gaz venƟng

Catalyst recycle

Processing

Cooler SG

Recycle MeOH CO Catalyst recycle

CO MeOH

Recycle

Figure 13.7 Chemical reactors for methanol carbonylation (left Cativa process, right CT-Acetica process).

13.8.2 Simulation of the Reaction System In this section we tackle the design issues regarding the engineering of the chemical reaction. The chemistry and kinetics are complex matters, as explained before. Fortunately, the reaction is fast and achieves practically total conversion with excellent yield. In the workable regime, the reaction rate can be mastered mainly by the catalyst amount, including co-catalyst and promoters, since the reaction is zero-order with respect to the concentration of reactive species. Thus, a first issue is recycling all the components of the catalyst recipe that may be carried out to the separation section. This is valid for homogeneous catalyst processes, but not for heterogeneous catalyst process. A second issue is managing the large amount of heat developed by reaction, which is 135.6 kJ/mol. For a production rate of 416.3 kmol/h the amount of developed energy corresponds to a large duty of 15.7 MW. A suitable reaction device is CSTR with external cooling loop, as pictured in Fig. 13.7 left-hand. Fig. 13.8 presents the flowsheet simulated with Aspen Plus 9.0. This comprises the reactor RSTOIC, which models the reaction, followed by a flash unit introduced for describing the VLE. The gas stream is condensed at 5 C, the liquid being recycled to the reactor, while the gas leaves as vent stream. The liquid reaction mixture is split further in a process stream and recycle stream that enters the cooling loop. The process stream is depressurised in a “flasher” unit in view of separating the catalyst. The catalyst is recycled as

497

498 Chapter 13 ACETIC ACID

Temperature (C) Pressure (bar)

HEATER

Figure 13.8 Simulation flowsheet of the reaction section.

Q

Duty (MW)

W

Power(kW)

Chapter 13 ACETIC ACID

bottom stream to the reactor. A supplementary amount of heat may be necessary to ensure full vaporisation of the mixture. Alternatively, the liquid part may be sent in the bottom part of the next distillation column for lights’ separation. The reactor model is based on stoichiometry. The main reaction 15.1 has a fractional conversion of 0.98. The following secondary reactions describe the formation of by-products, and impurities: • Formation of gaseous impurities, relations Eqs. 13.7 and 13.8, lumped in the reaction CO þ CH3OH ¼ CO2 þ CH4

(13.11)

The molar extent of reaction 2 kmol/h. • Esterification of acetic acid with methanol: CH3eCOOH þ CH3eOH ¼ CH3eCOOCH3 þH2O (13.12) Molar extent of reaction 2 kmol/h • Formation of propionic acid by the relation: CH3eCOOH þ CH3eOH ¼ CH3eCH2eCOOH þ H2O (13.13) Molar extent of reaction 4 kmol/h. Note the formal analogy of the stoichiometric Eqs. (13.12) and (13.13), although the true reaction should pass by acetaldehyde formation. The merit of the above relations is preserving the atomic balance and offering a realistic description of the reaction mixture involved further in separations. In addition, another input stream considers other species implied in reactions and accumulated by recycles, as methyl iodide and methyl acetate, as well as the methyl acetate used as solvent. The significant amount of heat developed by reaction can be recovered as useful energy. In this work we select a cooling procedure that includes a steam generator and a cooler, as shown in Fig. 13.8. This solution seems to be used in industry, as described in a recent patent from Celanese corporation (Zinobile, 2014). In view of maximising the steam amount, the reaction temperature is set at 220 C, which is also indicated in the same patent. The reactor pressure should be sufficiently high to ensure liquid phase reaction, in general between 20 and 50 bar. The reactant inputs are methanol of 424.8 and CO 438.2 kmol/h, or 13,612 and 12,274 kg/h, respectively, according to preliminary material balance formulated before. Methanol is preheated at 100 C, while the CO stream is compressed at 35 bar in a four-stage compressor with intermediate cooling at 50 C

499

500

Chapter 13 ACETIC ACID

except the last adiabatic stage where the temperature rises at 152 C. In addition, a recycle stream from separations is considered containing acetic acid (1000), methyl acetate (2000), methyl iodide (2000), and water (1500), all components in kg/h. The metal catalyst is not introduced explicitly in simulation because modeling constraints. The selection of the thermodynamic model is a key issue since it should ensure accuracy of both phase equilibrium and energy balance (Dimian et al., 2014). A suitable thermodynamic model is Wilson Hayden-O’Connell with Henry components. Note that using an equation-of-state model, such as SR-Polar (applicable in principle for polar species) gives large errors. The simulation reports also a reaction enthalpy in standard conditions (25 C, 1 atm and liquid phase) of 124.8 kJ/mol, 8% less than the above cited value but satisfactory for the conceptual design purpose. The goal in designing the reaction system is achieving constant temperature inside the reactor by manipulating the cooling capabilities. A powerful method is using an external cooling loop. The parameters of the cooling loop, as specified further, were tuned by a trial-and-error procedure. An intermediate temperature of 150 C is assumed for generating LP steam of 3.5 bar. The amount of reaction heat recovered as LP steam is in this case about 6.2 MW, which can be used for driving the reboilers of the separation section. The rest of cooling is performed by passing the hot stream through the heat exchanger COOLER. The final outlet temperature and the split fraction were tuned to bring the duty of the RSTOIC reactor close to zero. The results are the cooling temperature at 56 C and split fraction of 0.242, which results in a duty of the unit COOLER of 6.24 MW. The energy available in the hot effluent may be used for reactant preheating, as will be developed in the section devoted to process integration. Table 13.5 presents the streams leaving the reaction section. The stream VENT contains gaseous species together with traces of condensable species. The stream PROCESS is sent to the liquid separation section.

13.8.3 Reactor Sizing The reactor sizing on kinetic basis is difficult considering the complexity of the catalytic process. However, a shortcut calculation may be applied for getting the main design elements. The key assumption is that the reaction rate can be maintained at a constant value independent of the concentration of reactants

Chapter 13 ACETIC ACID

Table 13.5 Output Stream From the Reaction Section. Stream

Flow, Kg/h

CH3OH

CO

Acetic acid

Methyl Acetate

H2O

CH3I

Acid Propionic

CO2

CH4

Process Vent Total

32,046 340 32,386

4 0 4

235 312 547

25,662 0 25,662

2296 0 2296

1608 0 1608

1999 <1 2000

148 0 148

78 10 88

22 10 32

but ensuring only a constant catalyst composition, including rhodium/iridium complex, CH3I, water and promotors. This condition is fulfilled by the catalyst kept in reactor and by the amounts of species returned by recycles. Make-up of catalyst components may be provided in the case of losses. Monitoring continuously the catalyst amount is a requirement of the operating procedure. In this case study we consider the Cativa process for which a reaction rate R of 20 kmol/m3/h is achievable, as indicated in Table 13.3. The reactor effluent flowrate from simulation is 132,760 kg/h. The effluent density corresponding to composition from Table 13.3 is 655 kg/m3. It results a volumetric flow Qv ¼ 132,760/655 ¼ 202 m3/h. From the material balance the amount of acetic acid formed by reaction is: NA ¼ NMeOH  conversion ¼ 424.8  0.99 ¼ 420.55 kmol/h It follows that the reaction volume is VR¼NA/R ¼ 420.55/ 20 ¼ 21.03 m3. Accordingly, the reaction time is tR ¼ VR/Qv ¼ 21.03/202  3600 ¼ 373 s. A volume factor of 0.75 gives a reactor volume of 28 m3. For a cylindrical vessel with ellipsoidal heads the volume is VR ¼ p/4  D3  (H/Dþ0.166), which for H/D 1.4 leads to 2.3 m diameter and 3.2 m height.

13.9 Separation Section The strategy of separations starts by examining the state and composition of the reactor-outlet mixture (Dimian et al., 2014). The stream PROCESS that leaves the flasher as vapor at 1.6 bar

501

502

Chapter 13 ACETIC ACID

and 125 C contains acetic acid in large quantity, and much lower amounts of methyl acetate, methyl iodide and water (there are also small amounts of dissolved gases and entrained rhodium catalyst). Vapor pressure is the characteristic property for determining the separation sequence of components. Methyl iodide is the most volatile (nbp 42.4 C), followed by methyl acetate (56.9 C), methanol (64.7 C), water (100.0 C), acetic acid (117.9 C) and propionic acid (141.2 C). Accordingly, a direct sequence of three columns is appropriate (as shown in Fig. 13.9) although an intensified alternative using dividing-wall column (DWC) technology is also possible. The first column C-1 removes the lights, namely methyl iodide, methyl acetate, methanol and gaseous species as top distillate (vapor and liquid), while the bottom recovers completely the entrained rhodium catalyst as acetic acid solution. The acetic acid taken over as vapor side-stream is the feed for the next column C-2. Here the dehydration of acetic acid takes place, the distillate recovering completely water and lights. The vapor distillate from C-1 and C-2 enters column C-4 where the lights are recovered by absorption in methanol. A final cooling at 5 C removes the entrained methanol too. The condensate and the top distillates are recycled to the chemical reactor. In this way the catalyst composition is preserved. A makeup of methyl iodide might be necessary to compensate the losses in exit gaseous streams. The water inventory should also be balanced. The bottom stream containing the acetic acid is sent finally to the purification column C-3. High purity acetic acid is obtained as top product, while the heavy impurities are removed in the bottom. Note that the separations take place at top pressures close to atmospheric. The above sequence was simulated in Aspen Plus v9.0 using the VLE model Wilson-Hayden-O’Connell with Henry components. The sizing of internals is performed interactively by selecting different types, trays or packing. Tables 13.6 and 13.7 present selected streams and key results regarding the distillation columns. Column C-1 has 17 theoretical stages, with the vapor inlet stream location being close to reboiler, and it is equipped with FLEX-S trays. Note that the diameter is varied in three steps, 1.3/1.8/2.2 m, to accommodate the variation in vapor and liquid flows to an efficient hydraulic regime between 70% and 80% from flooding. The top condenser works at 35 C in order to take advantage of air cooling. However, cooling water may be used to adapt the operation to local temperature conditions. The column C-2 deals with the water removal (dehydration). This is a difficult separation, because the equilibrium curve

5 1.02 26 1.02 25

GAS H-5

S-2

ABS-OUT

1.50

Q=-0.02

C-4 MEOH-ABS

34 S-1 36

Temperature (C) Pressure (bar) Q

Duty (MW)

RECYCLE

1.05

35 VD-1

1.02

1.10

MEOH-R LD-1

35 LD-2

1.10

VD-2

125

C-1

PROCESS

C-2

126

1.27

1.30

SS

RAW-AA

33 1.05

H-6

C-3

ACETIC Q=-1.28

131 QC=-2.87 QR=2.50

125 1.30

QC=-9.66 QR=5.63

QC=-5.45 QR=5.36

1.35 HEAVIES

CAT

Figure 13.9 Flowsheet of the acetic acid separation section.

Chapter 13 ACETIC ACID

1.60

119

120 1.10

503

Table 13.6 Selected Streams by Acetic Acid Separation. Units Phase Temperature Pressure Mass flows CO CO2 CH4 METHANOL ACETIC MEACET H2O CH3I PROPION

C bar kg/hr kg/hr kg/hr kg/hr kg/hr kg/hr kg/hr kg/hr kg/hr kg/hr

PROCESS

S5

RAW-AA

GAS

MEOH-R

Recycle

ACETIC

Vapor

Liquid

Liquid

Vapor

Liquid

Liquid

Liquid

125.1 1.6 32,047 235.0 78.0 22.0 4.0 25,661.0 2292.0 1608.0 1999.0 148.0

120.5 1.1 25,000 0.0 0.0 0.0 0.0 24,916.3 0.0 34.0 0.0 49.7

126.2 1.3 25,400 0.0 0.0 0.0 0.0 25,225.6 0.0 34.0 0.0 140.4

5.0 1.0 290 212.3 44.5 15.6 17.3 0.0 0.0 0.0 0.0 0.0

36.0 1.1 6637 11.0 17.1 3.2 5934.2 0.0 413.5 13.0 245.1 0.0

34.3 1.0 5290 11.7 16.5 3.2 52.5 15.7 1878.0 1557.9 1753.9 0.3

33.0 1.1 25,000 0.0 0.0 0.0 0.0 24,916.3 0.0 34.0 0.0 49.7

Table 13.7 Sizing Characteristic of the Distillation Columns by Acetic Acid Separation. Item

Unit

C-1

C-2

C-3

Feed Distillate V/L Side stream Bottom Stages/feed/ss Top pressure Pressure drop Top temp. Bottom temp. Reflux ratio Internals Stages/diameter

kg/h kg/h kg/h kg/h e bar bar C C mol/mol

m MW MW

29,800 400/4000 e 25,400 35/27 1.2 0.32 20 131 5.8 Sieves 2-20/1.9 21-26/2.25 27-34/1.9 20 9.7 5.6

25,400 25,000 e 400 30/15 1.1 0.26 121 131 1.0 Nutter-BDP 2-29/2.1

Total height Condenser duty Reboiler duty

32,047 200/1625 29,800 422 15/17/15 1.3 0.15 10 125 9.9 Flex-S 2-10/1.3 11-14/1.8 15-16/2.2 10 2.8 2.5

/m

17 5.5 5.4

Chapter 13 ACETIC ACID

1 0.9

Mass fractions

0.8 0.7

Acetic acid Me-Acetate Water Me-Iodide

0.6 0.5 0.4 0.3 0.2 0.1 0

1

6

11

16

21

26

31

Stages Figure 13.10 Concentration profiles as mass liquid fraction by the dehydration of acetic acid.

shows a tendency to azeotrope formation in the region of water separation (see Fig. 13.10). The column needs 35 theoretical stages. Simple sieve trays reveal to be the most suitable. Changing the tower diameter is necessary to ensure an efficient hydraulic regime, this time larger in the middle, as 1.9 m in the range of stages 2e20, 2.3 m for stages 21e26 and 1.9 m for stages 27e34, with the maximum flooding approach of 75%, 79%, 78% respectively. The vapor feed is located above the stage 27. The total tower height reaches 22 m with theoretical stages, which could rise to 30 m with real stages. The split of the column in three regions might surprise, but this correspond to different separation requirements, which may be identified by analysing the composition profiles (Fig. 13.10). The water removal increases continuously from bottom to the top, while opposite happens for the acetic acid. In the middle section above the feed the vapor flowrate is high, due to the vapor feed, and thus need a larger cross section to avoid flooding. An interesting result is obtained switching on structured packing. Selecting Mellapack from Sulzer company reveals that the tower diameter could drop significantly, to diameters of 1.4, 1.8 and 1.45 m and flooding approach of 78%, 72% and 80% respectively. Since HETP is about 0.4 m the total column’s height is reduced to 13 m. A notable effect is pressure drop reduction from 0.32 bar with trays to only 0.02 bar with packing. Similar results can be obtained using Super-Pak structured packing from Raschig company. Thus, from cost viewpoint employing structured packing seems very attractive.

505

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Chapter 13 ACETIC ACID

Finally, the acetic acid is obtained as top product in the column C-3 with a high purity over 99.7%. The bottom stream contains heavies, simulated here as the propionic acid. The column C-3 is provided with 30 trays. Note that the separation is tedious, the K-value of acetic acid to propionic acid being around 1.20. Accordingly, the separation needs high reflux flowrate, about equal with the top product flowrate. After several trials it comes out that Nutter-BDP trays give the best results. These are special rectangular-shaped valve trays with improved efficiency/capacity profile and turndown ratio compared to conventional round-valve trays (Sulzer Chemtech, 2008).

13.10 Process Integration 13.10.1 Integration of Reaction and Separation Sections The exothermic reaction develops an important amount of energy, equivalent to a duty of 15.7 MW. The reactor needs cooling but creates also opportunities for supplying energy to other units. Table 13.8 presents data regarding the heat exchangers implied in the reaction system, as duty, temperature range and exchange area calculated from partial heat transfer coefficients. One can see that the unit ST-GEN can transfer a significant amount of energy to the separation section as steam at 3.6 bar and 140 C, the reducing in heating needs being of 45% (6.2 vs. 13.5 MW). Moreover, the Pinch Analysis indicates that a

Table 13.8 Heat Exchanger Units by the Reaction Section. Unit

FLASHER H-1 H-2 ST-GEN COOLER H-3

Type

Heater

Cooler

Duty

Hot Side

Cold Side

Area

MW

T_in C

T_out C

T_in C

T_out C

m2

1.46 0.89 0.34 6.19 6.24 0.26

150 133 118 219 149 219

133 118 113 149 55 50

120 20 20 133 20 20

123 100 100 134 35 35

60 12 5 115 130 3

Chapter 13 ACETIC ACID

507

Table 13.9 The Heat Exchangers of the Separation Section of Acetic Acid Manufacturing. Unit

Utility

T/( C)

P/(bar)

Duty/(MW)

Reboiler C-1 Reboiler C-2 Reboiler C-3 Condenser C-1 Condenser C-2 Condenser C-3

LP steam

125 126 131 10 35 110

1.3 1.30 1.35 1.1 1.1 1.1

2.5 5.6 5.4 2.9 9.7 5.5

Brine Steam gen.

supplementary 16% reduction in utilities is possible in the reaction section. Heating is necessary for methanol and recycle preheating, as well as for pushing the full vaporization in the unit FLASHER, with a cumulated duty of 2.69 MW. On the other side the unit COOLER may also supply 6.24 MW on the hot side. An H-curve plot shows that 2.75 MW is available on the interval 150 to 110 C, enough for driving the units FLASHER, H-1 and H-2 and ensuring a driving force of at least 10 C. Accordingly, the duty of COOLER diminishes with 2.69 MW from 6.24 to 3.5 MW. Thus, the total energy saving opportunity regarding the reactor effluent is 72%. Table 13.9 presents the heat exchange units implied in the separation section. With respect to separation section, the total heating and cooling utilities are 13.5 and 18.1 MW, respectively. Pinch analysis indicates that no saving by process/process streams is possible. However, a closer physical analysis of the distillation system will suggest several energy saving opportunities. A first possibility is exporting the steam generated in the reaction section. The best coupling is with the column C-2 by which 6.2 MW from steam-generator can cover completely 5.4 MW for reboiler. The temperature driving force is feasible, above 10 C. It may be noted that lower pressure drop favors lower temperature in reboiler. This is possible if packing is used instead of trays, at least in the upper part of the tower.

13.10.2 Heat Pump Assisted Distillation A second possibility for energy saving is installing a heat pump between the condenser and reboiler of the purification column C-3. Heat pumps can be used in distillation to increase the energy

508

Chapter 13 ACETIC ACID

efficiency of columns, by upgrading the low temperature energy from the top (condenser) to higher temperature levels, such that the recovered heat can be reused to heat a lower column stage (e.g., reboiler). This can bring significant energy (utilities) savings, but the energy required to increase the pressure (and temperature level) is of higher quality and price than hot utilities (Luo et al., 2015). A simple criterion depending on the Carnot efficiency can be used to decide whether a heat pump is worth considering. This can be simplified to the following form (Kiss and Infante Ferreira, 2016): Q / W ¼ 1/h ¼ Tc / (Tr  Tc) > 10

(13.14)

where Q is the reboiler duty, W the work provided, h is the Carnot efficiency, Tr the reboiler temperature and Tc the condenser temperature. When the Q/W ratio exceeds 10, then a heat pump should be considered, whereas when the ratio is lower than five then using a heat pump will not bring any benefits (Plesu et al., 2014). In case of the acetic acid separation, the temperature difference top-bottom is less than 20 C and Q/W is 18.6 (Tr ¼ 391 K, Tc ¼ 373 K) which is clearly favourable. Water is suitable as thermal agent. Steam can be generated in top at atmospheric pressure and temperature close to 100 C, and then compressed to about 4.5 bar at 148 C, sufficient for driving the reboiler with a temperature difference of 15 C. Note that compressing low pressure water vapor is a widespread technique for energy saving in industries dealing with water evaporation, as seawater desalination, food, pulp and paper, etc. This method is based on mechanical vapor recompression (MVR). A quick economic comparison can be done in term of credit of saving thermal energy versus the cost of compressor. Assuming 10 $/GJ for LP steam and 8000 h annual operation time gives the cost of steam for reboiler of 10  5.4  8000  3600/1000 ¼ 1.55 M$. Considering cooling water at 2 $/GJ gives a supplementary cost of 0.31  106 $ and a total cost of 1.86 M$. On the other hand, a raw estimation of the compressor cost can be obtained by means of the correlation C ¼ a þ bSn, (Towler and Sinnott, 2013) in which a ¼ 580,000, b ¼ 20,000, n ¼ 0.6, S being the power in kW and the resulting cost in USD. We consider a 4-stage compressor with intermediate cooling and an overall efficiency of 0.75. The simulation in Aspen Plus gives a power of 1111 kW, that leads to estimate a compressor purchased cost of 1.92 M$. Considering an installation factor of two gives a total installed cost of 3.84 M$. By assuming 0.07 $/kwh gives anoperating cost of 1111  8000  0.07 ¼ 0.616 M$, only 33% from the annual cost of thermal energy.

Chapter 13 ACETIC ACID

The estimation of profitability can be done based on Total Annual Cost. This considers that the capital of installed compressor is paid back within 3 years, while the operating cost each year. As result the TAC is 3.84/3 þ 0.661 ¼ 1.89 M$, to be compared with the cost of thermal energy 1.86 M$. This estimation suggests that the operation may be feasible if an advantageous price may be obtained for the compression system.

13.10.3 Heterogeneous Catalyst Process If heterogeneous catalyst is used, as in CT-ACETICA process, the column C-1 for catalyst recovery is no longer necessary. The separation system reduces to two columns. The first column deals with the recovery of lights and water in top, while the second column with the acetic acid purification. The energy for the dehydration column can be supplied by the steam generated in the reaction section, while the purification column can be driven by a heat pump system. This alternative has the minimum utility requirements. The only energy necessary is for driving the MVR system. From section the above calculation the compressor energy is 1111 kW, equivalent with 3.3 MW thermal duty. The energy saving with respect to the three-column separation sequence is (1e3.3/13.5)  100 ¼ 76%.

13.11 Economic Evaluation The economic analysis starts with the estimation of the total capital cost (TCI). This case study illustrates the capabilities of Aspen Plus Economic Analyser (APEA). This interactive tool may perform a full estimation of capital and operation costs, including the cost of utilities. The calculation is based on detailed sizing and cost accounting of various components, as heat exchangers respecting TEMA norms, various vessels, compressors and distillation columns. APEA calculation is based on industrialproved methods supported by a large database with equipment, materials, utilities and cost elements. Some complex units are treated as “modules”, which spares tedious calculations. For example, a distillation unit is assembled from components, as tower, internals, reboiler, condenser, vessels and pumps. Key results from APEA are the purchased equipment cost (PEC) and installed equipment cost (IEC) for the mapped items of equipment, as well as the consumption of utilities. These data may be exported in EXCEL and further exploited. For example, when the calculation is confronted with unavailable or partially solved items

509

510

Chapter 13 ACETIC ACID

spreadsheet calculation may be employed considering correlations and data from literature (Towler and Sinnott (2013), Turton et al. (2013); Dimian et al. (2014)). Then combining all these elements gives the total installed equipment cost, known as ISBL (Inside Battery Limits) cost, as well as the total utilities cost, which are at the fundamental elements of the economic analysis. The procedure developed below illustrates how combining APEA estimations with user data in spreadsheet calculations. TCI is obtained by estimating the fixed capital (FC) and working capital (WK). FC cumulates inside battery limits (ISBL) costs, offsite or outside battery limits (OSBL) costs, Engineering and Construction and Contingency (E&C) costs. The basis for determining FC is the ISBL cost, since the other contributions may be estimated as ratios to ISBL. In this project we consider the ratios recommended by Douglas (1989): OSBL ¼ 0.45  ISBL, E&C ¼ 0.25  (ISBL þ OSBL), WK 0.15  TCI and ST ¼ 0.10  FC, which lead to the expressions FC ¼ 1.8125  ISBL and TCI ¼ 2.36  ISBL. Table 13.10 shows the estimation of purchase and installation costs of equipment grouped on equipment type. One may see that the largest cost regards the distillation section, followed by reaction section, which includes the cooling system. The cost of compressor is also important. Note that the type of construction material was included being accessible as parameter. The acetic acid is very corrosive and therefore the equipment needs special steels. Hence AISI SS321 is used for the reaction section, SS316 for distillation columns, and SS304 for flash and storage vessels. The application is specific for each item, for example in the case of steam generator SS321 for tubes and SS304 for shell. The estimation for the base equipment cost (BEC) is 10.4 M$ for installed cost with 6.84 M$ for purchased cost. Notably, the ratio of 1.52 is significantly lower than proposed by factorial methods, for example 3.3 for fluid processes (Towler and Sinnott, 2013). In addition, the ISBL costs considers additional equipment not included in simulation and storage facilities, as 0.3 and 0.2 ratios to BEC, plus the costs for catalyst and CO production. In this way the ISBL costs raise to 20.9 M$. Then the fixed capital FC cost may be got easily by applying the above-mentioned contributions, namely off-sites, engineering and construction, contingencies, as well as working and start-up expenses. Here the utility system is included in OSBL costs. The estimation of the total project cost is finally 49.2 M$. On this basis a provision for capital depreciation of 4.9 M$ is determined as a linear reimbursement over 10 years.

Chapter 13 ACETIC ACID

511

Table 13.10 Capital Cost Calculation for the Acetic Acid Production Process. Equipment Reaction section Reactor ST GEN COOLER Reaction section Compressor Heat exchangers H-1 H-2 H-3 Total HX Vessels FL-2 Flasher Total vessels Distillation columns C-1 C-2 C-3 C-4 Total distillation Base equipment Aux. Equipment Storage Catalyst plant CO production Total cost Fixed capital TIC

Ratio

0.3 BE 0.2 BE

PEC, $

IEC, $

399,400 1,183,700 132,900 1,716,000 1,895,000

591,700 1,526,800 271,100 2,389,600 2,166,500

14,600 12,000 11,200 37,800

80,800 84,100 79,200 244,100

60,200 172,600 232,800

157,800 367,800 525,600

691,200 1,081,000 1,184,100 57,100 2,956,300 6,837,900 2,051,370 1,367,580 750,000 1,000,000 12,006,850

1,322,300 1,808,100 1,960,300 183,900 5,090,700 10,416,500 3,124,950 2,083,300 2,250,000 3,000,000 20,874,750 37,835,484 49,264,410

1.82 ISBL 2.56 ISBL

The design efficiency may be assessed by calculating the profit after taxes and the payback time of the capital based on given values of prices for materials, products and utilities. The following prices per metric ton (MT) or 1000 kg are considered: 350 $/ MT-methanol (Intratec solutions, 2019), 100 $/MTeCO internal price by syngas manufacturing, as well as 10% fees for catalyst

512

Chapter 13 ACETIC ACID

and additional chemicals. With the above elements the raw material cost is RM ¼ 1.1  (13,612  350 þ 12,274  100)  8000/ 106 ¼ 52.73 M$/year. The cost of utilities results from the energy balance. The prices of utilities are: electricity 0.07 $/kWh, steam 10 $/GJ, cooling water 2 $/GJ or 0.05 $/m3 by tower cooling, refrigeration 15 $/GJ. Table 13.11 shows the consumptions in reaction and separation sections. Note that the cooling of the reactor effluent after process/process exchange is done by air, thus zero cost. The cooling water is limited at 10% from the total duty of condensers, at 1.8 MW. In addition, a temperature difference of 10 C temperature difference between the inlet and outlet of the cooling tower is assumed. Otherwise air cooling is used. The result is a total cost utility cost of 5.0 M$. Note that the calculation takes as credit the LP steam generated by reaction. Steam is by far the largest consumer, followed by electricity. In Dimian et al. (2014) a simple equation for estimating the operation expenses OPEX was proposed: OPEX ¼ 1.11  (RM þ U) þ 0.183  FC þ 0.028  S (13.15) The advantage is using information available directly from simulation, namely the requirements in raw materials (RM) and utilities (U), the fixed capital (FC) and the estimated sales (S). The process profitability target is the product cost that ensures a ROI of 20% for a tax rate of 35%. Table 13.12 presents the computation. The result is a product cost of 456 $/MT. Comparing with 231.5 $/MT as average price of raw materials, the process gives a value-added of 214.5 $/MT.

Table 13.11 The Duty and Cost of Utilities by the Acetic acetic Manufacturing.

Electricity Steam MP Steam LP Cooling Refrigerant Total utilities, $/h Total utilities, M$/year

Reaction MW

Separation MW

Reaction $/h

Separation $/h

Total $/h

1.59 0 6.22 0 0.24

e 5.4 13.5 1.8 0.05

112.4 0.0 193.5 0 12.96 68.1 L0.545

e 194.4 486.0 7.5 2.7 696.9 5.575

112.4 194.4 292.5 7.5 15.7 622.5 4.980

Chapter 13 ACETIC ACID

Table 13.12 Product Cost and Profitability Calculation. Description

Value

Unit

Raw materials Utilities Product cost Sales Depreciation OPEX Profit before taxes Profit after taxes Cash flow Return of investment Payback time

52.73 5 456.00 91.20 4.93 73.77 12.50 8.13 9.85 20.0 5.00

M$ M$ $/MT M$ M$ M$ M$ M$ M$ % Years

13.12 Sustainability Analysis Sustainability metrics for design projects are presented in the Chapter 17 of the authors’ book (Dimian et al., 2014) with reference to the vision of the Institution of Chemical Engineers in the UK issued in 2002. This contains a list of indicators regarding the usage of resources, then emissions, effluents and wastes, and additional environmental impact factors. A simpler but efficient approach capturing the essential of the above method is applied in this case study. The approach proposed by industrial experts (Schwarz et al., 2002) is based on six metrics: material intensity, energy intensity, water consumption, toxic emissions, pollutant emissions, and greenhouse gas (GHG) emissions. The metrics are determined by reference to an output, which may be the unit of product(s), the revenues (sales), or the value-added (VA). Determining VA is a complex matter, but it should refer as the process ability to create more value by products and services. A simple but effective definition VA is the difference between Sales and Raw Materials plus Utilities. Considering the data from Table 13.12 gives the result: $VA ¼ 91.2 e (52.73 þ 5) ¼ 33.28 M$/year or 4160 $/h. The following sustainability metrics are briefly outlined. Note that lower values of metrics mean better process performance.

513

514

Chapter 13 ACETIC ACID

• Material intensity expresses the mass of wasted materials per unit of output. This may be determined as the mass of raw materials minus the mass of product(s). • Energy intensity is the energy consumed per unit of output for satisfying the heat and power requirements. For consistency reason, a difference should be done between the energy developed by the utility and the primary energy used for producing it. For example, electrical energy may be obtained in a power plant by using various energy sources, as coal, petroleum, gas, biomass, and nuclear. Following the 2017 data of US Energy Information Administration (EIA) typical “heat rate” was 2.3 for natural gas and 3.1 for coal and nuclear sources. As a rule of thumb, a unit of electrical energy needs three times more thermal energy to produce it. But, when using combined heat and power (CHP) systems, the heat rate (1/electricity production efficiency) may drop even below 2. • Water consumption includes water lost through waste treatment or disposal, as well as the amount of fresh water injected to keep the inventory of the cooling water system because evaporation and losses. • Toxic emissions are the amounts of toxics materials that are hazardous for the operating personnel. The toxicity may be characterised by the EPA database in USA or the REACH system in Europe for the registration, evaluation, authorisation and restriction of chemicals. • Pollutant emissions designates the materials affecting the environment, as air acidification, water eutrophication, ozone depletion, acidification of fresh water. • Greenhouse gas emissions contribute to the global climate change. These regard chiefly the CO2 emissions but also other species, as methane, which are converted in equivalent emissions. Carbon dioxide emissions from the steam and electricity generation are included, even though these utilities are imported rather than generated onsite. The applications of these metrics for the acetic acid manufacturing by methanol carbonylation is illustrated in the following part. Tables 13.13 and Table 13.14 present a synthesis. • The material intensity may be assessed from the global material balance as follows: Waste ¼ ((13612 þ 12273)25000)/25000 ¼ 886/25000 ¼ 0.03544 kg/kgAA The waste produced by value-added dollar is 886/4160 ¼ 0.213 kg/$VA

Chapter 13 ACETIC ACID

515

Table 13.13 GHG Emissions by Acetic Acid Manufacturing. Utility

Amount

Unit

Conversion

kg/h CO2e

Electricity Steam Cooling water Total

1600 48.6 154.8

kWh GJ m3/h

0.707 65 2.77

1131 1722 429 3282

Table 13.14 Sustainability Metrics by Acetic Acid Manufacturing. Metrics

Homogeneous Heterogeneous Benchmark

Material

Energy

Water

CO2e

kg/kg

kg/$VA

MJ/kg

MJ/$VA

kg/kg

kg/$VA

kg/kg

kg/$VA

0.0354 0.0248 e

0.213 e 0.177

2.07 1.57 e

12.46 e 15.9

0.098 0 e

0.585 e 3.574

0.131 0.076 e

0.789 e 1.107

• The energy intensity results from the energy balance, as shown in Table 13.11. The reaction section can supply 6.2 MW LP steam to the separation section. Here the heat consumers are the reboilers of the distillation columns C-1, C-2 and C-3 with duties of 2.5, 5.6 and 5.4 MW respectively, in total 13.5 MW. LP steam of 3.6 bar and 140 C is convenient for the first two columns. MP steam of at least 4 bar is suitable for the last. The hot utility for the separation section is 13.5 e 6.2 ¼ 7.3 MW for driving the reboilers. The electricity consumption counts 1.6 MW needed for CO compression, which is equivalent with 3  1.6 ¼ 4.8 MW thermal energy. The process needs also energy for cooling by refrigeration, in total 0.3 MW, as well as the energy involved with the cooling water 1.8 MW. The total energy consumption for the reaction and separation sections is 7.3 þ 4.8þ0.3 þ 1.8 ¼ 14.2 MW. Finally, the energy intensity metrics are 14.2  3600/25,000 ¼ 2.07 MJ/kg and 14.2  3600/ 4160 ¼ 12.46 MJ/$VA. If referring only to hot utility the energetic intensity is 6.32 MJ/$VA.

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• Water consumption is based mainly on the cooling water requirements. The assumption introduced previously is that most of the cooling in the separation section is done by employing air-coolers, and only 10% from load by watercoolers, or 1.89 MW. The temperature “range” from process to cooling tower is 10 C, from example from 20 to 30 C. Accordingly, the cooling capacity of water CP is 4.186  10 ¼ 41.86 kJ/ kg. The flowrate is 1.8  1000/41.86  3600/1000 ¼ 154.8 m3/h. Because the water cooling is obtained by evaporation, this must be compensated by a make-up with fresh water. Following the 7% rule (Schwarz et al., 2002), the loss is 0.07  154.8 ¼ 10.836 m3/h. Another estimation method can be found in Chapter 12 of Perry’s Handbook (2008) dealing with the design of tower coolers. The fraction of water lost by evaporation can be found with the relation Fev ¼ 0.00,085  (Tin-Tout), with temperature in  F. Considering 3% condensation for drift loss and blowdown, and expressing the temperature in C, the relation becomes Fev ¼ 0.00,158  (Tin-Tout). For 10 C difference the fraction evaporated 1.58%, or a loss of 2.561 m3/h. The water consumption intensity is 2.44  1000/25,000 ¼ 0.098 kg/kg and 0.585 kg/$VA. Note that Turton et al. (2013, Chapter 8) estimates the water loss by evaporation as 1.83% from the inflow for 10 C range. Thus the above 7% rule seems to overestimate the make-up and energy needs for cooling water. • The estimation of GHG emissions cumulates the amounts involved in the production of heat and power, electricity, and recycling the cooling water. Table 13.13 shows the input data. The key problem is the knowledge of the conversion factors. Thus, for electricity the conversion factor depends on the energy mix used by electricity production, which depends on the national grid or the industrial site. Here we adopt the USA average value of 0.707 kg/h CO2e per kwh (EIA, 2018). The CO2 conversion factor for steam is also variable. A reasonable assumption is considering a primary energy ratio electricity/steam of three that leads to 65 kg/h CO2e per GJ. With respect to cooling water, a value of 2.77 kg/h CO2e per m3/h was obtained from the APEA tool. When examining the values in Table 13.13, it can be observed that the CO2 emissions for steam supply are the largest. Reducing the steam load creates an opportunity for dropping the CO2 emissions. The CO2 metrics are: 3282/25,000 ¼ 0.131 kg/kg product and 0.789 kg/$VA. • The toxicity metrics characterise the threats on health of species leaving in the process outputs. The main exit points of gaseous streams in the process is VENT (Fig. 13.8) and

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GAS (Fig. 13.9). The most toxic species is methyl iodide. The summary of EPA description is: In humans, acute (shortterm) exposure by inhalation may depress the central nervous system, irritate the lungs and skin, and affect the kidneys. Massive acute inhalation exposure led to pulmonary oedema. Chronic (long-term) exposure of humans by inhalation may cause skin burns. EPA has not classified methyl iodide for potential carcinogenicity. Methyl iodide is very volatile and difficult to separate from a large amount of gas components, as CO and CH4. Two methods may be envisaged: deep cooling and absorption in methanol. The first is applied here to the gas reactor effluent. By cooling at 5 C about 98.8% of methyl iodide is condensed and recycled, but an amount of 0.2 kg/h is still lost in the stream VENT. The absorption in methanol works better with over 99.8% recovery. However, the exit gas stream may carry some methanol, which should be recovered by cooling. Hence, for the recovery of lights the combination of absorption in methanol with deep cooling offers the best solution. • The pollution metrics characterise the potential damage of process effluents on the environment. The gas streams analyzed so far contain large amounts of CO and CH4. These may be submitted to VOC treatment that takes place usually by combustion. About 1000 kg/h CO2 leaves as emissions the plant. This amount is important, but less than the CO2 created by the utilities. As may be seen by examining Table 13.14, the sustainability metrics determined in the presented project are comparable with the benchmark values. However, there is place for improvement. Using heterogeneous catalysis brings substantial simplifications in technology. The second line displays the target. The usage of raw may be improved by raising the methanol yield from 0.98 to 0.99 and of CO from 0.95 to 0.96. The material intensity drops to 0.0258 kg/kg-product. The energy needs diminish substantially. Thus, the column C-1 and C-2 are lumped in a single column dealing with lights and water recovery. Note that the immobilized catalyst needs considerably less water, maximum 2%, which means less energy spent for dehydration. The heat for driving the reboiler can be ensured entirely by the steam generated in the reaction section. Then the purification column may be equipped with a heat pump. The result is that the imported steam for the separation section becomes zero. The energy injected in the process involve only the energy for CO compression in the reaction section and for heat pumping by MVR in the separation section. Translated in

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thermal energy these are equivalent with duties of 4.8 and 3.3 MW, or 8.1 MW in total. The intensity metrics become are 8.3  3600/25,000 ¼ 1.20 MJ/kg and 8.3  3600/4160 ¼ 7.18 MJ/$VA. Finally, these savings reduce the energy requirements by more than 75% compared with the non-integrated base case. Another beneficial action is on water consumption. In principle, the cooling may be based on air everywhere, except the final condensing of lights by refrigeration. The availability of a cooling tower is no longer necessary. The CO2 are reduced only at the electricity consumed by compressors. We encourage the reader to work out the design project of the heterogeneous process.

13.13 Conclusions Acetic acid is an important chemical intermediate that can be produced by a sustainable process by methanol carbonylation using homogeneous or heterogeneous rhodium/iridium catalysts. Since methanol is available from renewable biogas or biomass, acetic acid may be produced as 100% bio-based product. In this work the emphasis is set on energy efficiency, since the exothermal reaction releases a large amount of energy. Thus, the design of the reaction and separation systems is integrated. In the homogeneous process the separation section consists of three columns, for lights’ removal, dehydration and acetic acid purification. The LP steam generated in the reaction section may cover most of the requirements of the first two columns, and thus about 45% from the heat needed in the separation section. Substantial energy saving may be obtained by applying heat pumping to the purification column. The energy available in condenser at temperature slightly above 100 C is upgraded for reboiler heating by using mechanical vapor recompression with water as thermal agent. By combining steam generation with heat pumping the energy requirements may be reduced by 75% with reference to the base case. An economic analysis is performed for estimating the capital and operation costs, as well as for determining the product price that would bring a profitability of 20% ROI. The case study shows also how to calculate comprehensive sustainability metrics for process design. These regard the use of raw materials, energy, water, greenhouse gas effects, toxicity and pollution. The homogeneous-based catalyst designed process has good sustainability metrics, in agreement with an industrial benchmark. Moreover, a substantial leap forward may be achieved

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by adopting the heterogeneous catalyst and energy saving methods developed in this work.

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