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1/53

.

Report No. 53Al

Interim

SYNTHETIC ETHANOL

AND

ISOPROPANOL

by PARK L. MORSE

January 1974

A private report by the

PROCESS ECONOMICS PROGRAM

STANFORD RESEARCH INSTITUTE

I

MENLO PARK, CALIFORNIA

Ethanol and Isopropanol, Synthetic, Supp. A. Part 1, January 1974


2/53

1

2

3

4

5

6

CONTENTS

INTRODUCTION

........................

SUMMARY

...........................

ISOPROPANOL BY DIRECT HYDRATION OF PROPYLENE:

REVIEW'OF PROCESSES

.....................

Tokuyama Technology

.....................

Deutsche Texaco Technology .................

ISOPROPANOL BY TOKUYAMA TECHNOLOGY

.............

Process Description

.....................

Process Discussion

.....................

Cost Estimates

.......................

ISOPROPANOL BY A PROCESS USING CATION EXCHANGE

RESIN CATALYST /

......................

Process Description

.....................

Process Discussion .....................

Cost Estimates

.......................

UPDATED PROCESSES FOR VAPOR PHASE DIRECT HYDRATION

OFOLEFINS

.........................

APPENDIX A

DESIGN AND COST BASIS

...............

APPENDIX B

SUMMARY OF WASTE STREAMS

..............

CITEDREFERENCES

........................

PATENT REFERENCES BY COMPANY

..................

1

3

7

7

13

19

19

27

28

35

35

45

46

53

55

59

63

67

V



3/53

ILLUSTRATIONS

3.1

Equilibrium Conversion of Propylene to Isopropanol . . . .

17

3.2

Isopropanol Production Rate with

Deutsche Texaco Catalyst . . . . . . . . . . . . . . . . .

18

4.1

Isopropanol by Tokuyama Technology . . . . . . . . . . . .

23

4.2

Isopropanol by Tokuyama Technology

Effect of Capacity and Operating Level on

Production Cost . . . . . . . . . . . . . . . . . . . . . .

33

5.1 Isopropanol by a Process Using Cation Exchange

Resin Catalyst . . . . . . . . . . . . . . . . . . . . . .

41

5.2

Isopropanol by a Process Using Cation Exchange

Resin Catalyst

Effect of Capacity and Operating Level on

Production Cost . . . . . . . . . . . . . . . . . . . . . .

51

Vii



4/53

TABLES

2.1

2.2

3.1

3.2

4.1

4.2

4.3

4.4

4.5

5.1

5.2

5.3

5.4

5.5

6.1

B.l

Process for Producing Isopropanol by Liquid Aqueous

Phase Direct Hydration of Propylene

Comparison of Economics . . . . . . . . . . . . . . . . . .

4

Process for Producing Isopropanol by Liquid Aqueous

Phase Direct Hydration of Propylene

Comparison of Technologies . . . . . . . . . . . . . . . .

5

Isopropanol by Direct Hydration of Propylene

Patent Summary . . . . . . . . . . . . . . . . . . . . . .

Tokuyama Experimental Data . . . . . . . . . . . . . . . .

9

14


Bases for Reactor Design . . . . . . . . . . . . . . . . .

19


Major Process Equipment and Utilities Summary . . . . . . .


Stream Flows . . . . . . . . . . . . . . . . . . . . . .

22

25


Total Capital Investment . . . . . . . . . . . . . . . . .


Production Costs . . . . . . . . . . . . . . . . . . . . .

29

31


Resin Catalyst

Bases for Reactor Design . . . . . . . . . . . . . . . . .

35


Resin Catalyst

Major

Process Equipment and Utilities Summary . . . . . . .


Resin Catalyst

39

Stream Flows . . . . . . . . . . . . . . . . . . . . . . .

43


Resin Catalyst

Total Capital Investment . . . . . . . . . . . . . . . . .


Resin Catalyst

Production Costs . . . . . . . . . . . . . . . . . . . . .

Ethanol or Isopropanol by Vapor Phase Direct

Hydration of Olefins . . . . . . . . . . . . . . . . . . .

Summary of Waste Streams in SRI Design Cases . . . . . . .

47

49

54

61

ix



5/53

1 INTRODUCTION

This report emphasizes the economics and technology for isopropanol

produced by the direct hydration of propylene.

Design cases based in

part on Tokuyama Soda and Deutsche Texaco technologies are included.

In

addition, design cases for ethanol and isopropanol manufacture that were

presented in Process Economics Program Report No. 53 (issued in 1969) are

updated.

Technical information for the study was taken from nonconfidential

sources. Tokuyama Soda and Deutsche Texaco have commercial isopropanol

plants that presumably use technologies somewhat similar to those described

in this report.

Appendix A contains a definition of terms, design conditions, and

the cost basis used.

'Appendix B presents waste disposal data.

Special acknowledgment is given the Loprest Company (fabricators of

ion exchange equipment) of Rodeo, California,

and Dow Chemical Company

(manufacturers of ion exchange resins similar to those used in the report)

for their help in the design work.

1



6/53

2 SUMMARY

This report emphasizes the manufacture of isopropanol (IPA) by the

direct hydration of propylene.

SRI also updated designs for ethanol and

isopropanol manufacture (see Section 6) by direct hydration that were

included in PEP Report 53.

The earlier report emphasized processes wherein hydration is carried

out by water and olefinic reactants in the vapor phase. The current re-

port evaluates the more recently developed reaction systems that operate

with liquid water and a high density propylene phase.

Table 2.1 summarizes the results of SRI's economic evaluation of the

Tokuyama Soda process,

and a process using a cationic exchange resin cata-

lyst.

When feasible, SRI used information authored by Deutsche Texaco

personnel for the latter process.

Even though the two processes have

significantly different reaction systems, the overall economics appear

to be quite similar.

Both of these processes apparently offer superior

economics to the vapor phase direct hydration process (Hibernia-Scholven*

technology) that was updated in Section 6. However, the SRI design for

the updated process contains appreciable uncertainties that could well

exceed in significance those encountered by SRI in evaluating the pro-

cesses shown in Table 2.1.

The stoichiometric equation for the hydration of propylene to IPA is

C3% + %O -

(CH3)sCHOH

A technical comparison of the two SRI design cases for liquid aqueous

phase direct hydration is shown in Table 2.2.

Both processes have been

commercialized in the last several years and appear to offer advantages

*

Now Veba-Chemie.



7/53

Table 2.1

PROCESSES FOR PRODUCING ISOPROPANOL BY LIQUID AQUEOUS PHASE

DIRECT HYDRATION OF PROPYLENE

COMPARISON OF ECONOMICS

Capacity =

360 Million lb/yr (163,000 metric

tons/yr) at 0.9 Stream Factor

CE Cost Index = 142

Capital investment (million $1

Battery limits,

excluding catalyst

Utilities and tankage

Total fixed capital, not including

waste disposal

Production cost ($/lb)

Labor (operators, maintenance,

control laboratory)

Propylene at 2.8$/lb

Miscellaneous materials

Utilities

Overhead, taxes, and insurance

G&A, sales,

and research

Interest on working capital

Depreciation of fixed capital

Fuel credit

Net production cost, excluding

waste disposal and royalty

Confidence rating

Process Using

Process Using

Tokuyama Soda

Cation Exchange

Technology

Resin Catalyst

7.0

3.3

11.8

11.4

0.12

0.12

2.17

2.31

0.07

0.07

0.88

0.92

0.16

0.16

0.75

0.75

0.07 0.07

0.33

0.32

to.071

(0.12)

4.48

C

6.6

3.3

4.60

C



8/53

Table 2.2

PROCESSES FOR PRODUCING ISOPROPANOL BY LIQUID AQUEOUS PHASE

DIRECT HYDRATION OF PROPYLENE

COMPARISON OF TECHNOLOGIES

Capacity = 360 Million lb/yr (163,000 metric


Plant yield on propylene (%)

Conversion of propylene per

pass (%I

Selectivity on propylene (%)

Average reaction temperature (OF)

Reaction pressure (psia)

Molar ratio of water .to olefin

at reactor inlet

Number of reactors

Type of reaction system

Reactor space-time-yield

[g IPA/(hr x liter)]

Utilities,*

per lb of alcohol

Steam (lb)

Cooling water (gallons)

Power (kwh)

Natural gas (Btu)

Operators (men/shift)

Process Using

Tokuyama Soda

Technology

95.7

Process Using

Cation Exchange

Resin Catalyst

89.4

65.0

75.3

98.5

96.0

490

282

3000

1200

27.7

2

*

13.8

4

t

260

108

4.4

6.1

22.9 22.4

0.041

0.026

1,780 0

3

4

*

Packed columns,

countercurrent flow with catalyst soluble in aqueous

phase.

t

Trickle cocurrent flow over resin catalyst.

*

Battery limits.

5



9/53

over other recently patented direct hydration processes.

A brief descrip-

tion of the two processes follows:

IPA by Tokuyama Technology

The hydration of propylene to IPA in this technology is conducted

at about 490'F and 3000 psia in the presence of a soluble catalyst of

Na9H[SiW~,~~)41m

Liquid propylene is heated to 465'F before being intro-

duced in the bottom of a packed reactor. The olefin then moves counter-

current to liquid water containing the catalyst, a dilute crude alcohol

being discharged at the base of the column.

The unreacted olefin is sepa-

rated from the crude alcohol by lowering the pressure, the unreacted pro-

pylene being recycled to the reactor.

The resultant aqueous stream con-

tains most of the IPA.

The IPA is then concentrated by distillation,

the predominately aqueous fraction being returned to the reactor, the IPA

concentrate being sent to the purification section of the plant.

By means

of distillation a 91 ~01% IPA product and an anhydrous IPA product are

produced.

The 91 voi% IPA is used for acetone manufacture.

The anhydrous

IPA is passed over activated carbon and marketed as premium grade isopro-

panel.*

IPA by a Process Using Cation Exchange Resin Catalyst

The hydration of propylene to IPA in this technology is conducted

at about 280'F and 1200 psia by passing propylene and an aqueous phase

downward over an acidic cation exchange resin.

Liquid propylene is

heated to 265'F by admixing with hot water before entering the reactor.

A dilute aqueous alcohol solution and a gas phase are discharged at the

base of the reactor.

The unreacted gaseous olefin is recycled to the

reactor after being separated from a liquid phase.

The liquid phase, con-

taining dilute IPA, is then sent to the purification section for concen-

tration and purification.

The latter operations are similar to the pro-

cedures described under the Tokuyama Soda technology; however, a much

larger quantity of water is separated (as a liquid phase) from the

alcohol and returned to the reactor as recycle.

*

Tokuyama customers have not required this step.

6



10/53

3 ISOPROPANOL

BY DIRECT HYDRATION OF

PROPYLENE:

REVIEW OF PROCESSES

This section of the report

contains a patent review and a discussion

of the more pertinent technological advances since the 1969 PEP report on

ethanol and isopropanol.

Table 3.1 summarizes the patents issued since

the 1969 report,

Major technological advancements have been made by

Tokuyama Soda and Deutsche Texaco.

Their technologies are summarized

below.

Tokuyama Technology

Tokuyama Soda has developed and commercialized a high-pressure,

high-temperature process for converting propylene and water to IPA by

use of a liquid phase reaction.

A small concentration of the catalyst

in an aqueous phase selectively converts at a high rate liquid propylene

to IPA.

The catalyst formula can be written as XmHn[Si(Wa010)4] where X is

hydrogen,

ammonium, methyl ammonium, ethyl ammonium, ethanol ammonium,

propyl ammonium,

or a water-soluble, salt-forming metal such as sodium,

potassium, lithium, copper, beryllium, magnesium, calcium, barium, stron-

tium, zinc, cadmium, aluminum, chromium, manganese, iron, cobalt, or nickel.

If the valence of X is equal to a, the sum of am + n is equal to 4 (m and

n are both positive integers).

Tokuyama Soda has used the abbreviation SW for silicotungstate

l3~w3qo)41.

A few of the catalysts tested are H4SW, NaaHSW, LiaHSW,

and Cui., HSW.

The pH of the aqueous catalyst solution must be controlled

to around 2.0 to 4.5. At lower pH's polymerization of the olefin occurs.

Typically, the pH of the catalyst solution is adjusted to around 3 by

the addition of acids or bases, the choice depending on the degree of

acidity of the catalyst in question. During the conversion of the olefin

the pH rises and presumably must be adjusted before the catalyst solution

7



11/53

Table 3.1

ISOPROPANOL BY DIRECT HYDRATION

OF PROPYLENE

PATENT SUMMARY

Reference No./ Priority

Patent No. Application

Patent Examnle

67521 J 43-14621

J 45-29163 Mar. 8, 1968

67543

Brit 1,238,556

US 749,308

Aug. 1, 1968

OC Atm Catalyst

Raw Materials

Products

Notes

Gelatinized

silica sol

Propylene and water

IPA

Performance

Yield = 36% Molar ratio CeBs/RaO = 20.

Sulfonated co-

polymer of

styrene and di-

vinylbenzene

Propylene, water, IPA and diiso-

and l&PO4 propyl ether

Yield = 30.8% Deactivation of catalyst is combated by addition

of 0.1 wtX RaPGa. Formerly the use of stainless

steel would deactivate system.

IPA/DIPE = 8 wt

ratio.

Acetic acid and

molybdophosphoric

acid

Propylene and water IPA and isopropyl

acetate

Probably olefin first forms ester with the acid.

Then ester is hydrolyzed to alcohol.

Diatomaceous

earth + Rap04

Oxides of Ti and

Zn

Propylene and water IPA

Yield = 33.1%

Conv.

= 65%

Yield = 2.1%

Propylene and water IPA Yield = 2.2%

Electric dis-

charge

w-n-0

Propane and Co, IPA and n-propanol

100,000 vo lts/cm and 200 set used for reaction.


Dealkalinated

zeolite


Yield = 8.0%

sv

= 380

Yield = 5.1% on

water

Molar ratio CsRs/BaO = 1.

Molar ratio CsRs/RsO = 2.4.

H3p04 on boro-

silicate

Propylene and water IPA Yield = 5.5% The propylene vapor leaving the reactor is con-

densed and recycled to the reactor.

H3p04-MOO3

paste Propylene and water

Acetone, IPA, and

acids

Acetone yield = 29%

IPA yield = 8

Yield = 27%

Reaction time = 30 minutes in autoclave.

Silica-alumina

Propylene and water

IPA Reactor effluent is cooled at 458 atm and

organic ph ase is separated from an aqueous

liquid phase that is recycled to the reactor.

The organic phase is flashed, the olefin being

separated from IPA and recycled.

Assignee

Asahi Chem.

Ind.

Celanese

100 14

150 103

67549

US 772,817

US 3,644,497

Nov. 1 1968

Celanese

160

67524

J 47-23524

J 44-31449

Apr. 25, 1969

Mitsui Toatsu

67523

J 47-23523

J 47-23523

June 17, 1969

Mitsui Toatsu

46209

US 3,497,436

us 606,759

Jan. 3, 1967

Monsanto

67477

us 3,450,777

J 39-54792

Sept. 29, 1964

Tokuyama Soda

67525

J 47-45323

Toray

67532

Fr 2,098,801

J 43-77010

Oct. 24, 1968

USSR 3

USSR

100688

Ital 22,094

Edison

Brit 1,166,121

Oct. 4, 1965

International

67494

US 561,836 Standard Oil

US 3,548,013

June 30, 1966 of Ind.

300

1

180 4

1

200 10

200 10

180

16

250 45

257 458

240

N+HSi(W3C&O)q

Propylene and water IPA Yield/pass = 70%

Selectivity = 99%

STY = 182

Molar ratio water/olefin = 27.

Alcohol, water,

and a completely soluble catalyst le ave the

reactor. After separation, an aqueous phase con-

taining the catalyst is returned to the reactor.

Also see Tokuyama Soda patents identified by

reference numbers 67562, 67563, 67564, and 67565.

Propylene and water IPA Yield = 32% Molar ratio water/olefin = 2.4.

67499 J 34798 Tokuyama Soda 280

Brit 1,281,120 May 8, 1969

67529

US 853,489

UOP

Fr 2,059,246

Aug. 27, 1969

150

80

MeS&H and

dioxane

9



12/53

Table 3.1 (Concluded)

ISOPROPANOL BY DIRECT HYDRATION

OF PROPYLENE

PATENT SUMMARY

Reference No./ Priority

/

Patent Example

Patent No.

Application

Assignee

OC

Atm Catalyst

Raw Materials

Products Performance

Notes

323154 Ger P1,768,207 Veba-Chemie HsPGa on carrier Propylene and water IPA Low molecular weight polymers are normally en-

Brit 1,269,553

Apr. 13, 1968 trained with recycle olefin and carried back to

the reactor where they deactivate th e catalyst.

Polymer formation is greatly reduced by keeping

recycle o lefin concentration at 95% or higher.

Also see 67531.

67569

Ger 2,147,737

67570

Ger 2,147,739

67571

Ger 2,147,740

67572

Ger 2,147,738

Sept. 24, 1971 Deutsche

c

Texaco

135-155 100 Amberlite@ 252, Propylene and water IPA and diiso-

etc. propyl ether

67578

J 46-59942

Mitsui Toatsu

200

J 48-26711

Aug. 10, 1971

67579

J 46-64744

Mitsui Toatsu

250

J 48-32809 Aug. 26, 1971

67580 J 46-64745

J 48-32810

Aug. 26, 1971

67568

us 3,705,912

US 127,030

Mar.

22, 1971

Mitsui Toatsu 200

UOP

140

18

STY = 108-126

Re2

7

Propylene and water

IPA

Yield = 34.2%

Conv. = 37.5%

Pyromellitic

anhydride

Propylene and water IPA, Me,CO

Yield = 36%

(trace), and

Conv. = 37.5%

isopropyl ether

Sulfonic acids:

CFaSOaH

Propylene and water

IPA, acetone, and

Yield = 65.0%

isopropyl ether

Conv. = 66.5%

45-84

Mo(V1) oxalate Propylene and water IPA

r Trickle flow of aqueous phase over catalyst and

downflow of gas with molar rati o of water/olefin

2 13 give high selectivity and yield/pass.

Heated water enters top of reactor, cooler water

is added along reactor length to serve as a

quench. Catalyst has high activity for at least

8,000 hours.

L

45 minute batch reaction.

Molar ratio water/

olefin = 6.4.

1 hour batch reaction.

Molar ratio water/

olefin = 6.4.

90 minute batch reaction. Molar ratio water/

olefin = 15.

16 hour batch reaction.

Molar ratio water/

olefin = 2.

11



13/53

is used again.

Other than the drop in pH,

the catalyst is stable and

requires little replacement.

The catalyst concentration is about 0.001 mol per liter.

Lower con-

centrations cause a falloff in activity; however, higher concentrations

cause no substantial benefit.

Either agitated batch-type or continuous column-type reactors are

suitable for the olefin conversion.

It would appear that a packed column

feeding the aqueous phase at the top and olefin at the bottom would be

ideal.

The product, together with catalyst solution and unconverted ole-

fin would then be removed as bottoms.

Inert gases or liquids present would

be expected to accumulate at the top of the columns where they could be

easily removed.

It is believed that Tokuyama Soda used a similar type

reactor to obtain the single pass performance data shown in Table 3.2.

The material of construction used in commercial reactors has not

been disclosed.

The low pH and high temperature would indicate that

carbon steel would corrode quite rapidly.

In addition, an Fe concentra-

tion >lOppm lowers catalyst activity. Accordingly, the use of ion ex-

changers, nonferrous materials, or chelating agents has been discussed

in a Tokuyama patent (67562).

Apparently diisopropyl ether and possibly trace amounts of polymer

and acetone are the only significant by-products formed. Tokuyama Soda

has stated that acids and aldehydes are not by-products.

The Tokuyama Soda technology has been used by SRI for a design case

in Section 4. The reader is referred to the design case for Tokuyama

recovery and purification technology.

Deutsche Texaco Technology

Deutsche Texaco has developed and commercialized a process for con-

verting propylene and water to IPA under a moderate pressure and temper-

ature.

The conversion is conducted with mixed phases of liquid and gas

in contact with an acid-type ion exchange resin.

Selectivity is high,

only a small amount of polymer and diisopropyl ether being formed.

13



14/53

Olefin feed

Propylene

Run 2

Run 3

Propylene

Propylene

Ethylene

Butene*

Aqueous solution+

Catalyst component

Molar concentration

of SW (mol/liter)

PH

.Na,HSW

0.001

3.0

Cq,sHSW

AlHSW

0.001 0.001

3.1 3.0

Na,HSW

0.001

3.0

Na,HSW

0.001

3.0

Feeding rates [kg/

(liter of reactor

vol x hr)]

Olefin

Aqueous solution

0.26 0.26

0.26 0.24

0.56

3.0

3.0 3.0

3.0

3.0

Reaction conditions

Temperature ('C)

Pressure (kg/en?)

280

250

IPA,

71

280

250

IPA

69

99

68

220

200

Product

Conversion (96)

Selectivity (%)

Yield (%I

STY [g alc./(liter of

reactor vol x hr)]

280

250

IPA

73

99

72

300

300

Ethanol Butanols

44

61

99

70

95

95

42

58

260

267 252 166

429

*

The starting butene was a mixture of 40% isobutylene and 4m butene-1, the

balance being substantially butane,

and the main product was a mixture of

secondary and tertiary butanol at a mixing ratio of about 1:l.

'SW is abbreviation for [Si(W,CJo),].

Table 3.2

TOKUYAMA EXPERIMENTAL DATA

Run 1

Run 4

Run 5

14



15/53

Except as noted,

the technology in this subsection is based largely

on articles published by Deutsche Texaco personnel (67575, 67576).

Commercially,

the reaction is carried out at 130 to 150C and 60 to

100 atm.

The moderate temperatures and pressures are feasible because of

the favorable thermodynamic equilibrium that is established and because

of the high activity of the catalyst.

Figure 3.1 is a Deutsche Texaco

equilibrium diagram that clearly shows that at temperatures as low as 250

to 300F, and pressures around 71 atm,

90% of a pure propylene feed theo-

retically can be converted to IPA.

Even when 20% inerts are present it

is possible to obtain equilibrium conversions around 80%.

Figure 3.2

shows the reaction rate as a function of propylene concentration in the

organic feed.

A patent (67472) assigned to Rheinpreussen (integrated with Deutsche

Texaco) may be the basis for some of the technology practiced commercially.

The patent data indicate that the catalyst declines about 6% in activity

over the first 1,000 hours. It is conjectured that regeneration could be

conducted with sulfuric acid.

A more recent patent (67570) shows a 15%

falloff in activity for an Amberlite@ 252 catalyst over an 8,000 hour

period.

These excellent results were achieved by increasing the tempera-

ture from 125 to 155'C to compensate for the tendency of the catalyst to

lose activity.

Presumably the decrease in activity continues until re-

generation is an economic necessity.

The optimum water-to-propylene molar ratio in the reactor is 12.5 to

15:l.

If lower ratios are used,

or if water distribution in the reactor

is poor, polymer formation is favored. Good distribution is provided by

allowing the liquid phase to trickle downward over the catalyst, the gas

phase moving concurrently with the liquid.

Commercial propylene usually contains some propane, and hence simple

recycle of unreacted feed to extinction is not feasible.

However, two

options would appear open to the IPA manufacturer:

15



16/53

Route unconverted Cs's to propylene plant

Recycle part of unconverted Cs's to reactor, using the re-

mainder for propylene plant, or as fuel.

The Deutsche Texaco technology has been used in part by SRI for a

design case in Section 5.

The reader is referred to that section for

information on recovery and purification technology.

16



17/53

Figure 3.1

EQUILIBRIUM CONVERSION OF PROPYLENE TO ISOPROPANOL

125 I50 175

Temperature, OC

200 225 250

275 300 325

loo

90

80

70

60

50

40

30

20

IO

( \

212 atm

250 300 350 400 450 500

Temperature, OF

550 600 650

Source: 67576

17



18/53

Figure 3.2

ISOPROPANOL PRODUCTION RATE WITH DEUTSCHE TEXACO CATALYST

I

I

I

I

I

I

I

-

74 78 8 86 90 94

PROPYLENE IN FEED, mol%

Source:

67576.

18



19/53

4 ISOPROPANOL BY TOKUYAMA TECHNOLOGY

This section presents a design case based on Tokuyama technology.

Section 3 contains a review of patents and the technologies of Tokuyama

Soda and Deutsche Texaco.

Process Description

The flow diagram for SRI's design case is shown in Figure 4.1. The

plant is composed of the following two sections:

100 section:

Propylene and water are converted to

crude IPA

200 section:

Crude IPA is concentrated and refined to

form -91 vol%IPA and anhydrous IPA.

The design is based on technical material forwarded to SRI by

Tokuyama Soda (67573) and a Tokuyama Soda patent (67499).

Table 4.1

shows the bases used by SRI for design of the reactor.

The Tokuyama

material did not include distillation facilities for producing 91 ~01%

IPA or for treating anhydrous IPA with activated carbon followed by dis-

tillation.

SRI added each of these process steps.

Table 4.1


BASES FOR REACTOR DESIGN

Reaction temperature (OF)


Molar ratio of water/olefin in feed to reactor

Catalyst

Catalyst concentration in water (mol/liter)

pH of catalyst solution

Conversion of propylene per pass (%I

Selectivity of IPA on propylene (96)

Yield of IPA on propylene per pass (%I

STY [g IPA/(hr)(liters of reactor volume)]

19

465-518

3000

27.7

Na3HCSi(%qd41

0.001

3

65.0

98.5

64.0

260



20/53

a j or equipment and utility requirements are given in Table 4.2.

Stream rates for producing about 360 million lb/yr of 100% IPA are given

in Table 4.3.

The production is split between -91 ~01% IPA (30.4 million

gal/yr to feed an acetone plant)*

and anhydrous IPA (27.7 million gal/yr

for marketing).*

Fresh liquid propylene (95 mol%) and recycle propylene are pumped

through heater E-102 and into the base of reactor R-101.

Followingup-

ward movement through the packing,

contact is made with a downward-flowing

aqueous phase (stream 5) that contains the soluble silicotungstate cat-

alyst.

Inert gas (propane, etc.) and some propylene are discharged at the

top of the reactor. IPA and unconverted reactants are discharged at the

bottom at about 3000 psia,

and then flow to separator V-101 where the

pressure is let down to 65 psia.

The flashed vapor phase, together with

propylene from C-101 is recycled to the reactors.

The liquid phase from V-101, containing the catalyst as well as most

of the water and IPA discharged from the reactor,

is fed to the azeotropic

column C-101. The catalyst and a great share of the water is taken off

as bottoms.

A small stream (stream 50) is bled off to prevent the buildup

of high boiling polymers in the catalyst system,

t

The remainder of the

bottoms flow to one of three catalyst storage tanks (T-lOlA-0. Each of

the tanks operates automatically on the following three hour cycle:

b

Receiving, 1 hr

b

Discharging, 1 hr

0 Makeup of chemicals and mixing, 1 hr.

Catalyst (stream 20),

and an acid (stream 19) to adjust for pH's higher

than 3, are added to the makeup tank to allow for losses in the system.

The solution being discharged from catalyst storage is returned to the

reactor after first being mixed with fresh water.

*

About 180 million lb/yr of 100% alcohol is produced for each of the

two products.

t

Tokuyama Soda has informed SRI that high boiling oligomers do not.form.

Accordingly,

stream 50 is probably not required.

20



21/53

The overhead from C-101 is partially condensed.

The vapor phase

(stream 9), which is mostly propylene,

is recycled to the reactor after

a small stream (stream 10) is bled off to prevent a buildup of nonreactive

*

components. The liquid phase (stream 11) is crude IPA that contains water,

ether, and polymer as impurities.

The crude IPA is first mixed with a small amount of caustic solution

to neutralize any acidic components and then is fed to light ends column

c-201. Ether, water,

and a small amount of IPA and polymer are distilled

overhead and condensed to form two liquid layers. The bottom layer is

primarily water and is discarded or incinerated. Part of the top layer,

consisting of mostly diisopropyl ether, is drawn off and used as fuel,

while the remainder is refluxed to the column. The bottoms from C-201

are split,

one half being used for manufacture of 91 ~01% IPA and the

remainder being used for producing pure IPA.

The 91 ~01% IPA is produced as a distillate in C-202.

In the pro-

duction of pure IPA the bottoms from C-201 are first dehydrated.

Benzene

(stream 32).is used as reflux in C-204 and acts as an azeotroping agent

for water. The water distilled off is condensed, separated from a benzene

layer, and eventually is removed as bottoms from C-203.

In the design

the bottoms are recycled to C-101; however, SRI has learned that it is

acceptable to return them directly to the reactor.

Anhydrous IPA is taken

off as bottoms from C-204 and then removed as a distillate from C-205.

The anhydrous distillate is then passed over activated carbon and filtered.

The use of activated carbon is discussed further in the following sub-

section.

A summary of waste disposal streams is included in Appendix B.

*

SRI uses two bleed streams (6 and 10) to prevent a buildup of a range

of molecular species in the system. Tokuyama Soda has indicated that

its existing commercial design requires but one bleed stream.

21



22/53

Table 4.2


MAJOR PROCESS EQUIPMENT AND UTILITIES SUMMARY

Capacity =



Yajor Process Equipment

Equipment

Height

Number

N8lne (it)

Reactors

Diameter

(ft)

R-101

Reactors (2 units)

35

6.0

Coluallls

c-101

c-201

c-202

c-203

C-204

C-205

C-206

Aaeotropic column

Light

enda

Heavy ends column

Benzene recovery column

Drying column

Finishing column

Activated

carbon

treaters

(2 units)

30

92

42

62

75

32

15

E-101 Condenser

60 4.20

Carbon steel Carbon steel

E-102 Heater

400

22.60

Carbon steel

Carbon

steel

R-103 Exchanger

5,000

135.00 316 8s 316 as

E-104 Beater

630

30.00

Carbon

steel

316 8s

E-105

Cooler

110

1.00 Carbon steel Carbon steel

R-106

Condenser

620 16.50

Carbon

steel

Carbon steel

E-107 Reboiler

2,200 43.00 Carbon steel

316 ss

R-106

Condensers (2

units)

6,000

50.50 Carbon steel

Carbon steel

E-109

Exchangers (2 units)

4,OW 21.60

316 8s

316 ss

E-110 Cooler

660

11.60 Carbon steel

316 88

E-201 Reboiler


Carbon

ateel

g-202 Condenser

3,920 33.00

Carbon

steel Carbon steel

E-203

Reboiler

460

9.30

Carbon

steel Carbon steel

E-204

Condenser (air cooled)

1,400 10.90


E-205

Reboiler

600 12.00

Carbon steel

Carbon steel

E-206

Condenser 6,700 73.30 Carbon steel Carbon steel

E-207

Reboiler

5,100 61.30


E-206

Reboiler

1,090 13.10

Carbon steel

Carbon

steel

E-209

Condenser (air cooled)


Carbon steel

v-101

v-102

v-201

v-202

v-203 ) 4

v-205

T-101

T-161

T-201

Exchangers

Vessels and Tanks

Beparator

Reflux

drum

Reflux drum

Reflux drum

Reflux drums

Surge vessel

(3 units)

Propylene storage tank

Material of Construction

316

ss-clad shell;

porcelain packing

12.4 316 8s clad 316 88

6.3 Carbon steel

Carbon steel

4.0


4.2 Carbon steel

Carbon steel

11.0 Carbon steel

Carbon steel

4.9 Carbon steel

Carbon steel

3.0

Carbon

steel


Shell Trays

Heat Load, ea

Gize, ea

(million


(sq ft)

Btu/hr)

Shell

TUbeS

Vol, ea (gal)

Baterial of Construction

2,600

316 ss

10,000

Carbon steel

3,000

Carbon steel

600 Carbon steel

6,000

Carbon steel

3,000

Carbon steel

60,000 Fiberglass

65,000

Carbon

steel

200

Carbon steel

Remarks

30 ft of packing each

12 valve trays, 24 in. spacing






22



23/53






Equipment

Number

Name

Major Process Equipment

Vol, ea (gal)


Remarks

Vessels and Tanks

(Continued)

T-202

T-203

T-252

T-253

(2 units)

(2 units)

Alcohol storage tanks

(2 units)


(2 units)

33 ) 00 Carbon steel

37,000

Carbon steel

1,200,000

Carbon steel

1,400,000

Carbon steel

Size (bhpl

K-101

K-102

Compressors

Compressor

Compressor

Pumps

710

300

Carbon steel

Carbon steel

100 section:

14 operating, no spares; 688 operating bhp

200 section: 21 operating,

no

spare*;

56 operating bhp

Utilities Summary (Average Conaumptions)

Cooling water (gpm)

Process water (gpm)

Electricity (kw)

Steam

at 150 psig (lb/hr)

Natural gas (million Btu/hr)

Inert gas, low pressure (scfh)

Battery Limits

100

Total Section

17,500

3,500

56 56

1,969 1,767

202,000

50,000

62

82

20,000 17,000

Additional Requirements

Utility

To Operate Utilities

Electricity (kw) Makeup Water (gpml

Steam

226

20

Cooling

water 261

350

-

-

Total

487

370

200

Section

14,000

102

152,000

3,000



24/53

Figure 4.1

ISOPROPANOL BY

TOKUYAMA TECHNOLOGY

150Fb Cw

vE-IM utr,

t

.- .--.

Lqdd

Propylene

hilr

Cad.nsat.

R-101

v-101

c-101

RMCbl

s9$umt~

l-lOlA,,a6C

(2 hih)

Azwlroplc Column

cddyst 5torape

Crud* Ale&d

1 -

Acid

30% NoOH

V-205

Pnhydmm IPA

Fresh Actvated

T202A&B

-

Lne Flter

c-201

c-202

Lgh? End,

Hwvy Ends

Collmm

CdW

C-203

Benzene Recovery

CdWt

c-20(

DrybaRColumn

C-205

Fnishing Cdumn

i

C-206A&B

:

Actvated Cmko

Treater

23



25/53

Table 4.3


STREAM FLOWS

Capacity = 360 Million lb/yr

(163,000 metric tondyr)

at 0.9 Stream Factor

Stream Flows (lb-mol/hr)

(1)

(2)

(3)

(4)

(6) (7) (9) (11)

(12)

(13) (14) (15) (16)

(17) (18) (19)

---

(5) -

(8) - -

10)

- --

-

-

--

5.66

-

-

-

63.35

-

6.02

--

--

-

-- -

-

10.74

-

tr

--

927.59

-

-

-- --

4.72 --

tr

-

664.24

-

-

-

-- -

- --

4.72 -

- tr

777.61 66.42

- -

- --

-- -

-

--

- -

--

-

--

- -

-- -

tr -

- -

tr -

*

86.42 -

- -

86.42 -

-

--

--

-

-

5.65 lb

--

-

-

106.7

(20

-

-

4.72

-

777.61

--

1.22

--

-

1.24

765.42

(22) (23)

(24)

--

-- 11.45

1.22 30.00

-- -

-- -

-- --

776.96 30.970.91

11.45 -

-- -

11.45 -

--

--

- -

30.00 --

--

- -

-- --

--

-

- --

- --

- --

-

- -

-- I

--

-

30.970.91 --

- - 31,077.61 --

--

-

- --

-

-- --

-

1,005.36

1,000.64

-

-- --

-- 146.24

133.56

Activated carbon

Ben?Zne

Diisopropyl ether

Acid

Isopropanol

Na0H[Wh4c). 1

Polymert

Propane

Propylene

Sodium hydroxide

water

--

tr

66.39

3.35

11.45

-

11.45

30.00 -- 31.29

--

41.94

32.72

tr

6.00

417.39

--

--

-

32.639.06

-- 32.919.94

11.45

-

31.22

17.63

17.65

227.43

225.16

-

--

31,656.58

--

-

- -

-- 257.73

256.52

--

-- -

--

-- -

--

- -

191.70 469.86

67.84

- -

-- 0.08

-- 74.66

-- 1,215.40

- -

1,566.11 1.063.59

--

42.12

600.35

--

0.06

32.54

415.11

--

1.063.36

0.18

2.27

Stream Flows (lb-mol/hr)

(25) (2.6) (27) (28)

(29) -------

30) (31) (32) (33) (34)

(35)

(36) (37)

(36)

(39)------

40) (41)

(42)

(43)

(44) (45)

(46)

(47) (48) (49) (50)

----

-- -

--

--

-

12.05

--

--

-

-

765.14

-- -

--

- - -

380.57 --

362.67

-

-

--

364.57

--

-

2.106.70 2,108.70

-- -

-

-

805.14 805.14

--

50.49

-

--

205.99

50.49

-

-

205.99

-

2.159.19

-

-

1,011.13

--

--

-

--

-

- -

384.57

--

-

-

--

-- -

-

--

362.57 2.00 383.34

-

-

-

--

--

1.07

--

-

-

-

-

-

-

--

-

--

-

--

-

-

0.15

-

-

-

--

--

-

0.15

--

-

--

--

-

1.10

--

--

-

-

2.30

-

3.5 lb/hr

3.5 lb/hr --

-- -

- -

-- --

- 4.72

- -

-- --

382.27 -

- 0.63

--

--

-

-- -

-

tr

-- -

-

tr

--

-

1.24

363.40

-

-

--

-

-- --

--

--

--

-

920.16

728.46

--

-

-

-

-- --

-

-

-- -

634.08

442.36

-

-

-

tr

--

-

-

--

-

0.62

1,362.54

--

-

-

-

tr

-

--

- --

-

0.62

- -

-

tr

-

--

0.62

1.10

Activated carbon

Benzene

Diisopropyl ether

Acid

Isopropanol

Ne.aHbi(Ws4c)cl

Pkymert

Propane

Propylene

Sodium hydroxide

water

-- tr

- tr

-- --

- --

1.24 --

6.45 298.39

- -

-- --

- -

--

1.21

- -

-- -

- - -- -

-- --

0.62 --

- -- 10.61

0.32

- - --

-

--

-

-

401.70

tr

--

--

0.62

191.70

tr

tr

--

- --

-

--

- -

0.62

1.24 --

191.70

9.02 183.78

--

11.90

*

Add acid to achieve pH = 3.

t .

Mixture of high and low molecular wt polymers.

25



26/53

Process Discussion

It was necessary to make several assumptions in closing the material

balance.

It was assumed that polymers of varying chain length and di-

isopropyl ether are the only by-products formed in significant quantity.

Some of the polymers were assumed to have higher boiling points than

water and accordingly to circulate between the catalyst storage (T-101)

and the reactor.

A buildup of the polymer is prevented by use of bleed

stream 50.

Most of the low boiling polymers are assumed to distill over-

head in column C-201, the remaining polymers being removed as bottoms

from C-202.

The vapor-liquid equilibrium constants required for estimating

the phase split in V-101 are not known,

hence the material balance at

this point in the process is uncertain.

A small amount of acid probably must be added to maintain the catalyst

system at a pH of 3.

The preferred acid has not been revealed; however,

acetic acid would permit adequate catalyst activity and not pose the

corrosion problem that hydrochloric or sulfuric acid would.

The liquid from the reactor is considered to be corrosive because

of the presence of the acidic catalyst solution.

It is not known whether

the distillate from C-101 contains components corrosive to carbon steel.

Nonetheless, because of the low cost entailed,

SRI treated the distillate

with caustic before beginning purification of the IPA.

If corrosive com-

ponents are present in the distillate,

it will be necessary to change the

choice of materials for E-108 and V-102 from carbon steel to stainless.

SRI is not informed on the commercial technology used for treatment

of IPA with activated carbon.

In addition, the purpose and extent of this

practice is not known.

As a result, SRI's design for the treatment is a

speculative one. It would seem probable that the use of activated carbon

is justified commercially on the basis of improving odor.*

*

Tokuyama Soda has informed SRI that its IPA customers do not require

an activated carbon treatment of the product.

27



27/53

The following sequence is used in each of two columns:

.

Impurities adsorbed from anhydrous IPA by use of a flow

of 2,800,OOO gal of IPA per charge of carbon (1 lb

carbon/l,000 gal).

.

Na blow to remove IPA from bed, the IPA being returned

to c-208.

.

Sweetening off cycle. Process water is used to remove

residual IPA from bed, the effluent liquid being returned

to c-202.

.

Column recharged with fresh carbon.

Sweetening on cycle. Pure IPA is pumped briefly through

bed.

Adsorption (repeat of first item).

During adsorption there may be a tendency of carbon fines or im-

purities to plug the flow.

A once-a-day backwash with pure IPA should

alleviate the problem.

Cost Estimates

The battery limits and utilities investment, together with other

capital requirements, are given in Table 4.4. Production costs are given

in Table 4.5. Figure 4.2 shows production cost as a function of plant

capacity and operating level.

Tokuyama Soda has made public cost and performance data for their

process (67573).

This information is compared with SRI's as shown below:

Tokuyama

Soda

(67573)

SRI

Battery limits investment for 30,000 metric

tons/yr (million $1

$2.6* $2.3

Raw materials and utilities

Propylene (lb/lb) 0.72-t 0.73t

Steam (lb/lb)

3.5

4.4

Electricity (kwh/lb) 0.09 0.04

*

Based on 266 yen = US$l.

t

As pure propylene.

28



28/53

Table 4.4


Battery limits equipment,

f.o.b.

Reactors

Columns

Vessels & tanks

Exchangers

Compressors

Pumps

S 431,200

231,500

185,300

671,500

183 200

299,900

$ 431,200

69,900

93 900

732,100

163,200

276,700

Total

Battery limits investment

2,202,600 0.62 0.72 1,767,000

7,006,000

0.77

0.67 5,569,000

Utilities & tankage

Cooling water

Process water

steam

Inert gas

Tankage

Dowtherm@

566,600

117,700

4,000

4,000

1.053,200 260,700

72,300

61,500

904,600

133.100

244,200

244,200

Total

2,866,900

0.89

0.63

s 921.200

Utilities 0 tankage

investment 3,265,OOO 0.86

0.80

$ 935,000

0.75 0.66 2,330,000

BATTERY LIMITS &

Ul'ILITIES OST

10,273,000 0.80

0.71

6,504,000

b.80

0.72 3.769,000

General service facili-

ties et 15% of above

1.541.000

TOTAL FIXED CAPITAL

$11,814,000

Interest on construction

loan at 9.5%/yr*

start-up cost

Working capital

746,000

933 ) 000

2,609,OOO

TOTAL CAPITAL INVBS'I?dBNT,

not including land

16,104,000

TOTAL CAPITAL INVESTMENT



CE Cost Index = 142

Total

Reaction-Recovery

Section

cost

capacity Capacity

Exponent

Exponent

J L-

Down

cost

l L-

Down

0.95

n.95

0.95

0.62

0.60 0.59

0.95

0.94

0.41

0.40

0.H9

0.62

0.65

0.76

0.81

0.74

0.95 0.79

0.76 0.79

0.91 0 39

0.95

0.52

0.73 0.73

0.50 0.50

0.76 0.70

Purification Section

Capacity

Exponent

cost

ulr

DOWll

--

S

161,61Kl

91,400

139,400

23,200

415,600

sl.439.oor)

470,900

792,500

10,900

771,500

2,045,700

0.64

0.47

0.61

0.40

O.XG

0.82

0.26

0.22

0.66

0.54

Il.61

0.44

0.95

0.79

0.91

n.t49

0.95

0.52

0.95

0.95

0.93

0.89

0.91

n.xs

0.80

0.68

*

Interest calculsted over half of construction Period of 16 months,



29/53

Table 4.5


Labor

Operating 3 men/shift, $6.25/man-hr

Maintenance

3%/yr of battery limits cost

Control laboratory 20% of operating labor

Total labor

Materials

Propylene

2.8c/lb (6.17c/kg)

Catalyst

$l.OO/lb ($2.2O/kg)

Activated carbon 0.4$/lb (0.882$&g)

Caustic

4$/lb (8.82c/kg)

Maintenance


Operating

10% of operating labor

Basis or Unit Cost Units/lb

0.0001 man-hr

0.7741 lb 0.7741 tons

0.00012 lb 0.00012 tons

0.00759 lb

0.00759 tons

0.00108 lb

0.00108 tons

Total materials

2.24 4.93 8,100

Utilities

Cooling water

Zc/l,OOO gal (0.528c/cu m)

22.99 gal

191.9 cu m

Steam

$1.35/1,000 lb ($2.98/tori)) 4.424 lb

4.424 tons

Process water

35$/1,000 gal (9.25c/cu m)

0.0746 gal

0.6228 cu m

Electricity

1.35$/kwh (1.35'$/kwh) 0.0409 kwh

90.24 kwh

Natural gas

9Ochillion Btu (0.357c/ton cal) 0.0018 million Btu 997.7 ton cal

Inert gas (low pressure)

15$/1,000 scf (0.53c/cu m)

0.438 scf 27.34 cu m

Total utilities

TOTAL DIRECT OPERATING COST

Plant overhead

8oo/o f total labor

Taxes and insurance 2%/yr of fixed capital

Plant cost

G&A, sales, research

Cash expenditures

Depreciation

10o/o/yr f fixed capital

Interest on working capital 9.5%/yr

TOTAL PRODUCTION COST

By-product fuel credit,

streams 6 and 10 60c/million Btu

NET PRODUCTION COST

PRODUCTION COSTS

Capacity

= 360 Billion lb/yr

(163,000 metric tons/yr)


CE Cost Index = 142

Costs by Section (thousand $/yr)

Thousand

Reaction-Recovery Purification

Units/l,000 kg c/lb

0.1609 man-hr

0.05

0.06

0.01

0.12

c/kg

0.11

0.13

0.02

0.26

2.17 4.78

0.01

0.02

0.06

0.13

0.05

0.60

0.06

0.16

0.01

0.88

3.24

0.09

0.07

3.40

0.75

4.15

0.33

0.07

4.55

0.11

1.32

0.13

0.35

0.02

1.93

7.12

0.20

0.15

7.47

1.65

9.12

0.73

0.15

10.00

(0.07)

4.48

(0.15)

9.85

Wyr

164

210

33

407

7,803

43

11

16

210

17

165

'2,150

9

199

582

24

3,129

11,636

326

236

12,198

2,700

14,898

1,181

248

16,327

(240)

16,087

Section

Section

55

109

167 43

11

22

233 174

7,803

43

11

16

167 43

6

11

8,019

81

33 132

532

1,618

9

188

11

582

20

4

1,364 1,765

9,616

2,020

31



30/53

Figure 4.2


EFFECT OF CAPACITY AND OPERATING LEVEL ON PRODUCTION COST

6.5

4.0

3.5

3.0

l-

\

I -

\

\

\

I

I

I

I

I

I I

I

.5 .52 .54 .56 .5B .6

,fi .7

.75

.a .05 .9 .951.0

OPERATING LEVEL, fraction of design capacity



31/53

5 ISOPROPANOL BY A PROCESS USING CATION

EXCHANGE RESIN CATALYST

This section presents a design case based on a cation exchange resin

catalyst.

Although SRI drew heavily on literature published by Deutsche

Texaco, the design is not intended to represent the commercial practice

followed by Deutsche Texaco.

Section 3 contains a review of patents and

the technologies of Tokuyama Soda and Deutsche Texaco.

Process Description

The flow diagram for SRI's design case is shown in Figure 5.1.

The

plant is composed of the following two sections:

100 section:

Propylene and water are converted to crude IPA

200 section:

Crude IPA is concentrated and refined to form

~91~01% IPA and anhydrous IPA.

The design is based largely on Deutsche Texaco articles and patents

(67569, 67570, 67571, 67572, 67575).

A patent (67472) assigned to

Rheinpreussen,

which is integrated with Deutsche Texaco, is believed to

also have application.

Table 5.1 shows pertinent details of the reaction system.

Table 5.1

ISOPROPANOL BY A PROCESS USING CATION


BASES FOR REACTOR DESIGN

Reaction temperature (OF)

265-300


1200

Molar ratio of water/olefin in feed 13.75

Conversion of propylene per pass (%) 75

Selectivity of IPA on propylene (%I 96

Yield of IPA on propylene per pass (%I

72

STY [g IPA/(hr x liters of catalyst)] 108

35



32/53

The exact processing steps used by Deutsche Texaco for purifying

two grades of alcohol,

if indeed this was the intent, were not clear from

the literature.

As a result SRI made assumptions in the design of the

purification section.

It is understood that Deutsche Texaco uses a

significantly different purification train and only produces anhydrous

premium grade alcohol.

Major equipment and utility requirements are tabulated in Table 5.2.

Stream rates for producing

N360million lb/yr of 100% IPA are given in

Table 5.3.

The production is split between

-91~01% IPA (30.4 million

gal/yr* that is to feed an acetone plant) and pure anhydrous IPA (27.7 mil-

lion gal/yr* for marketing).

Recycled hot water (stream 2) and fresh liquid propylene (95 mol%)

are mixed to form a two-phase (vapor-liquid) system. The olefin-water

mixture is admixed with recycle stream 3 and the combination then enters

the top of reactor R-101 at 265'F and 1200 psia.?

The liquid phase

trickles downward through a series of four beds packed with a cation ex-

change resin (such as Rohm and Haas' Amberlite 8 252),the gas phase moving

concurrently with the liquid; as IPA is formed, the heat of reaction is

compensated for by the addition of quench water (stream 4) into each of

the packed beds.

The reaction mix leaves the reactor at the base and flows to high

pressure separator V-101 where gas and liquid phases are split. Pressure

is controlled in the reactor by regulating the flow of gas from the sep-

arator, the off-gas then being recycled to the reactor. The liquid phase

flows to low pressure separator V-102, where most of the remaining soluble

gas is flashed, and then compressed and recycled to the reactor. However,

part of the recycle gas is bled from the system (s.tream 10) and returned

to the propylene plant for removal of propane.

*

About 180 million lb/yr of 100% alcohol

is

produced for each of two

products.

t

Deutsche Texaco does not recycle this stream to the reactor, but returns

it, along with stream 10, to the propylene plant.

36



33/53

Crude alcohol flows from the low pressure separator to light ends

column C-201 where ether, water,

some IPA and a small amount of polymer

are distilled overhead, condensed into two liquid layers, and drawn off.

The upper, ether layer is used as fuel and also refluxed to the column

while the lower,

water layer is discarded or incinerated.

The bottoms from C-201 are fed to C-202, where 91 ~01% IPA is dis-

tilled overhead. Half of the product is sent to the acetone plant, the

remainder is pumped to C-204 for dehydration.

Benzene (stream 32) as used as reflux in C-204 and acts as an

azeotroping agent for removing water from IPA.

The water that is distilled

is condensed, separated from a benzene layer, and eventually removed as

bottoms from PAC-101.

The aqueous bottoms (stream 28) from C-202 contains sodium ions and

possibly some iron ions that must be removed before the water can be re-

used in R-101. This is accomplished by routing stream 28, as well as

stream 22, through ion exchange resins in PAC-101.

In SRI's design

PAC-101 operates with two columns (A&C) on stream while duplicate columns

undergo some stage of regeneration.*

Stream 28,

in combination with water

from various other sources in the process,

is fed (in stream 19) to

column c,

which contains a weak acid cation exchanger such as

Dowe

x@ CCR-2

to remove Na+ and possibly other cations.

The effluent is passed down-

ward through a second column containing a weak base anion exchanger, such

as Dowe

# WGR, to remove S04-- ions. Regeneration of the cation exchanger

is accomplished by washing with 0.075 wt% HaSO, (stream 16) followed by

a thorough water rinse (stream 15).

The cycle for each of two fully automated cation exchange columns

is as follows:

.

8 hours on stream

.

10 minutes backwash

*

It is understood that Deutsche Texaco uses a significantly different

design than SRI for PAC-101.

37



34/53

.

40 minutes acid wash

0 40 minutes rinse.

Because of the lack of information on SO," content, the size of

the anion exchange columns was arbitrarily made the same as that of the

cation exchange column. No effort was made to fix the sodium hydroxide

(stream 14) and rinse requirements (stream 13) for the regeneration of

column A.

Anhydrous IPA is taken off as bottoms from C-204 and then removed

as a distillate from C-205.

The anhydrous distillate is then passed over

activated carbon and filtered.

The use of activated carbon is discussed

further in the following subsection.

A summary of waste disposal streams is included in Appendix B.

38



35/53

Table 5.2




Capacity =,360 Million lb/yr (163,000 metric



Number

Name

Height Diameter

(ft) (ft)


Remarks

-101

Reactors (4 units)

Columns

Ether column

Aseotropic column

Benzene recovery column

Drying column

Finish column

Activated carbon treaters

(2 units)

-210

Exchangers

Heater 100

Exchanger

1,900

Cooler

40

Condenser

30

Condenser

10

Cooler '

2,200

Reboiler 4,100

Condenser

7,700

Reboiler

3,380

Condensers (2 units)

6,000

Reboiler

600

Condenser

6,700

Reboiler

5,100

Reboiler

1,090

Condenser 1,700

2xchanger

80

Vessels R Tanks

6eparators

Reflux drum

Reflux drum

Reflux drums

6urge vessels

Propylene storage tank

(2 units)

(2 units)


(2 units)


(2 units)

40

8.0

316 ss-clad shell

92 12.8 Carbon steel Carbon steel


62

4.2



32 4.9 Carbon steel

Carbon steel

15 3.0 Carbon steel

Size, ea

(sq ft)

Heat Load, ea

(million

Btu/hr)


Shell Tubes

3.40 Carbon steel 316 ss

36.40 316 6s 316 ss

0.60

Carbon steel

316 ss




80.60 Carbon steel

Carbon steel






61.30 Carbon steel

Carbon steel

13.10 Carbon steel

Carbon steel



~01, ea (gal)

1,200

316 ss clad

9,000

Carbon

steel

5,000

Carbon steel

6,000

Carbon steel

20,009 Carbon steel

13,000 Fiberglass

65,000 Carbon steel

200

Carbon steel

33,000 s Carbon steel

37,000

Carbon steel

2,000

Carbon steel

1,200,000

Carbon steel


Shell Trays


1.400,000 Carbon steel

44 valve trays,

24 in. spacing

12 valve trays,

24 in. spacing

40 valve trays,

24 in. spacing

36 valve trays,

24 in. spacing

15 valve trays,

24 in. spacing

39



36/53





Capacity =




Equipment

Number

Name

Size


Remarks

Compressors

K-101

Compressor

13 bhp

Carbon steel

K-102

Compressor

100 bhp

Carbon steel

Package Units

PA0101

Ion exchanger

500 gpm

Pumps

100

section:

8 operating,

no

spares;

645

operating

bhp

200 section:

18 operating,

no spares; 115 operating

bhp

Utilities Summary (Average Consumptions)

Battery limits

100 200

Total Section

Section

--

Cooling water (gpm)

17,200

2,700 14,500

Process water (gpm)

83

63

Electricity (kw)

1,lSS

794 394

Steam used at 150 psig

(lb/hr)

279,000

4,000 275,000

Inert gas, low pressure (scfh) 23,000 lS,OOO 5,000

Additional Requirements

To Operate Utilities

Utility Electricity (kw)

Makeup Water (gpm)

Steam

312

2s

Cooling water

257

-

- 44

Total

569

372



37/53

Figure 5.1

ISOPROPANOL BY A PROCESS USING

CATON EXCHANGE RESIN CATALYST

,........................................................................................................................RE~,o~EcoMRy SECTION

. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .

E-104 I _-. -

Rm0

Aqu-

limo

b-1

water Far PAC-101

Lquid Qs to

Pmpyheu Pkmt

I

quid

PwI*ll*

s

ooOF

1200 pi0

I -1 -I

Rockwash

I 1x1 I

1x1 I

0

R-101

RM.Zlor

(4 Units)

Reoston Quench

v-101

HI& Fmwm

Se$amtaI

Cd. IPA

I

v-102

Low Ptmwa

Separo~

wear tkPAC-101

rkgeonnotm

(Column A)

PAC-I01

la Exsho~~

Fmm C-2026262

Jwb NoOH

212F

15OF

15 pi0

I T 201

c-201

Ether Column

c-202

Azeatropic

Column

c-203

hnnrma Recovery

COIUM

C-2Q4

Drying Column

C-205

Fnishing Column

To C-202

C-206ARB

Ac&otad Cohn

TrtStCf



38/53

Table 5.3


CATION EXCHANGE RESIN CATALYST

STREAM FLOWS




*

stream Flows (lb-mols/hr)

15) f6l

-I

(71

-_

((11

(9) (10)

(11)

(12) (15)

(16)

(19) ( 20)

(21) ----

22) (23)

(24)

(25)

(26)

-- ---

2)

-

-

--

-

-

-

0.04

-

--

-

12.667.59

(3)

--

(4)

-

--

--

--

- -

-

--

15.07

- - -

-

-

780.84

--

0.01

--

--

-

-

2,165.Ol

--

270.00

1,080.OO

-

12.685.04

- -

1.18

1.18

270.00 21.26

266.76 21.00

- --

14.054.18 14,034.23

- - - -

- - --

-- 15.07 --

- - --

-- 780.84 -

-

--

--

--

-

--

-

--

-

340.00

-

-

--

-

-

0.82

-

--

-

--

592.00

--

-

-

--

-

--

-

--

-

15.01

-- -

-

780.84

--

-

0.04

-

-- --

-

-

1.48

-

14,832.60

2,310.36

-

1.18

-

-

1.57

14.041.59

--

--

-- -

-

3.209.91 3,194.84

-- -

440.20 426.12

-- --

251.34 250.16

-- -

- --

-- --

621.94 218.36

-

-

--

--

- -

-- -

12.07 766.76

--

--

--

0.05

- -

--

--

-

1.57

402.55

13.638.01

Activated carbon

Benzene

Diisopropyl ether

Acid

Isopropanol

Sulfuric acid

--

Polymer

--

Propane 45.15

Propylene 857.85

Sodium hydroxide

--

Water

-

-

-L

15.07

-

780.84

-

--

tr

248.75

245.76

-

19.94

- - -

- 1.18 --

45.15 -- -

44.61 -- --

- - 0.09

3.50 14,033.22 932.31

--

--

224.85

222.15

21.26

21.00

-

17.45 1.01

184.56

(27)

Stream Flows (lb-mols/hr)*

(28) ----------

29) (30) (31) (32) (33) (34) (35) (36) (37) (38) (39) (40)

o----

(42) (43) (44)

(45) (46)

(47) - -

48) (49)

- -- --

-- 2.108.70 2,108.70

- - --

- -- -

384.49 805.14 805.14

-

50.49

-

-

205.99

--

50.49

-

-

205.99

-- -

- -

-- -

-- -

-

--

634.08 449.52

-

--

2,159.19

--

- --

-- -

1.011.13 384.49

-- -

-

--

- -

-- --

--

--

1,362.54

--

-

-

-

-

382.57

--

--

--

--

-

-

--

--

-

-

1.92

-

--

-

-

--

--

--

-

-

-

383.34

--

-

-

--

--

--

-

-

--

-

1.07

--

-

--

-

-

-

- -

--

--

-

-

0.15

--

-

--

--

-

1.10

- --

3.5 lb/hr 3.5 lb/hr -

- - - - -

- -- - -- 15.07

- - -- -- -

2.22 382.27 -- -- 2.01

Activated carbon

Benzene

Diisopropyl ether

Acid

Isopropanol

Sulfuric acid

Polymer

Propane

Propylene

Sodium hydroxide

Water

-

-

--

--

--

--

-

-

1.57

8.00

-- - -

-- -- --

-- -- -

766.76 -- 384.49

-

--

- -

- --

--

0.15

-- -- -- -- -

-- - - --

1.18

- -- -- - -

- - - - --

- - -- -- --

1.10 - -- 10.75 1.03

-- - -

0.05 0.05 -

-- - -

-- - -

1.57 1.57 --

13,638.01 13.270.00 184.56

- - --

- -- -

-- -- --

- - -

-- -- --

184.56 913.02 728.46

- -

- --

_- --

- --

11.85 -

*

Streams 13, 14, 17, and 18 were not estimated.

43



39/53

Process Discussion

It was necessary to make several assumptions in closing the material

bal ante.

It was assumed that polymers of varying chain length and di-

isopropyl ether are the only by-products formed in significant quantity.

Some of the polymers were assumed to have higher boiling points than

water and accordingly to circulate between T-102 and the reactor. A

buildup of the polymer is prevented by use of stream 12, which is used

along with acid for regeneration of the cation exchanger.

Most of the

low boiling polymers are assumed to distill overhead in column C-201.

The vapor-liquid equilibrium constants required for estimating the phase

split in V-101 are not known, hence the material balance at this point

in the process is uncertain.

SRI designed for a slight rise in temperature through the reactor.

A recent Deutsche Texaco patent (67570) describes operation with a fairly

uniform temperature through the reactor. Over a period of time temperature

is allowed to increase to compensate for a decrease in catalyst activity.

The SRI design for the ion exchange treatment is highly speculative.

This was caused by lack of knowledge of the identity and quantity of all

of the ions present.

SRI is not informed on the commercial technology used for treatment

of IPA with activated carbon.

In addition, the purpose and extent of

this practice is not known. As a result, SRI's design for the treatment

is a speculative one. It would seem probable that the use of activated

carbon is justified commercially on the basis of improving odor in top

quality cosmetics.

The following sequence is used in each of two columns:

Impurities adsorbed from anhydrous IPA by use of a flow

of 2,800,OOO gal of IPA per charge of carbon (1 lb carbon/

1,000 gal).

Na blow to remove IPA from bed, the IPA being returned

to c-202.

.

Sweetening off cycle. Process water is used to remove

residual IPA from bed,

the effluent liquid being returned

to c-202.

45



40/53

.

Column recharged with fresh carbon.

.

.Sweetening on cycle.

Pure IPA is pumped briefly

through bed.

.

Adsorption (repeat of first item).

During adsorption there may be a tendency of carbon fines or impurities

to plug the flow.

A once-a-day backwash with pure IPA should alleviate

the problem.

Cost Estimates

The battery limits and utilities investments, together with other

capital requirements, are given in Table 5.4. Production costs are given

in Table 5.5. Figure 5.2 shows production cost as a function of plant

capacity and operating level.

Deutsche Texaco has supplied cost and performance data for its

process.

This information is compared with SRI's in the tabulation that

follows.

Cooling water is much higher for SRI's design; possibly, Deutsche

Texaco used a higher At for cooling water and made more extensive use of

air coolers.

Deutsche

Texaco

SRI

Battery limits investment for

100,000 metric tons/yr (million $1

Raw materials and utilities

Propylene (lb/lb)

Steam (lb/lb)

Cooling water (gal/lb)

Process water (gal/lb)

Electricity (kwh/lb)

$4.9*

$4.6

0.74-t 0.78-t

6.3

6.1

4.2

22.4

0.08

0.11

0.06

0.03

*

Based on 3.22 DM = US$l.

t

As pure propylene.

46



41/53

Table 5.4



Battery limits equipment,

f.o.b.

Reactors

columns

Vessels & tanks

Exchangers

Compressors

Pumps

$ 441,600

206,000

185,500

326,100

47,900

504,700

Total

Ion exchanger

Battery limits investment

$1,713,800

0.84 0.76

42,600

$ 6,558,OOO

0.81 0.74

Utilities and tankage

Cooling water

Process water

Steam

Inert gas

Tankage

579,800

5,300

1,363,600

77.300

904;600

Total

2,930,600

0.94 0.86

Utilities & tankage

investment $ 3,336,OOO

0.91 0.83

$ 352,000

0.83

BATTERY LIMITS IE

UTILITIES COST

9,894,OOO 0.64

0.77

4,799,000

0.85

General service facili-

ties at 15% of above

1,484.OOO

TOTAL FIXED CAPITAL $11,378,000

Interest on construction

loan at 9.5$/yr*

Start-up cost

Catalyst cost

Working capital

676,000

1,072,OOO

140,000

2,671,OOO

TOTALCAPITAL IRVRSTMRRT,

not including land

$15,937,000

TOTAL CAPITAL INVESTMENT



CE Cost Index = 142

Total

Capacity

Exponent

-L L

ost

E

Reaction-Recovery

section

capacity

Exponent

cost

A JL

$ 441,600

0.95

53,200

0.70

71,900

0.79

47,900

0.76

479,300

0.92

$1,093,900

0.90

42,600

0.40

$4,447,000

0.85

91,000

0.95

5,300

0.77

19,500

0.95

60,500 0.95

133,100

0.73

$ 309,400

0.66

DOWll

-.

0.95

0.68

0.72

0.76

0.90

0.88

0.40

0.84

0.79

0.77

0.89

0.51

0.73

0.71

0.68

0.83

purification Section

Capacity

Exponcnl

cost

s

208,000

0.80

132,300 0.58

254,200 0.62

25,400 0.36

5

619,900

0.72

$2,111,000 0.71

48X,600

0.95

1.344,lOO

0.95

16,600 0.95

771,500 0.95

S2,621,200

0.95

$2,984,000 0.92

s5,095,000

fl.S4

DlJWll

0 55

0.43

0.7H

0.29

0.59

0.54

0.79

o.n9

0.51

0.95

o.nx

o.ns

0.71

*

Interest calculated over half of construction period of 15 months.

47



42/53

Table 5.5


CATION EXCHANGE RESIN CATALYST

Labor

Operating

Maintenance

Control laboratory

Total labor

Materials

Propylene 2.8'$/lb (6.17c/kg)

Caustic 3.8C/lb (8.38$&J

Sulfuric acid

0.9$/lb (1.98$/kg)

Activated carbon

0.4$/lb (0.88X-&)

Maintenance 30/o/yr f battery limits cost

Operating


Total materials

Utilities

Cooling water

Steam

Process water

Electricity

Inert gas (low pressure)

Total utilities

TOTAL DIRECT OPERATING COST

3.42

Plant overhead

80% of total labor

Taxes and insurance

2%/yr of fixed capital

Plant cost

G&A, sales, research

Cash expenditures

Depreciation

Interest on working capital

TOTAL PRODUCTION COST

By-product credit

Fuel credit, Stream 10

Fuel credit, Stream 48

NET PRODUCTION COST

4.60

2c/l,OOO gal (0.528c/cu m)

22.6 gal

$1.35/1,000 lb ($2.98/tori))

6.11 lb

35$/1,000 gal (9.25c/cu m)

0.1096 gal

1.35c/kwh (1.35$/kwh)

0.026 kwh

15c/1,000 scf (0.53$/cu Ill)

0.5037 scf

lOX/yr of fixed capital

9.5Wyr

60c/million Btu (0.10)

6O$/million Btu

(0.02)

0.8267 lb

0.00133 lb

0.00175 lb

0.00759 lb

0.8267 tons 2.31

0.00133 tons 0.01

0.00175 tons

0.00759 tons

0.05

0.01

2.38

Basis or Unit Cost Units/lb

Units/l,000 kg c/lb

C/kg $/yr

Section Section

4 men/shift, $6.25/man-hr



0.0001 man-hr 0.2146 man-hr 0.06

0.05

0.01

0.12

0.13

219

99 120

0.11

196 133 63

0.02

44

20

24

0.26 459 252 207

5.09

0.02

0.11

0.02

5.24

8,333

18

6

11

196

22

8,333

6

18

133

10

11

63

12

8,586 8,482

104

188.6 cu m 0.05

6.11 tons 0.82

0.9144 cu m

57.33 kwh 0.04

31.44 cu m 0.01

0.92

0.11

1.81

0.09

0.02

2.03

7.53

163

26

2,970 43

14 14

126

84

27 21

137

2,927

42

6

188 3,112

8,922

3,423

0.10

0.06

3.58

0.75

4.33

0.32

0.07

4.72

0.22

0.13

7.88

1.65

9.53

0.71

0.15

10.39

3,300

12,345

367

228

12,940

2,700

15,640

1,138

254

17,032

(0.22)

(0.04)

10.13

(351)

(90)

16,591

PRODUCTION COSTS




CE Cost Index = 142

Total Costs

Costs by Section (thousand $/yr)

Thousand Reaction-Recovery

Purification

49



43/53

Figure 5.2

ISOPROPANOL BY A PROCESS USING CATION EXCHANGE RESIN CATALYST

EFFECT OF CAPACITY AND OPERATING LEVEL ON PRODUCTION COST

7 0

6 5

3 5

3 0

I

I

I

I

I

I

I I

I-

I

I I

I

I

I

I

I

5 S2 54 56 58 6

65 7

75 a a5 9 951 0

OPERATING LEVEL, fraction of design capacity

51



44/53

6 UPDATED PROCESSES FOR VAPOR PHASE

DIRECT HYDRATION OF OLEFINS

Processes evaluated earlier in the report use direct hydration re-

action systems that operate with liquid water and a high density propylene

phase, the product mainly being contained in a liquid effluent.

PEP

Report 53, issued in 1969, evaluated processes wherein reactants and prod-

ucts within the reactor are principally in the vapor phase.

This section

of this report updates the earlier so-called vapor phase direct hydration

processes.

Table 6.1 summarizes the results.

The current results for the ethanol process were obtained by using

updated costs, the plant capacity and design remaining essentially the

same as in the 1969 report.

For the process producing isopropanol, the

following adjustments were made to the 1969 case:

.

Plant capacity was increased to correspond to designs

in Sections 4 and 5 of this report.

.

Provision was made to produce 91 ~01% IPA (produced

in 1969 report) and anhydrous IPA. The revised design

produces products corresponding to those of the designs

in Sections 4 and 5 of this report.

Costs were updated.

In the earlier report,

SRI also evaluated a process based on Pullman

technology.

No updating of that process was attempted in the current

report, because the confidence rating, which was poor in 1969, would not

be improved by information acquired since then.

53

Ethanol and Isopropanol, Sy

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