Production of Acetaldehyde From Acetic Acid

Department of Chemical & Biomolecular Engineering

Senior Design Reports (CBE)

University of Pennsylvania Year 2002

Production of Acetaldehyde from Acetic

Acid

Calvin daRosa Aurindam GhatakUniversity of Pennsylvania University of Pennsylvania

Claire PintoUniversity of Pennsylvania

This paper is posted at ScholarlyCommons.

http://repository.upenn.edu/cbe sdr/45

PRODUCTION OF ACETALDEHYDE FROM ACETIC ACID

Authors:

Calvin daRosa

Aurindam Ghatak

Claire Pinto

Faculty Advisor

Dr. John V ohs

April 9, 2002

Department of Chemical Engineering

University of Pennsylvania

/,., /1 L ,::;'1./6:.r/ClJ} - ~ 0rJ~11 6( I

09 April 2002

Dr. John Vohs Prof. Leonard Fabiano Department of Chemical Engineering University of Pennsylvania Philadelphia, P A 19104

Dear Dr. Vohs and Prof. Fabiano:

Enclosed in this report is the completed economic analysis of our proposed process. The

process is designed to produce and recover acetaldehyde at high purity from acetic acid.

This process design recovers 12,818 lblhr of acetaldehyde by extractive distillation at

99.6 % weight purity. A second commodity chemical, ethyl acetate, is produced as a

byproduct in this process and is purified by a series of distillation columns. Ethyl acetate

is produced at a rate of 1,139 lblhr and purity of 99.6 weight percent.

Capital cost estimations and profitability analysis have been completed for our process.

Financial modeling of our process assuming the price of acetic acid to be 0.16/lb yielded

an Investors Rate of Return (IRR) of 11.4 % and a Total Capital Investment (TCI) of $47,242,990. This scenario is not economically feasible. However, when the price of acetic acid is taken to be $0.12/lb, the IRR and TCI are 18.5 % and $47,224,990 respectively. In the light of that fact that the possible legislation ofMTBE out of gasoline

might make this process more economically attractive, the group recommends further

research into the feasibility of such a plant and the possible future construction of the

facility given the realization of the second scenario.

Sincerely,

~0?d4--- ~ CO'{)jJaU~~ Calvin P. daRosa Aurindam K. Ghatak Claire L. Pinto

TABLE OF CONTENTS

Abstract. ................... .......... . .................................. .. ........ . .... 5

Introduction.... . , .... . .................. .. ................................... , ........ 7

Process Flowsheet. .................. . ................................................ 13

Material Balance ...... . .................. . ................................... . ....... 16

Process Description ................... . ......................................... . .... 23

Energy Balance and Utility Requirements ..... ............................ . ......35

Unit Descriptions ............ . .... . .................. . ................................ 39

Absorber . . .................... . .... . ....... . ............... 39

Compressor. .. . .............................. . ... . . . ....... 40

Condensers. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 40

Decanter... . ... . . . ...... . ........ . ......................... . 44

Distillation Columns ..... . ......... . ..................... 44

Fired Heater. ...... . ................. . .............. . .. . ... 49

Flash VesseL ......... . .................................... 50

Heat Exchangers ......... . ........... . .................... 50

Mixers ....... . .............................................. 53

Pumps .. . . . ..... . .......................................... . 54

Reactor. ...... .. ............... ... .... . .. .. ............. . ... 59

Reboilers .. . ...... . .................. . .... . .... .. ....... . .. . 61

Reflux Accumulators ........... .. ......... . .............. 63

Refrigeration System .. . ..... . ................. . ..... .. ... 65

Splitters . . . . .... .. ..... .. . .. ....... . ....... .. ......... . ...... 65

Stripper ............ .. .... .. ... . ........... " . ... . . .. .... . .. 66

Tanks . . .. . .................................................. 67

Valve ................. . ..... . .......... ... . . ...... . ......... .. 70

Unit Specification Sheets .................. . ..... . ...... . ........... . ..... . .......... 71

Equipment Cost Summary......... . ..... . ..... . ..... . .. . .. . ...... . .... . ... . ........ 135

Fixed-Capital Investment Summary.................... ...... . ..... ... ... . ... . .... 137

Important Considerations ...... , .... . ... . .. . .... .. ... '" .... .. ..... . ..... . .. . ....... .141

Operating Cost and Economic Analysis ................. . ..... . ..... . ......... . . . 145

Conclusions and Recommendations ........... . .......... . ....... .. ........... . ... 159

Acknowledgements ........... . ..... . .................. . ..... . ... . . . ................. 161

Bibliography.... . ....... , .. .. . .... . ....... . .......... . , . ..... . .... .. ... . ...... . ....... 162

Appendix A: Unit Cost Calculations . . ........ ... ..... .. .... . ..... . ..... . ..... . .. . 163

Appendix B: Utility Cost Calculations ... .. .................... . ... .... .......... .201

Appendix C: Aspen Plus Results ... . ... . .... . . .. ........ .. ... . ... . ........ . ..... ... 207

Appendix D: Problem Statement. . .. . .......... . ........... .. ........... . ...... .... 269

Appendix E: Patent. .... . ....................... . . . .. . ..... . , ..... . ............ , ..... . 273

Abstract

Our group has designed a process to manufacture 101,520,000 lb/yr of

acetaldehyde by hydrogenation ofacetic acid over a 20% wt. palladium on iron oxide

catalyst. The reaction conditions used are the optimum according to a patent filed by

Eastman Chemical (Tustin, et.al., U.S. Patent No. 6,121,498): temperature is range is

from 557-599 of at a pressure of254 psi. The conversion of acetic acid in the reactor is

46 %, with selectivity of 86% to acetaldehyde. Major by-products are ethanol, acetone,

carbon dioxide, and the light hydrocarbons methane, ethane, and ethylene. Acetaldehyde

is purified in a series of steps: it is first absorbed with an acetic-acid rich solvent, then

distilled to separate acetaldehyde from heavier components. A refrigerated condenser is

then used to recover additional acetaldehyde from the vapor distillate of the main

separation. Acetic acid is purified and recycled to the reactor to limit the amount of

feedstock required. Ethyl acetate is produced as a by-product in the acetaldehyde

distillation column and is purified and sold.

The economics of the process is strongly dependent on the price of acetic acid,

and we examined scenarios under which acetic acid was available at either $0.16/Ib or

$0.12/Ib. The total capital investment in either situation is approximately $47,000,000.

If acetic acid is available at $0. 16/1b, we estimate an IRR of 11.3 %, but if acetic acid can

be purchased for $0.12/Ib the IRR is 18.5% after 20 years. It is our recommendation to

pursue more research into projecting both the cost of acetic acid and the market for

acetaldehyde. If acetic acid will be available at the lower price, the company should

pursue production of acetaldehyde.

5

Introduction

The main product manufactured in this process is acetaldehyde. Acetaldehyde

was chosen as the primary product because of its wide use in industry and the

profitability of its sale as a chemical of given purity. In addition to acetaldehyde, ethyl

acetate is produced as a side product. Table 1 below shows basic chemical information

concerning the products.

Table 1: P f f duct-

Acetaldehyde Ethyl Acetate

Synonym Molecular Formula Molecular Weight CAS No. Melting point Boiling point Density

(grimary product) Ethanal C2H4O 44.05

75-07-0 -190.3 OF 69.6 OF

0.6149 g/cmJ

{side product) Acetic acid ethyl ester

CH3COOC2Hs 88.0

141-78-6 -117 OF 171F

1.108 g/cmJ

I. Uses

Acetaldehyde is primarily used in industry as a chemical intermediate, principally for

the production of pyridine and pyridine bases, peracetic acid, pentaerithritol, butylene

glycol and chloral. It is used in the production of esters, particularly ethyl acetate and

isobutyl acetate (lARC V.36 1985; Chern. Prod. Synopsis, 1985). It is also used in the

synthesis of crotonaldehyde as well as flavor and fragrance acetals, acetaldehyde 1,1

dimethylhydrazone, acetaldehyde cyanohydrin, acetaldehyde oxime and various acetic

esters, paraldehyde halogenated derivatives (lARC V.36, 1985). Acetaldehyde has been

used in the manufacture of aniline dyes and synthetic rubber, to silver mirrors and to

harden gelatin fibers (Merck, 1989). It has been used in the production of polyvinyl acetal

7

L

resins, in fuel compositions and to inhibit mold growth on leather (lARC V.36, 1985).

Acetaldehyde is also used in the manufacture of disinfectants, drugs, perfumes,

explosives, lacquers and varnishes, photographic chemicals, phenolic and urea resins,

rubber accelerators and antioxidants, and room air deodorizers; acetaldehyde is a

pesticide intermediate (Sittig, 1985; Gosselin, 1984).

Acetaldehyde, an alcohol denaturant, is a GRAS (generally recognized as safe)

compound for the intended use as a flavoring agent (Furia and Bellanca, 1975; HSDB,

1997). It is an important component of food flavorings added to milk products, baked

goods, fruit juices, candy, desserts, and soft drinks. In 1976, approximately 19,000 Ib of

acetaldehyde were used as food additives, primarily as fruit and fish preservatives and as

a synthetic flavoring agent to impart orange, apple and butter flavors.

Ethyl acetate, the side product of this process, is widely used in printing inks, paints

and coatings, pesticides, pharmaceuticals, laminations and flexible packaging.

II. Production

i. Reasons for entering the market

Acetaldehyde was first produced commercially in the United States in 1916. U.S.

Production of acetaldehyde reached its peak in 1969 at approximately 1.65 billion lb

(lARC V.36, 1985). There has been an overall decline in the demand for acetaldehyde

due to the use of more economical starting materials for principal derivatives and a lower

demand for some acetal derivatives (Chern. Prod., 1985). However, in recent times due to

a decline in the number of suppliers and an increase in potential U.S. acetaldehyde

exports, there is a vast potential for profitability in manufacturing acetaldehyde. In 1985,

8

estimated U.S. exports of acetaldehyde were 1.2 billion lb (Chemical Products Synopsis,

1985).

ii. Alternative processes and their disadvantages

Acetaldehyde can be made commercially via the Wacker process, the partial

oxidation of ethylene. The major disadvantage of that process is that it is very corrosive

requiring very expensive materials of construction. Another major disadvantage is that

the reaction is prone to over-oxidation of the ingredient, the products ofwhich are

thermodynamically more stable than acetaldehyde which is the partial oxidation product.

This over oxidation of the ingredient reduces the yield of acetaldehyde produced and

converts expensive ethylene into carbon oxides (Tustin, et al.).

Acetaldehyde is also manufactured by oxidizing ethanol using air. A mixture of

air and ethanol vapor is fed into a multi-tubular reactor. Temperature is maintained

between 750-932 of (400-500 C), and the pressure at 29.4 psi. The catalyst used is

chromium activated copper. Vapor coming out of the reactor is passed through a scrubber

and unreacted ethanol is separated and recycled. However, this process gives a relatively

poor yield of acetaldehyde

The process investigated in this report converts acetic acid into acetaldehyde.

Acetic acid is relatively inexpensive and is available at $0.12-$0.16/lb. It can be

generated from inexpensive methanol. Due to the possible legislation ofMTBE out of

gasoline, there may be a worldwide glut ofmethanol, so any chemicals that use methanol

may become much more economically attractive. That is why acetic acid is our starting

material of choice.

9

The catalyst used in this process is 20% palladium on an iron oxide support. This

catalyst gives a selectivity of 86% to the desired reaction at 46% acetic acid conversion.

Though this process can also be effectively catalyzed by mercury compounds, the toxic

nature of mercury makes it unfeasible.

iii. Discussion of Production Method

The reaction is carried out in a packed bed reactor at a temperature range between

557 and 599 of. The following reactions occur in the reactor:

CH)COOH + H2 -> CH)CHO + H20 (main reaction) (1)

CH)COOH + 2H2 -> CH)CH20H + H20 (2)

2CH)COOH -> CH)COCH) + CO2+ H20 (3)

3CH)COOH + 9H2-> 2CH4 + C2H6 + C2~ + 6H20 (4)

Under reaction conditions, the selectivity to reaction (1) is 86%. This facilitates a good

yield of acetaldehyde and further justifies the cost of the reactor conditions. The product

is then passed through an absorber and then separated as the distillate using a fractional

distillation column. The following reaction occurs in the distillation column to produce

ethyl acetate:

CH)COOH + CH3CH20H -> CH3COOCH2CH3 + H20 (5)

This generation of ethyl acetate in situ is beneficial as it facilitates an ethyl acetate-water

azeotrope in the acetic acid separation column, which makes it easier to separate the

acetic acid. This acetic acid is then recycled back to the reactor. After being separated

from the water, first in a decanter and then by distillation, the ethyl acetate is purified to

99.6 % wt, and can be sold.

10

The Gulf Coast is the location of choice for this plant. This is primarily due the region

being an industrial belt. As a consequence, storage facilities as well as raw materials are

readily available and cheap. As mentioned in the problem statement, due to this choice of

location, it is assumed that hydrogen can be purchased over the plant fence for $0.50Ilb at

200 psig. Additionally, the prices of utilities are relatively inexpensive. Natural gas is

available at $2.30IMMBTU, cooling water is purchased at $0.33IMGal and steam at 35

psi at $2.46IMLbs.

III. Environmental issues and potential safety problems

EP A regulates acetaldehyde under several Acts such as the Clean Air Act (CAA) and

the Clean Water Act (CWA). EPA has established water quality criteria, effluent

guidelines, rules for regulating hazardous spills, general threshold amounts and

requirements for the handling and disposal of acetaldehyde wastes. Process enclosures,

local exhaust ventilation and other engineering controls must be used to maintain

airborne levels below maximum exposure limits.

Acetaldehyde is an extremely flammable liquid and vapor. Its vapor may cause flash

fires. It forms explosive peroxides and polymerizes, resulting in hazardous conditions.

Acetaldehyde is therefore sold in stainless steel tanks with a refrigerating system to

ensure that the temperature of the product does not rise above 15C.

Acetaldehyde is also a potential cancer hazard. High vapor concentration may cause

drowsiness or irritation of the eye and respiratory tract. For eye protection, safety glasses

with side shields and a face shield need to be worn by people with risk of exposure.

Additionally, chemical resistant gloves, boots and protective clothing appropriate for the

11

risk of exposure need to be worn. Decontamination facilities such as eye bath, washing

facilities and safety shower must be provided.

Ethyl acetate, on the other hand, is not subject to EPA emergency planning

requirements under the Superfund Amendments and Reauthorization Act (SARA) (Title

III) in 42 USC 11022. However, ethyl acetate is an irritant of the eyes and upper

respiratory tract at concentrations above 400 ppm [NLM 1992]. Ethyl acetate

occasionally causes sensitization, with inflammation of the mucous membranes and

eczema of the skin [Hathaway et al. 1991]. As a consequence, ethyl acetate is stored in a

cool, dry, well-ventilated area in tightly sealed containers. Splash-proof chemical safety

goggles or face shields and coveralls should be worn during any operation involving

potential exposure to ethyl acetate.

12

w

9 Temperature ;, FEED TODC-SOO) PressureY

TO

~ :s:: :p .....

o

TO

~ :s:: JJ

~ o

LEGENDS

PURGE2

SOLVENT FROM P-540

F-230

LEGENDS OFFGAS

9 Temperature y Pressure

...l-

SOLVENT FOR Z AB-320

COOLANT

"

l iY2GJ

HAC PRODUCT2

HAC PRODUCT1

SPLIT FROM DE-720

DC-10

ACETALDEHYDE FROM ACETIC ACID PROCESS (PART 3 OF 3)

40

TO P-620

28

DE720 STEAM __._~ . ~ ~ cw ~ ~ u.

LEGENDS STEAM y Temperature -~ y Pressure ~WASTEWATER

MATERIAL BALANCE

..J

Item Number: 5206 5206a 5206b 5207 5208 5301 5302 5303 5304 5305

Temperature (oF) 556.8 556.8 556.8 202.6 268 280.1 113 113 159.2 126.4 Pressure (psia) 250.9 250.9 250.9 262 .8 245.92 245.7 244.7 244.7 235.0 234 .9 Total Mass Flow (Ib/hr) 62,977.88 42,195.18 20,782.70 20,402.42 42,195.18 62,979.67 62,979 .67 23,469.70 62,277.91 65,870.50 Components (Ib/hr): HYDROGEN 7,088.04 4,748.99 2,339.05 7,781.47 4,748.99 7,088.04 7,088.04 7,088.02 trace 0.02 CO2 5,425.89 3,635.35 1,790.54 5,137.53 3,635.35 5,425.88 5,425.88 5,328.15 trace 122.96 METHANE 1,565.20 1,048.68 516.52 1,495.12 1,048.68 1,565.20 1,565.20 1,543.27 trace 28.45 ETHYLENE 916.61 614.13 302.48 855.34 614.13 916.63 916.63 893.76 trace 27.13 ETHANE 921.10 617.13 303 .96 855.42 617.13 921.10 921.10 898.61 trace 31 .92 ACETALD 13,468.60 9,023.96 4,444.64 621.46 9,023.96 13,471 .78 13,471.78 6,792.88 6.47 6,167.54 ACETONE 665.97 446.20 219.77 285.40 446 .20 665.76 665.76 119.04 785.89 615.79 ETHYLACE 908.03 608.38 299.65 906.51 608 .38 907.31 907.31 126.56 2,775.68 1,983.88 ETHANOL 771 .86 517 .15 254.71 17.25 517.15 771 .84 771.84 33.52 110.70 126.75 WATER 7,201.54 4,825.03 2,376.51 933.97 4,825.03 7,201 .08 7,201.08 266.36 13,552.93 12,873.06 HOAC 24,045.05 16,110.18 7,934 .86 1,512.96 16,110.18 24,045.05 24,045.05 379.53 45,046.25 43,892.99

*** VAPOR PHA5E *** Density kg/cum 4.69 4,685.00 4.69 2.83 5.90 6.34 6.34 3.74 Viscosity cP 0.02 0.02 0.02 0.012 0.015 0.015 0.015 0.012

*** LIQUID PHA5E *** Density kg/cum 869.51 944.18 941.97 Viscosity cP 0.28 0.49 0.55 5urface Ten dyne/cm 26 .59 42.46 L...... 42.39

.g

Item Number: 5-306 5-401 5-401a 5-402 5-403 5-501 5-501 a 5-502 5-503 5-504

Temperature (oF) 113 159.4 159.4 77 152.3 120.9 121 .8 101.8 98.3 101.81 Pressure (psia) 244.7 233.5 233.0 214.7 212.7 229.9 43.5 32.0 30.1 32 .0, Total Mass Flow (Ib/hr) Components (Ib/hr):

39,509.77 19,877.11 19,618.71 785.61 20,404.31 105,380.47 105,379.58 10,354.66 2,463.56 2,885.40

HYDROGEN 0.02 7,088.00 6,995.85 785.61 7,781.46 0.04 0.04 trace trace 0.04 CO2 97.73 5,205.19 5,137 .53 5,137.53 220.69 220.69 5.53 1.30 215.16 METHANE 21.93 1,514.82 1,495.13 1,495.13 50.38 50.38 2.39 0.54 47 .99 ETHYLENE 22.87 866 .63 855.36 855.36 50 .01 50.00 8.07 1.88 41 .93 ETHANE 22.49 866.69 855.42 855.42 54.41 54.41 4.78 1.09 49 .63 ACETALD 6,677.43 631.81 . 623.59 623.59 12,846.44 12,843.29 10,306.52 2,455.29 2,527.19 ACETONE 546.90 289.14 285.38 285.38 1,162.51 1,162.71 trace trace trace ETHYLACE 781 .38 918 .35 906.41 906.41 2,764.63 2,765.36 trace trace ETHANOL 738.33 17.47 17.24 17.24 865.07 865.08 trace trace WATER 6,935.17 946.23 933.93 933.93 19,807.78 19,808.00 27.372 3.456 3.46 HOAC 23,665.52 1,532.78 1,512.85 1,512.85 67,558.51 67,559.63 trace trace

*** VAPOR PHASE *** Density kg/cum Viscosity cP

2.85 0.012

2.85 0.012

1.19 0.009

2.48 0.012

3.48 0.012

2.48 0.013

3.71 0.0098

**- LIQUID PHASE --Density kg/cum 926.34 936.37 935.77 747.63 750.34 Viscosity cP 0.56 0.56 0.55 0.19 0.19 Surface Ten dyne/cm 41 .01 41 .92 41.85 18.96 19.10

::0

Item Number: 5-505 5-506 5 -506a 5 -507 5-601 5-602 5-603 5 -604 5 -701 5-702

Temperature (oF) 10 262 .7 262 .6 158 262.7 263.2 117.5 118.1 275.2 182.5 Pressure (psia) 28.0 37.8 37.8 31.8 37.8 100.0 35.0 98 .0 95.0 40.0 Total Mass Flow (Ib/hr) 421 .84 92,139.53 62,277.57 62,277 .91 29,861 .62 29,861.62 22,326.31 22,326.31 31,444 .17 5,111 .13 Components (Ib/hr): HYDROGEN 0.041 trace trace trace trace trace < 0.001 < 0.001 < 0.001 < 0.001 CO2 213.86 trace trace trace trace trace trace . trace trace trace METHANE 47.45 trace trace trace trace trace ETHYLENE 40.05 trace trace trace trace trace ETHANE 48.54 trace trace trace trace trace ACETALD 71.90 9.57 4.87 6.47 3.10 3.10 17.83 17.83 20.93 4.82 ACETONE trace 1,162.71 785.95 785.89 376.83 376.83 1,598.79 1,598.79 1,975.61 487.45 ETHYLACE 4,106.58 2,775.97 2,775.68 1,330.91 1,330.91 17,200.99 17,200.99 18,531 .88 4,044 .88 ETHANOL 163.78 110.71 110.70 53.08 53.08 110.58 110.58 163.66 31.88 WATER < 0.001 20,051.42 13,553.36 13,552 .93 6,498.49 6,498.49 1,746.22 1,746.22 7,878.58 540.20 HOAC 66,645.46 45,046.71 45,046.25 21,599.21 21,599.21 1,651.91 1,651.91 2,873.51 1.90

*** VAPOR PHASE *** Density kg/cum 3.05 Viscosity cP 0.011

*** LIQUID PHASE *** Density kg/cum 868.23 868.12 945.04 868.23 867.80 895.15 894.70 784.92 834 .37 Viscosity cP 0.27 0.27 0.49 0.27 0.27 0.41 0.41 0.27 Surface Ten dyne/cm 34.14 34.23 42 .56 34.14 34.10 34.16 34 .11 32.29

~

Item Number: 5-703 5-704 5-705 5-706 5-707 5-801 5-802 5-803 5-804 5-805

Temperature (oF) 214.1 113 182.3 181.9 184.3 117.5 259.4 234.1 182.1 167.6 Pressure (psia) 39.0 36.0 24 .7 25.0 24.9 35.5 35.0 27.4 25.0 22 .0 Total Mass Flow (Ib/hr) 36,555.16 Components (Ib/hr): HYDROGEN < 0.001 CO2 trace METHANE ETHYLENE ETHANE

36,555.16

< 0.001 trace

5,111.13

< 0.001 trace

4,356.10

< 0.001 trace

755.20

trace

8,295.16

trace

2,161 .83 9,262.68 1,194.31

trace

439.28

trace I

I

ACETALD 25.83 25.83 4.82 4.66 0.17 3.63 0.32 3.31 3.16 ACETONE 2,463.10 2,463.10 487.45 424.01 63.50 439 .38 145.94 293.44 229.99 ETHYLACE 22,576.48 22,576.48 4,044.88 3,434.98 610.00 803.75 19.51 784.24 174.35 ETHANOL 195.55 195.55 31.88 29.15 2.74 55.58 50.45 5.13 2.39 WATER 8,418.79 8,418.79 540.20 462.28 77.92 6,208.52 2,161.83 8,263.04 107.31 29.39 HOAC 2,875.41 2,875.41 1.90 1.02 0.88 784.29 783.42 0.88 trace

*** VAPOR PHASE *** Density kg/cum 5.02 Viscosity cP 0.011

1.33 0.014

3.56 0.01

3.08 0.0096

*** LIQUID PHASE *** Density kg/cum 834.76 910.22 834.57 834.49 835.00 943.08 891 .97 Viscosity cP 0.25 0.50 0.27 0.27 0.26 0.57 0.25 Surface Tenjyne/cm_ 40.45 48.51 32 .31 32 .37 31 .98 64.57 54.45

t:

Item Number: 5-901 5-901a 5-902 5-903

Temperature (oF) 167.6 117.5 221 175 Pressure (psia) 22 .0 35.0 28 .0 16.0 Total Mass Flow (Ib/hr) Components (Ib/hr) : HYDROGEN CO2 METHANE ETHYLENE ETHANE

439.28

trace

5,933.43

< 0.001 trace

1,577.33

trace trace

1,139.31

ACETALD 3.16 4.66 < 0.001 < 0.001 ACETONE 229.99 424.90 0.89 0.89 ETHYLACE 174.35 4,571.37 1,136.39 1,134.92 ETHANOL 2.39 29.39 0.24 0.24 WATER 29.39 464.07 1.79 1.79 HOAC trace 439.05 438.03 1.48

*** VAPOR PHASE *** Density kg/cum Viscosity cP

3.08 0.0096

*** LIQUID PHASE *** Density kg/cum 895.16 853.48 838.96 Viscosity cP 0.41 0.26 0.25 Surface Ten dyne/cm 34.16 16.21 17.39

Process Description 1. Reaction Section

The feedstocks for the reaction of interest are acetic acid and hydrogen.

Hydrogen gas is available at 200 psig, and the acetic is available as a liquid. We assumed

both starting materials to be at 77 OF. These starting materials are each mixed with

recycle streams composed primarily of the respective reactant before proceeding to the

reactor. The pure acetic acid feed stream S-101 is combined with recycle streams S-102

and S-1 03, which originated at the ethyl acetate distillation column DC-91 0 and the

acetic acid distillation column DC-61 0, respectively. In preparation for the high

pressures required for reaction, the acetic acid feed is pumped to 263 psi by the acetic

acid feed pump, P-110 (see Unit Description on p. 54). The pure hydrogen feed, S-402,

is mixed with the mostly-hydrogen recycle stream S-40 1, and this combined stream is

compressed from 213 psi to 263 psi in the compressor CP-410 (see Unit Description on p.

40). This is a very expensive part of the process because of the price of both the

. compressor itself and the high electricity requirement.

The combined acetic acid stream S-201 is passed through heat exchanger HX-200

(see Unit Description on p. 50), where hot reactor effluent in stream S-206b, heats the

acetic acid from a liquid at 242 OF to a partial vapor at 465 OF. In the process, the reactor

effluent is cooled to its dew point of280 OF. A different portion of the reactor effluent,

stream S-206a, is used to heat the hydrogen feed S-207 from 203 OF to 478 OF in the heat

exchanger HX-21 0 (see Unit Description on p. 51). The cooled reactor effluent S-208

exiting HX-21O is also at its dew point of280 OF and 246 psi, this was done for ease of

mixing when two reactor effluent streams are reunited. The hot acetic acid and hydrogen

23

streams, S-203 and S-204 respectively, are both fed to the fired heater F-230 (see Unit

Description on p. 49), which raises their temperature to 599F. Optimal reaction

conditions occur at approximately 570 of, but the reaction is endothermic, so heat would

have to be supplied in order to maintain a constant temperature. Instead, we decided to

heat the reactants to a higher temperature rather than attempt to insulate the reactor. The

feed temperature of 599 of is well within the range of suggested temperatures for this

reaction (Testin, et al.). F-230 is designed to operate on natural gas, but waste material

streams OFFGAS and ACETONE WASTE are rich in hydrogen and hydrocarbons and

are also burned in F-230. These streams are used to furnish 8,547,300 Btu/hr of the

15,291,600 Btu/hr required to operate the furnace . More energy could be taken from

these streams as well as the purge streams, but limiting the amount of energy derived

from waste streams to approximately 55% makes controlling the heating rate of the

furnace more reliable. Any portion ofthese streams that is not used in the furnace is sent

to the flare stack to be burned.

The reactor RX-240 (see Unit Description on p. 59) is a cylindrical vessel

containing a packed bed of20% wt. Pd- Fe203 catalyst pellets. In order to ensure the

proper oxidation state for the desired conversion, the hydrogen and acetic acid are fed in

a 5/1 molar ratio. Conversion of acetic acid is only 46 %; this is in order to improve the

selectivity to acetaldehyde, which is 86 % under the given conditions. Side products

formed in the reactor include ethanol, acetone, carbon dioxide, and light hydrocarbons.

The reactor effluent S-206 is at 557F, and contains, by mass, 11 % hydrogen, 21 %

acetaldehyde, and 38 % acetic acid.

24

2. Acetaldehyde Purification

The reactor effluent S-206 is then split into two streams and used to preheat the

acetic acid and hydrogen feeds. It is split instead of passing sequentially into the heat

exchangers so that both the hydrogen and acetic acid feeds can be heated to higher

temperatures. After passing through the heat exchangers, the temperature of the

combined stream is 280 of, its dew point. This is hot enough to supply energy

sequentially to the reboilers of the azeotropic distillation column (DC-900) and the

acetone waste column (DC-81 0), but the amount of steam and cooling water utilities

saved would not counteract the need to move the hot fluid over long distances and the

associated control complications.

After passing through the acetic acid heat exchanger HX-200 and the hydrogen

heat exchanger HX -210, the separate reactor effluent streams are mixed together, forming

stream S-301. This stream must be cooled further in order to achieve high recovery in the

absorber column; cooling water is used to cool the stream to 113 of in HX-300 (see Unit

Description on p. 51). This partially condenses the stream, and the cool effluent, S-302 is

fed to the flash vessel FV -310 (see Unit Description on p. 50) to separate the liquid and

vapor phases. Only the vapor stream S-303 exiting the flash vessel is sent to the absorber

AB-320 (see Unit Description on p. 39), which it enters on the bottom stage. The solvent

fed to the top stage of the absorber is the acetic acid-rich bottoms product from

acetaldehyde distillation column DC-500 (see Unit Description on p. 44). Under the

conditions of high pressure (the top stage operates at 233.5 psi), this solvent

preferentially absorbs the acetaldehyde, allowing hydrogen, carbon dioxide, and other

25

light materials to escape: 85 % of the acetaldehyde is recovered, but only 3.5% of the

ethane (the next-heaviest component) is recovered.

The amount of material recycled to the absorber is an important variable in the

economics of the process. As the recycle amount increases, the recovery of acetaldehyde

increases, but at the expense of larger diameters for both the absorber and the distillation

column, as well as a more difficult separation of the acetaldehyde from the water and

acetic acid. The selected solvent recycle allows recovery of 95% of the acetaldehyde,

while also leaving a feasible separation and relatively low column costs. The pressure of

the absorber is also as high as possible in order to improve acetaldehyde recover and limit

the amount of work the compressor CP-410 must do. The mainly-hydrogen vapor exiting

AB-320 is recycled to the reactor, with a 1.3% purge taken to prevent excessive buildup

of light components. This purge stream PURGE2 is burned in the flare stack. The

bottoms product from the absorber, S-305, is combined with S-306, the liquid exiting the

flash vessel FV-310. S-305 and S-306 contain 10% and 17% acetaldehyde by mass,

respectively. At the time of mixing, the pressure is still high, 230 psi. The combined

stream S-50 1 is passed through a valve to reduce the pressure to 43 psi before being fed

to the acetaldehyde distillation column DC-500.

Design of the distillation column to recover acetaldehyde was complicated by

several factors. The boiling point of acetaldehyde at atmospheric pressure is 70 OF, so it

is preferred to keep the pressure elevated to reduce the need for refrigerants to condense

the vapor. However, acetaldehyde fonns an azeotrope with water at pressures higher

than 30 psi, and this becomes more water-rich with increasing pressure. In addition, the

presence oflight components such as carbon dioxide and methane require adjustments to

26

be made in the distillation. First, they lower the boiling point ofthe mixture further,

making it more difficult to use cooling water to condense the distillate. Second, their

concentration is high enough that even if the acetaldehyde is completely separated from

the heavier components, its purity will still be only 97.2% wt., which is not high enough

to be sold. Thus, this column requires a partial condenser with both liquid and vapor

distillate: the liquid is taken as the acetaldehyde product, while the vapor distillate is fed

to another column where additional acetaldehyde can be removed from the lighter

components. Several alternatives were considered before choosing to operate the

condenser at 32 psi. Since the liquid-liquid separation of acetaldehyde and water actually

increases with temperature, one alternative examined carrying the high pressure from the

absorber into the distillation column, then allowing the two liquid phases to separate in a

decanter. This method did not produce acetaldehyde at the required purity, was difficult

to control, and required very high reboiler temperatures. A second alternative involved

using the main column as selected, but also adding an absorbing column, in which the

acetic acid-rich bottoms product would be used to absorb acetaldehyde. This did not

produce sufficient separation, and the added flow into the distillation column increased

the column's cost.

The selected design involves operating the condenser of the acetaldehyde

distillation column at 32 psi. The distillate vapor fraction was varied to keep the

temperature of the distillate at 102 OF, which can be cooled using cooling water. The

liquid product, HAC PRODUCT, contains 10,355 Iblhr (82,012,000 Ib/yr) of99.5% wt

acetaldehyde. The vapor distillate S-504 is 22 % by mole of the total distillate and

contains 2,527 Iblhr of acetaldehyde. Since this is a significant portion of the product, it

27

was necessary to recover as much of this stream as possible. The vapor distillate is then

sent to the bottom stage of a small secondary column, DC-51 0 (see Unit Description on

p. 46), which has a refrigerated condenser at 10 of. A second product stream, HAC

PRODUCT, exits the bottom stage ofDC-510 and contains 24641blhr (19,515,000 lb/yr)

of99.7 wt% acetaldehyde. Sixty tons of refrigeration is required at the condenser, and

because of the temperature of 10F is fairly moderate, this can be provided by an

ammonia absorption system RF-520 (see Unit Description on p. 64). Decreasing the

temperature further would increase the yield of acetaldehyde, but at the expense of more

expensive equipment and a larger heat duty. In this design, 71.9 Iblhr (569,000 lb/yr) of

acetaldehyde is lost to the stream OFFGAS. This stream, with its high levels of methane,

ethane, and ethylene is burned in the fired heater, F-230, to reduce natural gas costs.

The acetaldehyde distillation column is also the location of an esterification

reaction between ethanol and acetic acid, which forms ethyl acetate and water.

Equilibrium for this reaction is achieved wherever acetic acid and ethanol are present

together, but for the purposes of this design, it was assumed this reaction occurs only in

the bottom stage ofDC-500, where the high temperatures and the presence of acetic acid

and ethanol in the liquid phase especially favor this reaction. When equilibrium is

reached, over 80% of the ethanol has been reacted. The presence of ethyl acetate is very

important in the acetic acid separations section of the process, where ethyl-acetate forms

an azeotrope with water, easing the separation of water from acetic acid. The bottoms

product S-506 from the acetaldehyde distillation column is split, with part proceeding to

the acetic acid separation sequence, and the remainder being cooled by cooling water in

the heat exchanger HX-530 (see Unit Description on p. 52) and then recycled to the top

28

stage of the absorber AB-320, where it acts as the solvent to preferentially absorb

acetaldehyde.

3. Acetic Acid separation

The main goal of the acetic acid separation column DC-610 (see Unit Description on

p. 46) is to obtain a pure stream of acetic acid, which can be recycled to the reactor feed.

The reasons for this are twofold. The primary reason is that the high cost of acetic acid

makes it economically feasible for us to reuse the unreacted acetic acid rather than

dispose of it. This is particularly relevant because of the low conversion in the reactor,

which results in a significant amount of unreacted acetic acid in the system. The second

reason is that acetic acid is an impurity in water and its substantial presence in the

wasteYv'ater stream will increase costs of wastewater treatment.

There are two streams entering the acetic acid distillation column. The feed stream, S

602, is the bottoms from the acetaldehyde separation columns and enters DC-61 0 at the

1i h of 30 actual trays. It primarily consists of the unreacted acetic acid and the products

of side reactions such as water, ethyl acetate, acetone, and ethanol. The recycle stream S

603 fed to the second stage is rich in ethyl acetate and is used to facilitate the low boiling

water - ethyl acetate azeotrope, which makes the separation of acetic acid in the bottoms

easier. The feeds S-602 and S-603 are pumped to 100 psi and 98 psi, respectively, before

entering the column. The condenser is operated at 95 psi because as pressure increases,

the water-ethyl acetate azeotrope becomes more water-rich, easing the separation from

acetic acid. This significantly increases the purity of acetic acid that is collected at the

bottoms. lfthe same process were operated at 37.8 psi (the pressure of the feed stream),

then the mass fraction is only .92-.93 as opposed to .982 at this pressure. Though the

29

higher pressure increases the cost, it is a cost that is well incurred since the acetic acid

increase in the bottoms is critical as it lowers the acetic acid that is lost in the wastewater

stream. Further increases in pressure beyond 95 psi cause very marginal increases in the

acetic acid mole fraction and do not justify the additional costs.

The nwnber of equilibrium stages calculated is 18. The increased separation with

additional stages is minimal after 18 stages and does not justify the increasing cost. The

calculated tray efficiency is 61 %, meaning 30 actual trays are required. We use a kettle

reboiler in this process and a total condenser. A total condenser is used because the

distillate must be fed to the decanter DE-720 (see Unit Description on p. 44) as a liquid.

Before the distillate enters the decanter, it goes into a mixer, M-700. The mixer

incorporates the distillate stream of the acetic acid separation colwnn with the distillate

stream of the DC-900 col umn, whiCh primarily consists of ethyl acetate and water, and

the bottoms of the acetone separation column, DC-81 0, which contains ethyl acetate that

was contained in the water-rich stream coming out of the decanter and water and acetone

impurities. The pressure across the mixer drops to 39 psi. This is done because operating

the remaining distillation columns in the separation sequence at higher pressure does not

produce results that are significantly favorable enough to account for incurring the higher

cost when the separating columns were operated at a higher pressure. In addition, this

places less of a load on the pump P-730 (see Unit Description on p. 56), which must

increase the pressure of recycle stream S-705 to only 40 psi before sending it to the

mIxer.

The exit stream from the mixer has a temperature of214 of and is cooled to 113 of

by passing it through the heat exchanger HX-71 0 (see Unit Description on p. 52). The

30

exit stream from this heat exchanger is fed to the decanter DE-720, which is used to

separate the water from the ethyl acetate in the feed stream. The water rich stream S-801

contains 74.8% by mass of water and S-901 contains 77.0% ethyl acetate.

4. Ethyl Acetate Separation

The ethyl acetate-rich organic layer exiting the decanter in stream S-90 1 is sent to the

ethyl acetate splitter, which recycles 78.9% of the stream to the acetic acid distillation

column where it enhances the separation of water from acetic acid. The rest of the stream

is fed into the distillation column DC-900 (see Unit Description on p. 48).

Column DC-900 is used to separate water from the acetic acid and ethyl acetate in

S-90 1. The low-boiling azeotrope of water and ethyl acetate is taken in the distillate

along with acetone and acetaldehyde, leaving a bottoms product that is primarily ethyl

acetate and acetic acid. The acetic acid is purified and recycled to the reactor, and the

ethyl acetate is sold as a product. The distillate stream of this distillation, S-706, is

recycled to the decanter DE-720. This is to separate the water from the ethyl acetate,

which is then either recycled to the acetic acid distillation column, or purified for sale.

The separation in DC-900 was achieved by using a condenser pressure of 30 psi and 12

equilibrium stages (20 actual trays). A total condenser and kettle reboiler are used in this

process.

The bottoms product of DC-900, S-902, is 72% ethyl acetate by mass, with acetic

acid composing most of the balance. This stream enters the ethyl acetate distillation

column, DC-91 0 (see Unit Description on p. 48), at stage 13 of23 (tray 25 of 43). The

main purpose of this column is to separate the ethyl acetate from the acetic acid at a level

31

of purity whereby the ethyl acetate can be sold. This column is operated at a condenser

pressure of 16 psi and the distillate contains 99.6% mass of ethyl acetate. The bottoms

product contains 99.7% by mass of acetic acid and is recycled back to the reactor via the

stream S-1 02. Ethyl acetate is stored in a cool, dry, well-ventilated area with a holding

capacity of 7300 ft3 (enough capacity to hold 14 days worth of ethyl acetate production)

in tightly sealed containers.

5. Acetone and Wastewater Disposal

The water-rich stream S-SO1 exiting the decanter is then sent to the stripper ST -SOO

(see Unit Description on p. 66). The condenser is operated at a pressure of25 psi. The

bottoms stream of this column is sent to a waste treatment plant for purification before it

is disposed off. The distillate, S-S04, primarily contains acetone and ethyl acetate and is

fed to the acetone separation column, DC-Sl 0 (see Unit Description on p. 47), at stage S

of 11 (13 th of IS actual trays). A partial condenser is used for the stripper because utility

costs are decreased by not condensing the vapor and feeding a dew point vapor to the

acetone distillation column.

The acetone distillation column, DC-S1 0, aims to remove all of the acetone from the

feed stream as the distillate and remove it from the system. The acetone in the distillate

cannot be made pure enough to be sold as a side product, unlike ethyl acetate. This

column has IS actual trays and operates at a condenser pressure of 22 psi. The bottoms

product S-707 mainly contains ethyl acetate and is mixed with S-706, the distillate stream

of the DC-900 column, via the mixer M-740. The combined stream, S-705, is at 24 psi

32

and is pumped to 40 psi before being fed to the mixer M-700 where it is mixed with the

acetic acid separator distillate to be fed into the decanter.

33

..

Energy Balance and UtiJity Requirements

Because of the high temperatures required for the reactor and the many dis61lation

columns, supplying energy for heating and cooling process streams is of paramount

concern for the economics of the process. The largest heating and cooling requirements

are found in the acetic acid distillation column, which requires 39,534,800 Btu/hr for the

reboiler at 370 of and 36,586,600 Btulhr for the condenser. These requirements are

satisfied with 300 psia stearn (dropped from its source at 600 psig) and cooling water,

respectively. The reboilers and condensers for each of the other columns are handled

similarly (using steam at the appropriate pressure), except for the low-temperature

acetaldehyde condenser in DC-51 O. The refrigeration unit RF-520 utilizes ammonia

absorption to cool a 40% ethylene glycol in water solution that circulates in the

condenser. The glycol removes 612,900 Btulhr of heat by partially condensing the

distillate ofDC-5IO at 10 OF. Because of heat leak to the surroundings and inefficiencies

in heat transfer, we assumed that the ammonia absorption system must supply 900,000

Btu/hr of refrigeration. Assuming the ammonia system operates at -10 OF, 507

Btu/min/ton are required for steam in the generator and 5.4 gpm/ton of water are required

for the conditioner (McKetta). The refrigeration load is 75 tons, requiring 2,497,500

Btulhr of steam at 300 OF (68 psi) and 24,000 gpm of cooling water.

The acetic acid and hydrogen streams, S-20I and S-207 respectively, which are

the reaction starting materials must be heated from 242 OF and 203 OF to 599 OF before

entering the reactor. This is above the optimal reaction temperature of 572 OF, but still

within the range of recommended temperatures (Tustin, et.a!.). It was heated to this

temperature because the primary reaction is endothermic, and by heating above the

35

optimal temperature we allow the temperature in the reactor to decrease as the reaction

progresses. This method was used because it was suggested that this would be more

efficient than attempting to insulate the reactor at the high temperature required. The

reaction consumes 2,677,000 Btu/hr, and the effluent is a total vapor at 557 OF. The hot

reactor effluent is split and sent to separate heat exchangers to heat the hydrogen and

acetic acid feed streams. The rate of energy transferred to the acetic acid in HX-200 is

4,305,300 Btu/hr, and 8,741,000 Btulhr is transferred to the hydrogen stream in HX-21 O.

Following this, the hydrogen and acetic acid feed streams still require 12,692,000 Btu/hr

to reach 599 OF; this is accomplished in the fired heater, F-230. Waste streams OFFGAS

and ACETONE WASTE are burned to produce 8,547,300 Btulhr, and natural gas is

required for the rest of the duty, which because of inefficient heat transfer is 15,291,600

Btulhr total. After being cooled to its dew point by the reactor feeds, the reaction product

S-401 is cooled further, to 113 OF by transferring 20,300,000 Btulhr to cooling water in

HX-300.

The option of using heat from the reactor effluent stream S-401 at 280 OF to heat

the reboilers ofDC-900 and DC-810 was examined, but ultimately rejected. There is

sufficient energy in S-401 to maintain a 45 OF driving force with the bottoms, but

substituting this method for steam heating did not justify the need to pump the hot fluid

over long distances and the more complicated control aspects. Heat integration among

the condensers and reboilers in the separation section was not attempted because the

operating pressures needed to optimize product composition and column costs does not

allow for productive stream matching.

36

The utility requirements are summarized in the table below.

Table 2: Heat Transfer Among Process Streams Cold Stream Cold Stream

Temperature Change

Hot Stream Hot Stream Temperature Change

Energy Transferred (BtU/hr)

S-201 242F to 465F S-206b 557F to 280F 4,305,300 S-207 203F to 478F S-206a 557F to 280F 8,741,000

Table 3: Cooling Water Requirements Process Stream or Condenser

Temperature Cooling Water Duty (Btu/hr)

Amount of Cooling Water Required (Gallhr)

Cost ($/ru:)

C-500 102 of 10,559,800 85,332 28.16 C-610 275.2 of 36,586,600 147,824 48.78 C-800 182F 994,200 4,020 1.33 C-810 184 of 428,700 1,732 0.57 C-900 182F 2,911,400 11,763 3.88 C-910 175F 509,600 2,059 0.68

RF-520 1,503,600 24,300 8.02 S-301 280 of to 113 of 20,300,000 82,020 27.07

S-506a 263F to 158F 2,736,000 11,050 3.65 S-703 259 of to 113 of 3,689,100 14,906 4.92

TOTAL REQUIREMENTS 80,219,000 385 ,006 $127.06

Table 4: Steam Utility Requirements Reboiler Temperature

of Stream Pressure of Purchased Steam (psig)

Steam Heat Duty (Btulhr)

Amount of Steam Required (lblhr)

Cost ($Ihr)

R-500 262.7 of 75 16,892,000 14,195 35.49 R-610 369.9 of 600 39,534,800 29,636 82.98 R-810 184.2 of 35 252,312 215 0.53 R-900 221.2 of 35 3,150,280 2,683 6.60 R-910 262.9 of 75 486,919 409 1.02 RF-520 300 of 75 2,497,500 2,099 5.25 TOT AL REQUIREMENTS \62,813,911 ~238 $132.00

37

Table 5: Natural Gas Requirements T JDota uty 0 fF dBIre eater (BtU/hr) 15,291,600 Energy from Waste Streams (Btu/hr)

OFFGAS 3,604,800 WASTE ACETONE 4,942,500

Natural Gas Required (Btulhr) 6,744,300

Table 6: Electricity Requi.ements Process Unit Electrici ty

Required (kW) Cost ($/hr)

CP-410 459.6 16.086 P-110 18.7 0.654 P-540 20.5 0.718 P-600 3.8 0.134 P-620 3.1 0.109 P-730 0.3 0.010 PB-500 7.S 0.262 PB-610 1l.2 0.392 PB-810 0.1 0.003 PB-900 0.4 0.013 PB-910 0.3 0.010 PR-SOO l.5 0.OS2 PR-SlO 0.2 0.008 PR-610 5.6 0.196 PR-800 0.3 0.011 PR-810 0.1 0.003 PR-900 1.1 0.039 PR-9l0 0.1 0.003 Total Requirements 534.1 $18.69

The costs listed are what would be paid if these utilities were purchased from an

outside source. This was not done for steam and cooling water, because it was

determined that building allocated facilities would be more profitable (see Appendix, p.).

38

Unit Descriptions

Absorber

AB-320 (see spec. sheet p. 72)

The absorber unit AB-320 is a trayed tower used to separate the acetaldehyde

from light hydrocarbons, carbon dioxide, and the unreacted hydrogen that results from

feeding that material in large excess. The primary objective of this unit is to recover as

much acetaldehyde as possible because 1.3 % of the material leaving through the top of

the column leaves the system as PURGE. 23,470 Iblhr of the vapor stream S-303 is fed

to the bottom stage of the column, and 62,278 lb/hr of S-304, the solvent recycled from

the bottoms of the acetaldehyde distillation column, is introduced on the top stage. This

solvent level was chosen to balance the amount of acetaldehyde recovered in the absorber

with the ease of separation of the acetaldehyde from the remaining liquid, which contains

a lower fraction of acetaldehyde as solvent flow increases. The recovery stream S-305

exits as a liquid from the bottom stage, and has a total flow of 65,871 lb/hr with an

acetaldehyde mass fraction of 0.094. S-401 exits the top of the column and contains

mostly unreacted hydrogen and light gas side products; 98.7% of this stream is recycled

to the reactor, the remainder is purged and burned in the flare stack.

The top stage pressure in the absorber is 233.5 psi, with the intention of keeping

the pressure as high as possible without resorting to compressing the reactor effluent.

The absorber was designed to recover over 85 % of acetaldehyde fed to it, so that the

resulting loss to purge would be only approximately 0.15 % of acetaldehyde. This design

led to a column containing 15 theoretical trays; the O'Connell correlation was used to

calculate the stage efficiency (Seader). This was found to be 51 %, requiring 30 actual

39

trays. Using equations and tables found in Seider, the calculated dimensions were 2.5 ft

diameter and a height of 74 ft; the associated bare module cost of the tower and trays is

$117,300. IPE calculated a 3 ft diameter, 76 ft. height, and an equipment cost of

$120,000. The material used for both the column and the trays is stainless steel because

of the concern of corrosion caused by acetic acid.

Compressor

CP-41O (see spec. sheet p. 73)

Compressor CP-41O is a major piece of equipment because of the large expense

associated with compressing gases, especially hydrogen. A reciprocal compressor is used

to compress 20,404 felhr of the mixed hydrogen stream S-207 from 212.7 psi to 262.8

psi. The temperature of the stream also increases, from 152.3 OF to 202.6 oF. The load

required was limited as much as possible by maintaining high pressures throughout the

reactor section of the process, while still ensuring that the feed to the compressor was

above its dew point. The power is 330.9 kW, and the efficiency is estimated by Aspen to

be 72%. This leads to an electricity requirement of 459.6 kW. The material used was

stainless steel, and the bare module cost was calculated to be $2,628,000.

Condensers

C-500 (see spec. sheet p. 74)

Condenser C-SOO is a partial condenser that is used for the acetaldehyde

distillation column, DC-500. A partial condenser with both liquid and vapor distillate is

employed for two reasons. First, the significant concentration of light compounds such

as carbon dioxide would cause the acetaldehyde purity in the product to be too low if a

40

single disti llate stream was taken. Second, condensing those light gases and all of the

acetaldehyde would require very low temperatures and a large amount of refrigeration.

Instead, the fraction of vapor distillate was set in order to have a distillate temperature of

102 of, which can be achieved with cooling water. To accomplish this, 22% of the

distillate remains vapor; this stream (S-504) contains approximately 20% of the

acetaldehyde produced, and this is recovered in the DC-51 0 column. The condenser is

made of stainless steel, and has a heat duty of 10,559,800 Btulhr. Using a heat transfer

coefficient of97.6 BtU/(hr-ft2-0F), B-JAC estimates a surface area of 13,109 ft2; because

of this large size and the temperature crossover of the hot and cold stream temperatures,

B-JAC split this exchanger into two in series. A 1-8 shell-and-tube heat exchanger was

designed, with a length of 8 ft., and shell size of 70 in. The estimated installed cost from

B-JAC is $209,960.


C-520, a partial condenser with all vapor distillate, is used for DC-51 0, the

refrigerated acetaldehyde recovery column. If a liquid distillate were condensed, the

temperature requirement would be unreasonably low, and there is little processing benefit

from having a liquid distillate. The heat duty of the condenser is only 612,900 Btu/hr at

10F, making refrigerant cooling a practical solution for this process. A 40%-ethylene

glycol in water solution is circulated in the condenser to cool the distillate. The heat

transfer coefficient estimated by B-JAC is only 32 .2 Btu/(hr-ft2-0F), and the required

surface area is 403.2 ft2. In order to limit the amount of glycol solution needed to cool

the condenser, the solution is heated from -5 OF to 30 OF; because of the temperature

41

crossover, two condensers, both 1-2, are required in series. The length ofthe tube is 14

ft., and the shel1 width is 10.75 in. The material used was stainless steel. The estimated

installed cost reported by B-JAC was $16,600.


The condenser used for the acetic acid distillation column is a total condenser,

because in the next step in the process the liquid distillate is fed to the decanter DE-720.

Cooling water is used, and the exiting distillate has a temperature of275 of. The heat

duty is 36,586,600 Btu/hr, and using an estimated heat transfer coefficient of 129 BtU/(hr

ft 2_OF), the required surface area is 2,494 ft2. A single 1-4 shell-and-tube heat exchanger

is used for this process. The tube length is 10ft., and the shell diameter is 38 in. Using

stainless steel as the material, B-JAC estimated the installed cost of this item to be

$64,710.


For the stripper, a partial condenser was used because sending S-804 to the

acetone distillation column DC-81 0 as a vapor reduced the overall utility requirements

without significantly altering the separation. The condenser C-800 is a 1-4 shell-and-tube

heat exchanger, made of stainless steel. The required heat duty is 994,200 Btu/hr, and

using a heat transfer coefficient of 112 Btu/(hr -ft2 -OF), the required surface area is 112.2

ft2. The shell diameter is 8.6 inches and the tube length is 16 ft. The installed cost is

$7,570.

42


The condenser for the acetone distillation column DC-81 0 is a 1-4 shell-and-tube

heat exchanger made of stainless steel. The heat duty is 428,700 Btu/hr, with a distillate

temperature of 168 of. A partial condenser with all vapor distillate was used because the

distillate stream WASTE ACETONE is being burned in the fired furnace, so it is

unnecessary to condense the stream. The heat transfer coefficient is 95 Btu/(hr-ft2_OF),

and the resulting area is 78 ft2. The tube length is 10ft., and the shell diameter is 8.6 in.

The installed cost is $6,700.


A total condenser is used for the near azeotrope distillation column DC-900. The

heat exchanger is a 1-2 shell-and-tube heat exchanger made of stainless steel. This

condenser subcools the product to 2 of below the saturation temperature so that when the

disti llate is mixed and fed to the pump P-730, the feed stream is a total liquid. The

distillate temperature is 18l.9 of, and the heat duty is 2,911,400 Btu/hr. Using a heat

transfer coefficient of 105 Btu/(hr-ft2 _OF), B-JAC estimated the required surface area to

be 275 ft2. The tube length is 14 ft., and the shell thickness of 12.75 in. The installed

cost is $9,250.


A total condenser is used for the ethyl acetate distillation column so that the liquid

ethyl acetate can be recovered and stored. C-910 is a 1-8 shell-and-tube heat exchanger

made of stainless steel. The distillate temperature is 175 of, and the heat duty is 509,600

43

Btulhr. The effective surface area based on an estimated heat transfer coefficient of72.6

Btul(hr-ft2 _OF) is 473 ft2. The tube length is 6 ft., and the shell diameter is 24 in. The

installed cost is $18,470.

Decanter

DE-720 (see spec. sheet p. 81)

The purpose of the decanter is to separate water from the ethyl acetate in the inlet

stream S~704. It achieves this to an extent of getting a water-rich stream, S-80 1 with

74.8% by mass of water and an ethyl acetate rich stream, S-901 with 77 .0% by mass of

ethyl acetate. The outlet temperature of the decanter is 117 .5F and its outlet pressure is

35.5 psi . The capacity of the decanter is based on a 10 minute residence time at half full

and equals 214 ft3. The decanter is a horizontal, stainless steel vessel, with a diameter of

4.5 ft. and length of 13.5 ft; using cost charts, we determined its bare module cost to be

$63,000.

Distillation Columns

DC-500 (see spec. sheet p. 82)

DC-500 is the major unit for separating acetaldehyde from the heavier

components in the reactor effluent: acetic acid, water, acetone, ethyl acetate, and ethanol.

Even though acetone's boiling point is closest to acetaldehyde's, the key heavy

component in this separation is water because of its high concentration. At pressures as

low as 30 psi, water forms an azeotrope with acetaldehyde, which becomes richer in

water as pressure increases. As a result of this, the pressure of the condenser was set at

32 psi to limit the recovery ofthis azeotrope. Acetaldehyde's relatively low boiling point

44

(10F) makes this low pressure undesirable because of difficulties in condensing the

material, but attempted higher pressure methods such as liquid-liquid separation from

water could not produce acetaldehyde in the needed purity. Fenske-Underwood-Gilliland

calculations provided the minimum number of stages as 20 and the minimum reflux ratio

as 1.61. At 1.3 times the minimum reflux (LID = 2.09), the Gilliland correlation yields

the estimated number of stages as 40. This actual conditions used in the simulation were

a reflux ratio of 2.40, and 40 equilibrium stages. The tray efficiency is 56%, necessitating

71 actual trays. Stream S-501 enters above the 50th stage to provide a larger rectifying

section, allowing better purification of the acetaldehyde distillate. Using cost charts

available in Seider, the distillation column dimensions are 4.5 ft in diameter, and 157 ft.

taB; the estimated cost of the column and trays was $1,013,500. IPE estimated the

column dimensions as a 4 ft. diameter and 171 ft. height. The material and labor cost

using lPE was estimated as $1,087,000. The height was determined by taking a 2-f1. tray

spacing, a 4-ft. disengagement height for the condenser, and a 10-ft. bottoms sump.

DC-500 also has the esterification reaction in which acetic acid and ethanol react

to form water and ethyl acetate (Reaction 5, p. 10). This reaction is very important to the

process, and the patent describes adding sulfuric acid to catalyze this reaction if it does

not occur in sufficient yield. Under the conditions of this column, that step was not

necessary. This reaction is important because ethyl acetate is needed in the acetic acid

separation section to form a low-boiling azeotrope with water in order to facilitate the

separation of water from acetic acid. The ethyl acetate-water product is favored

thermodynamically, and under the reaction conditions over 80% of the ethanol reacts.

45

The ethyl acetate can then be purified, and we are able to purify and sell 1,139 Ib/hr of

ethyl acetate.

HAC PRODUCT, the liquid distillate, contains 99.5 % acetaldehyde by mass at a

flow of 10,355 Ib/hr. The bottoms product, S-506, contains 92,140 Iblhr of 72.4 %wt

acetic acid, with water being the other major component. Only 9.6 Ib/hr of acetaldehyde

is in this stream. S-506 is split, with part returning to the absorber AB-320 as the solvent,

and the remainder proceeding to the acetic acid separations section, where pure acetic

acid is recycled to the reactor and ethyl acetate is purified for sale.


The column DC-51 0 contains only a rectifying section and is intended to

condense acetaldehyde from the vapor distillate of the main acetaldehyde distillation

column. S-504 enters the column on the bottom stage, from which the second

acetaldehyde product HAC PRODUCT is also taken. The temperature of the condenser

is 10F, and indirect refrigeration is provided via an ammonia absorption-ethylene glycol

system. Theoretically, the number of stages required to recover pure acetaldehyde in

good yield is only three, but low tray efficiency (26 %) caused by the large relative

volatility of ethane to acetaldehyde leads to the need to for eight actual trays. The height

of the column is 30 ft, and its diameter is 1.5 ft. Stainless steel is used for the tower and

trays. The bare module cost of the column, using cost charts, is $62,500.


The acetic acid separation column was designed for 18 equilibrium stages. The

efficiency is 61 %, and the actual number of trays is 30. Assuming a 2-ft. tray spacing, as

46

well as the sump and disengagement heights, the height ofthe column is 74 ft., and the

diameter is 10.5 ft. The acetic acid-rich stream recovered as the bottoms of the

acetaldehyde distillation column enters DC-610 at tray 17. The ethyl acetate-rich recycle

stream S-603 enters the colurrm on the second tray and is used to form the ethyl acetate

water azeotrope, which makes the separation of acetic acid from water simpler. The

bottoms rate is 20,744 lblhr, the molar reflux ratio is 2.5 and the condenser pressure is 95

psi. On the recommendation of the industrial consultants, stainless steel was the material

chosen for the tower and the trays. This is because acetic acid has a corrosive effect on

carbon steel and stainless steel is sufficiently corrosion resistant. The bare module cost of

this column with trays is calculated to be $1,491,200.


This acetone distillation colurrm is designed to remove the acetone from the system in the

distillate to be used as fuel in the fumace. The distillate stream, ACETONE WASTE,

primarily contains acetone, ethyl acetate and water with very small quantities of

acetaldehyde and ethanol. The bottoms stream, S-707, contains 80.8 % ethyl acetate and

10.3 % water. Eleven equilibrium stages were designed for the column; the tray

efficiency of 60.8 % implies the need for 19 actual trays. Using 2-ft. tray spacing, and 14

ft. total for the sump and disengagement heights, the column is 52 ft. tall, with a 1.5 ft.

diameter. Based on the flow rates a smaller diameter is necessary to avoid flooding, but

the 1.5-ft. diameter produces an aspect ratio of 35, which is much more reasonable. The

stripper distillate stream S-804 enters the column at tray 13. The distillate rate is 439

lblhr, the molar reflux ratio is 4 and the condenser pressure is 22 psi. Stainless steel was

also the material of choice for this column. Its bare module cost is $103,600.

47

....


Th.is column is designed to remove the acetone, ethanol and the ethyl acetate-water

azeotrope as the distillate and isolate the acetic acid and the remaining ethyl acetate in the

bottoms. The ethyl-acetate rich stream S-901 from the decanter is fed at the 10th tray.

Twelve equilibrium stages are needed for the separation, and the efficiency is 60%,

requiring 20 actual trays. To that end, this separation column is 54 ft in height and has a

diameter of 2 ft. The distillate rate is 4356 lblhr, the reflux ratio is 1.8 and the condenser

pressure is 30 psi. Stainless steel was also the materiaJ of choice for this column. Its bare

module cost is $150,532.


This ethyl acetate distillation column is designed to purify ethyl acetate as the distillate to

a level at which it can be sold. The bottoms product is acetic acid, which is recycled to

the reactor. Twenty-three equilibrium stages are needed for this separation; the tray

efficiency is 41.5 %, and the actual number of trays is 45. This separation column is 104

ft in height and has a diameter of 3 ft. Stream S-S-804 enters the column at tray 25. The

distillate rate is 1139 lblhr, the reflux ratio is 1.8 and the condenser pressure is 16 psi.

Stainless steel was also the material of choice for this column. Its bare module cost is

$433,400.

48

Fired Heater

F-230 (see spec. sheet p. 88)

In order to completely heat the hydrogen and acetic acid feed streams, S-204 and

S-203, to the desired reaction temperature, energy must be supplied by burning fuel in the

fired heater F-230. F-230 is a vertical cylindrical fired heater. 20,402Ib/hr of the

gaseous hydrogen feed S-204 at 478 of and 42,571 lblhr of the partial vapor S-203 at 465

OF are fed to F-230 separately because they are different phases. The heat duty required

to raise the temperature of the streams to 599 OF to prepare for the reactor is 12,692,000

Btulhr. Assuming a stack temperature of 650 of, according to McKetta the efficiency is

83% and thus the required total heat duty is 15,291,600 Btulhr. This is furnished by

burning natural gas, along with waste streams from the process: OFFGAS and ACTONE

WASTE. The cost for this unit was estimated based on a design heat duty of 20,000,000

Btu/hr. This leads to an installed cost of $609,400 (Walas). The material used is

stainless steel because it is able to handle the high temperatures and also will not corrode

in the presence of acetic acid. Of the 15,291,600 Btulhr required to heat the process

streams, 8,547,300 Btulhr (56 %) is supplied from burning the process streams. This

amount was chosen in order to balance the cost of natural gas that must be bought with

the control and safety concerns that would be associated with using waste streams for

nearly all of the energy in the heater.

49

Flash Vessel

FV -310 (see spec. sheet p. 89)

The flash vessel FY -310 is used to separate the liquid and vapor portions of the

reactor effluent leaving the heat exchanger HX-300. The volumetric flow rate into the

flash vessel is 101,146 ft3lhr. The required volume, determined by considering a 5

minute holdup time at half full, is 16,860 ft3. This is a large process vessel, and the

diameter was set at 17 ft., so that it would not have to be fabricated on the site. The

length height is 76.5 ft. This is made of stainless steel because of the corrosion caused by

acetaldehyde and acetic acid. The bare module cost for the flash vessel is $1,832,000.

Heat Exchangers

HX-200 (see spec. sheet p. 90)

HX-200 is used to increase the temperature of the acetic acid feed stream S-20 1

from 242 OF to 465 OF, where it is a partial vapor. To do this, 20,783 lb/hr of reactor

effluent S-206b is cooled from 557F to 280F, its dew point. The amount of heat

transferred is 4,305,300 Btulhr, and the estimated heat transfer coefficient between the

vapor and liquid was 10.8 BtU/(hr_ft2 _OF). This is relatively low for what is primarily

liquid-vapor heat exchange. It is a 1-1 shell-and-tube heat exchanger made of stainless

steel. The surface area is 8,343 fe, the tube length is 20 ft., and the shell diameter is 48

in. The cost estimated by B-JAC was $194,840.

so


In HX-210, 20,840 lb/hr of the hydrogen feed stream S-201 at 202.6 of 1S heated

to 477.6 of by cooling 42,1961b/hr of the reactor effluent S-206a from 557 of to 280F.

This temperature is chosen because it is the dew point of S-206a at the operating

pressure, and having only vapor will ease mixing with S-206b, which is also cooled to its

dew point. This is a 1-1 shell-and-tube heat exchanger, which requires three exchangers

in series because of the large surface area required . The amount of heat transferred is

8,741,000, and using B-JAC the heat transfer coefficient was estimated to be 16.7

Btu/(hr-ft2_OF). The total required surface area is 8920 ft2, and the dimensions are a 10ft.

tube length and 60 in . shell diameter. HX-210 is composed of stainless steel to prevent

corrosion. The installed cost estimated by B-JAC was $630,750 for the three in series.

HX-300 (see spec. sheet p . 92)

HX-300 is used to decrease the temperature of 62,979 Iblhr of the cooled reactor

effluent S-301 from 280F to 113F in order to increase the recovery of acetaldehyde in

the absorber AB-320. S-301 enters as a dew point vapor, and is partially condensed in

the heat exchanger. Cooling water is used to transfer 20,300,000 Btulhr from S-301. The

heat exchanger is a 1-8 shell-and-tube heat exchanger made of stainless steel. The

estimated heat transfer coefficient used was 70.9 Btul(hr-fe-OF), requiring a size of6687

ft2. The tube length is 12 ft ., and shell diameter is 58 in. The estimated installed cost for

HX-300 is $163,820.

51


Stream S-506a, the solvent for the absorber, flows at a rate of 62,278 Ib/hr and a

temperature of 263 of into the heat exchanger HX-530. Cooling water is used to lower

the temperature to 157 of because more acetaldehyde is absorbed when the feeds to the

absorber are at lower temperatures. The required heat transfer rate is 2,736,000 Btulhr.

The estimated heat transfer coefficient in B-JAC is 105 Btu/(hr-fe_OF), which appears to

be reasonable for liquid-liquid heat transfer. The required area is 320 ft2, and the

estimated price is $10,750. The material used for the heat exchanger is stainless steel,

and it is a 1-2 shell-and-tube exchanger. The tube length is 20 ft., and the shell diameter

is 10.8 in.


HX-710 employs cooling water to reduce the temperature ofthe S-703 from 260.3 OF

to 113.0F before it is fed to the decanter DE-720. The exchanger has a heat duty of

3,689,100 Btu/hr. A heat transfer area of 921 ft2 is calculated for HC-71 0 in B-JAC based

on an overall heat transfer coefficient of95.4 Btu/(hr-ft2-0F). It is a 1-6 shell-and-tube

heat exchanger with the following dimensions: tube length is 20 ft., shell diameter is 18

in. Stainless steel is the material of construction, and the installed cost is $15,940.

52

Mixers

M-220, M-310

Both M -220 and M-31 0 are used to mix reactor effluent, and a small pressure

drop is assumed across each. M-220 mixes 20,783 lb/lrr ofS-202 and 42,196 lb/hr ofS

208, both at 280 OF and 246 psi, the dew point of the vapors. These streams had been

used to heat the acetic acid and hydrogen feeds . These streams are mixed to allow a

single heat exchanger, HX-300, to cool the streams before feeding the resulting vapor

stream to the absorber AB-320. M-310 mixes 65,871 lblhr of the recovered absorber

bottoms S-305 and 39,510 lblhr of the liquid stream S-306 exiting the flash vessel. There

is a pressure drop of 5 psi across the mixer; this pressure drop can be almost arbitrarily

large, because the mixed stream is then passed into the valve V-50 1 to decrease its

pressure to 45 psi before entering the acetaldehyde distillation column DC-500.

M-400

Mixer M-400 combines the pure hydrogen gas feed at 215 psi a and 77 OF with the

hydrogen-rich stream S-402a at 159 OF and 233 psi. The pressure of the combined

stream is decreased to 213 psia. This pressure is low enough to ensure that the vapor is

above its dew point and no liquid will be fed to the compressor. The outlet temperature is

152 OF.

M-700

M-700 combines the 31,444 Iblhr liquid distillate from the ace6c acid distillation

column at 275.2F and 95 psi with the recycle stream S-702 which is at 182.5F and 40

53

psi and has a flow rate of 5, 111 Ib/hr. It integrates the two streams to liquid stream S-703

and lowers the pressure to 39 psi, with the corresponding temperature of214.1 of.

M-740

M-740 combines the 43561blhr liquid stream S-706 (distillate from the near

azeotrope distillation column) at 181.9 of and 25 psi with liquid stream S-707 (the

bottoms from the acetone distillation column) which is at 184.3F and 24.9 psi and has a

flow rate of755 lb/hr. The combined stream is a liquid with a flow rate of 5,111 Ib/hr at

182.3 OF and 24.7 psi.

Pumps

P-I10(seespec. sheetp. 95)

Pump P-ll 0 is used to increase the pressure of the combined acetic acid feed

stream S-1 04 to 263 psi, so that it will be at the proper pressure for reaction. The

entering stream is at ambient pressure because it includes the pure acetic acid feedstock

S-101, assumed to be at 14.7 psi. The net required power is 9.7 kW, but since the

efficiency is only 0.52 the total power supplied is 18.6 kW. This result from the

simulation agrees well with the hand-calculation for required power included in the

appendix. This is a centrifugal pump made of stainless steel. The purchase cost,

determined from cost charts, is $11,570 and the bare module cost is $57,860.

54

P-540 (see spec. sheet p . 96)

Pump P-540 is used to increase the pressure of the absorber solvent stream S-507

from 31.8 psi to 235 psi so that it can be fed to the top stage of the absorber AB-320,

which is at high pressure to maximize acetaldehyde recovery. The volumetric flow rate

is 1149 ft3/hr. The power requirement is 9.7 kW, but the efficiency is 0.52, so the total

power is 18.6 kW. This is a stainless steel centrifugal pump, and the bare module cost is

$57,860.

P-600 (see spec. sheet p. 97)

Pump P-600 is used to increase the pressure of the acetic acid distillation column

feed from 37.8 psi to 100 psi. The power requirement is 1.86 kW, but because of the

efficiency of only 0.485, 3.83 kW of electricity is necessary. This is a stainless steel

centrifugal pump, and the bare module cost is $30,860.


P-620 pumps the S-901 recycle stream from the stream splitter to the second stage

of the acetic acid distillation column DC-61 O. It increases the pressure ofthe stream from

85 psi to 98 psi . The temperature of the stream is virtual1y unaffected by this change in

pressure. An efficiency of 0.44 is used for the centrifugal pump. The electricity

requirement is 3.1 kW. The volumetric flow rate is 400 ft31hr. The estimated bare

module cost is $28,930.

55


P-730 pumps the recycle stream S-705 from the mixer M-740 to the mixer M-700. It

increases the pressure of the stream from 24.7 psi to 40 psi. The temperature of the

stream is increased from 182.3F to 183.4F by this change in pressure. An efficiency of

0.27 is used for the centrifugal pump. The electricity requirement is 0.27 kW. This is a

stainless steel centrifugal pump, and the flow rate is 98 ft3l1rr. The estimated bare module

cost is $12,860.

PB-500 (see spec. sheet p. 100)

The reboiler pump for the acetaldehyde distillation column DC-500 is a

centrifugal pump made of stainless steel, and used to pump the bottoms of DC-500. The

capacity is 289 gpm, and its power requirement is 7.46 kW. The estimated bare module

cost of the pump is $18,000.

PB-610 (see spec. sheetp. 101)

The reboiler pump for the acetic acid distillation column is constructed from

SS304. It has a design pressure of 93.65 psig, a design temperature of 408.30F and a

driver power of 11.2 kW. It is a centrifugal pump. The bare module cost is estimated to

be $45,000.

56

PB-SI0 (see spec. sheet p. 102)

The reboiler pump for the acetone distillation column DC-Sl 0 is constructed from

SS304. It has a design pressure of 15.00 psig, a design temperature of215.97F and a

driver power of 0.1 kW. Based on cost charts for centrifugal pumps, the bare module

cost is $10,930.


The reboiler pump for the near azeotrope column DC-900 is constructed from SS304.

It has a design pressure of 15.00 psig, a design temperature of 231.94F and a driver

power of 0.37 kW. It is a centrifugal pump, and the bare module cost for this pump is

$14,410.


The reboiler pump for the ethyl acetate distillation column DC-91 0 is constructed

from SS304. It has a design pressure of 15.00 psig, a design temperature of 312.94F and

a driver power of 0.3 kW. The bare module cost for this centrifugal pump is $13,540.

PR-500 (see spec. sheet p. 105)

The reflux pump for the acetaldehyde distillation column DC-500 is a stainless

steel centrifugal pump with a capacity of67.6 gpm. The work output is 1.49 kW, and the

bare module cost is $21,210.

57


PR-510, the reflux pump for the refrigerated acetaldehyde recovery column, is a

centrifugal pump made of stainless steel. Its capacity is 3.5 gpm at a driver power of 0.25

kW. The estimated bare module cost is $12,860.

PR -610 (see spec. sheet p. 107)

PR-610 is the reflux pump for the acetic acid distillation column. The required

power is 5.6 kW. It is a stainless steel centrifugal pump. The bare module cost is

$35,360.


The reflux pump for the stripper S-800 is made of stainless steel. It is a

centrifugal pump with required power of 0.25 kW. The bare module cost is $12,860.


The reflux pump for the acetone distillation column is a centrifugal pump and

made of stainless steel. The power required is 0.1 kW, and the bare module cost is

$10,290.


The power for the reflux pump for DC-900 is 1.1 kW. It is a centrifugal pump,

and the material is stainless steel. The bare module cost is $19,930.

58


The reflux pump for the ethyl acetate distillation column is made of stainless

steel. It is a centrifugal pump with driver power of 0.1 kW. The bare module cost is

$10,290.

Since the purchase cost for each of the pumps listed above is relatively small compared to

the total capital investment of this process, a spare for each pump was purchased to avoid

long delays if a pump goes out of service.

Reactor

RX-240 (see spec. sheet p. 112)

The process unit RX-240 is a packed-bed reactor filled withPd-Fe203 catalyst

pellets. The design ofthe reactor was limited somewhat by the reaction information

available in the Eastman Chemicals patent regarding this process (Tustin et aL, U.S Pat.

No.6, 121,498). The desired reaction (1) is the reduction of acetic acid to form

acetaldehyde.

CH3COOH + H2 -) CH3CHO + H20 (1)

In the hydrogen-rich environment of the reactor, acetaldehyde can be further reduced to

ethanol (2).

CH3CHO + H2 -) CHJCH20H + H20 (2)

In process simulations, reaction (2) was not modeled as a sequential reaction, but instead

the following direct hydrogenation of acetic acid was used:

59

(2a)

A major factor in the design of the catalyst for this process was to promote (1), while

suppressing (2). This can also be accomplished by lowering the acetic acid conversion:

at higher space velocity (low conversion) the selectivity to acetaldehyde is enhanced.

Other side reactions that were considered in designing the reactor were the production of

acetone from acetic acid (3), as well as the formation of light hydrocarbons (4).

Reaction (4) is used merely as a material balance to account for side products listed in the

patent. Selectivity to methane and C2 hydrocarbons (ethylene and ethane) was presented

as 2% total, so we assumed 1 % to methane and 0.5% for each ethylene and ethane.

The reaction conditions allow 46% conversion of acetic with reaction selectivities listed

in Table 7.

Table 7: Reaction Selectivity Reaction Number

Reaction Selectivity (Acetic Acid Conversion)

1 CH3COOH + H2 ~ CH3CHO + H2O 89% 2a CH3COOH + 2 H2 ~ CH3CH20H + 2 H2O 5% 3 2 CH3COOH ~ CH3COCH3 + CO2+ H2O 4% 4 3 CH3COOH + 9 Hz ~ 2 CH4 + C2~ + C2H6 + 6 H2O 2%

The gas hourly space velocity (GHSV) for volumes of gas per volume of catalyst was

reported as 2600 hr-1 under the reaction conditions. The catalyst consists of 20%

palladium on iron oxide pellets. Based on a feed gas flow rate into the reactor of

224,307 ft3/hr, this implies a catalyst bed volume 0[86.3 ft 3 (see Appendix for

Calculations). The reactor diameter is 4 ft., and the height of the catalyst bed is 6.9 ft. In

60

addition, the reactor has a 6 ft. footer, O.S ft. distributor, O.S ft. catalyst support, and a 3 ft.

header. The total height of the reactor is 17 ft. The bare module cost ofthe reactor, a

vertical vessel, is $267,400. A catalyst density of 42 lb/ ft3 was assumed, producing 3623

lb of catalyst in the reactor. At a price of $1 ,6S0/lb (Dr. Rob Becker), the cost for a

charge of catalyst is $S,979,000.

The operating conditions for the reactor are based on the optimal conditions

described in the patent: the feed enters at S99 of and 2S2 psia, with a hydrogen to acetic

acid ratio of SI1 by mole. Under the conditions of the patent, the catalyst does not show

significant degradation in perfonnance with time on stream, and because coking is not

considered a risk, we assumed a catalyst replacement of20 % per year (Vrana), or

complete replacement once every five years. The material of construction for the reactor

is stainless steel. One benefit of the catalyst described in the patent for the process is the

ability to use relatively inexpensive materials of construction compared to more corrosive

methods employed earlier. It was suggested that carbon steel may be satisfactory for the

material, since gaseous acetic acid is not expected to be corrosive. However, on the

suggestion ofBruce Vrana to limit possible sources of corrosion throughout the process,

we decided to use stainless steel for this vessel. Stainless steel is also stable at the

reaction temperature; it is classified as suitable up to SOO C (932 OF) (Perry).

Reboilers

R-SOO (see spec. sheet p. 113)

The reboiler for the acetaldehyde distillation column DC-SOO is a kettle reboiler

constructed from stainless steel. The heat duty is 16,892,000 BtU/hr at a bottoms

61

temperature of262.7 of. Steam purchased at 75 psig is used to provide the energy. Two

reboilers in parallel are used, and the effective heat transfer area is 8,108 ft2. The

estimated heat transfer coefficient is 105 Btu/(hr-ft2_OF). The installed cost calculated by

B-JAC is $153,460.

R-610 (see spec. sheet p. 114)

The reboiler for the acetic acid distillation column, constructed of stainless s

Production of Acetaldehyde From Acetic Acid

Documents

proposed process

process flowsheet

process description

process design recovers

price of acetic acid

lblhr of acetaldehyde

acetic acid authors

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