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PROCESS ECONOMICS

PROGRAM SRI INTERNATIONAL

l

Introduction

The production of bisphenol-A from phenol and acetone was evalu-

ated in detail in PEP Report 81, November 1972. The process evaluated

used an anhydrous hydrogen chloride catalyst and required extensive

facilities for recovery of the catalyst and for purification of the

bisphenol-A product. The presence of both HCl and water necessitated

Menlo Park, California

94025

PEP Review No. 82-l-l

BISPHENOL-A FROM PHENOL AND ACETONE WITH AN ION EXCHANGE RESIN CATALYST--UNION CARBIDE TECHNOLOGY

Yoshio Kosaka and Kenneth B. Sinclair

(September 1982)

ABSTRACT

Bisphenol-A is produced commercially by the acid catalyzed conden-

sation of phenol and acetone under mild conditions of temperature and

pressure. This review evaluates the use of a cation exchange resin con-

densation catalyst according to Union Carbide technology. The process

avoids the handling of highly corrosive streams usual in traditional

HCl catalyeed processes and appears to be capable of producing very

high purity polycarbonate grade bisphenol-A simply.

Compared with the HCl catalyzed process, the resin catalyzed pro-

cess has a 2-3c/lb lower net production cost for polycarbonate grade

product. Production costs for epoxy gr-ade product are the same for

both processes. The resin catalyzed process shows a higher return on

investment for both polycarbonate and epoxy grade products.

extensive use of exotic materials in contact with process streams. The

bisphenol-A was purified by recrystallization from benzene. Product

purity as 99.5%.

An alternative catalyst system now widely employed is based on the

use of a cation exchange resin as the condensation catalyst. Such pro-

cesses are believed to have been developed by Bayer, Dow, Rhone-Progil,

Shell, and Union Carbide. These processes have the distinct advantage

that the catalyst is noncorrosive.

In this review, Union Carbide's resin catalyzed process technology

is evaluated both to elucidate the characteristics of this catalyst

system and to update bisphenol-A production economics on the basis of

modern technology. Some information on reaction chemistry is included

in this review; Report 81 contains a more detailed discussion.

Industry Status

As shown in Table 1.1, world production capacity for bisphenol-A

in 1981 was about 800,000 metric tons per year, with virtually all of

it being in North America, Western Europe, and Japan.

Major end-uses of bisphenol-A are epoxy resins and polycarbonate

resins which, in the United States, account for 94% of total demand.

Until 1976, epoxy resins were the largest consumer. From 1977 to 1980,

consumptions for epoxies and polycarbonates were about equal, but the

higher growth rate projected for polycarbonates is expected to make

this the major end-use in future. Other uses of bisphenol-A include

polyarylates and specialty polyester resins, polysulfone engineering

resins, and certain types of flame retardants.


Table 1.1

BISPHENOL-A PRODUCTION CAPACITY

United States Canada Mexico

North America

France Federal Republic of Germany The Netherlands United Kingdom

Western Europe

Japan*

Total

Number of Nameplate Capacity, l/1/81 Producers (thousand metric tons/yr)

5 1 1'

7

1 2 2 1

6

2 -

15

424 10 2

436

45 105 90 22

262

85

783

*Capacity will expand to 130,000 tons per year by 1983.

Source: Chemical Economics Handbook, SRI International.

Physical Properties

Bisphenol-A is a white solid in which the molecules consist of two

phenol groups joined through the center carbon atom of a propane mole-

cule. Its physical properties are listed in Table 1.2.

Because a variety of synonyms are used for bisphenol-A in the

chemical and patent literature, its identity is not always apparent to

the casual reader. Some of these synonyms are:

l p,p'-Isopropylidenediphenol

l 4,4'-Isopropylidenediphenol

l 2,2-(4,4'-Mhydroxydiphenyl)propane

3


Table 1.2

PHYSICAL PROPERTIES OF BISPHENOL-A

Appearance

Odor

Specific gravity at 25/25oC

Bulk density, flakes (lb/ft3)

Molecular weight

Freeting point (OC)

Boiling point at 4 mm Rg (OC)

Flash point, Cleveland open cup (OC)

Vapor pressure,(mn Hg)

179oc

193oc

21ooc

240.8OC

273OC

339oc

360.5OC

Heat of fusion (Btu/lb)

Solubility, approximate (gm/lOO gm solvent at 250C)

Acetone

Benzene

Carbon tetrachloride

Ethyl ether

Heptane

Methanol

Toluene

Water (250C; 83OC)

White crystals, flake or prills

Mildly phenolic

1.195

36-42

228.28

157.0

220

207

0.2

1.0

2.25

10.0

40.0

400.0

760.0

55.2

120

0.2

co.1

>llO

co.1

>120

0.2

<O.l; 0.34

Toxicity: Low in acute oral toxicity and only mildly irritating to the skin and eyes. Solutions greater than 1X, in some solvents, are capable of marked irritation and injury both to the skin and eyes. Dusts may produce irritation of the upper respira- tory passages, with sneezing and a burning sensation in the nose.

Sources: 354251 and trade literature. 4


l 2,2-Bis-(4-hydroxyphenyl)propane

l 2,2-Bis-(p-hydroxyphenyl)propane

l 2,2-Di-(4-hydroxyphenyl)propane

l g,g-Bis-(4-hydroxyphenyl)propane

l p,p'-Dihydroxydiphenyldimethyl methane

l Diphenylolpropane (common in Europe)

Bisphenol-A is usually sold under two general specifications: an

epoxy grade containing as much as 5% impurities but normally being

about 99% purity, and a polymer grade with an assay greater than 99.5X,

normally about 99.8% purity or above. Typical specifications are shown

in Table 1.3. The impurities present are phenol, the 2,4*-bisphenol

isomer, trisphenol, and chromans (Dianin's compound) which are de-

scribed below. The main reason for the very high purity requirement

for the polymer grade is the tendency for these impurites to form color

bodies under the alkaline reaction conditions used in polycarbonate

production. The APHA colors of the melt and of caustic solutions are

thus commonly used to specify polymer grade product. The quantities of

impurities are determined by gas chromatography (354047).

Table 1.3

TYPICAL BISPHENOL-A SPECIFICATIONS

Epoxy Grade Polymer Grade

Melting point (OC) 155.0 min 156.5 min

APHA color, 5Og/70 ml MeOH 100 max 25 max

Phenol content (wt%) 0.2 max 0.1 max

Moisture (wt% as shipped) 0.15 max 0.15 max

Iron content (ppm) 1.5 max 1.0 max

Ash (wt%) 0.02 max 0.02 max

Source: 354251.


Chemistry

The bisphenol-A production process evaluated in this review uses a

sulfonated styrene-divinylbenzene cation exchange resin catalyst (e.g.,

Dowex@ 50-X-4) for the liquid phase condensation of phenol with ace-

tone:

ii OH + CH, - C - CH,

2 x 94 phenol

58 acetone

228 18 Bisphenol-A water

The resin catalyst is active only in its anhydrous (phenol swol-

len) form and the degree of conversion to bisphenol-A (BPA) is thus

limited by the water by-product. A large excess of phenol is used to

increase BPA yield per pass.

The reaction product is an equilibrium mixture of the 4,4'- and

2,4'-bisphenol isomers, and small quantities of impurities such as

trisphenols and isomers of Manin's compound as follows (354037).

(11 2-(2-Hydroxyphenyl)-2-(Chydroxy~enyl)propa~ or 2,4’-bisphenol-k

OH

Freazing point = 11 l°C


(2) 2,4-Bir(qcrdimethyll-hydroxybenzyl)phanol or bisphenol-X:

C"3 - 7 - CH,

6H

Fraeting point = 181OC

(3) Co’-klydroxyphenyl)-2,2,4-trimrthylchroman or codimar or Dianin’s compound:

Freezing point = 158OC

OH

(4) 2-(4’.HydroxyphanylI-2,4,4,trimathylchroman or isomeric codimar, an isomar of Dianin’s compound:

OH 0 \ I 2

0 0

\ I

CH3

C"3 CH3

Fraazing point - 133OC

The first two compounds result from a substitution reaction taking

place with hydrogen in the ortho position rather than in the para

position on the phenol molecule. The Dianin's compound isomers are

formed by reaction of phenol with trace amounts of mesityl oxide,

(CH&C=CH- (C=O)-CH3, present in the reaction mixture.

7


In addition to these by-products, higher-condensation products,

tarry resins, and very small proportions of highly colored compounds

characterized by intense absorption of ultraviolet light, are also

present in the reactor product (354134, 354218).

Since the 4,4'- and 2,4'-bisphenol-A isomers are in equilibrium in

&he reaction product, the 2,4'-isomer can be recycled continuously to

the reactor feed to ensure 100% yield of the 4,4'-isomer. The 2,4'-

isomer is thus not in itself an objectionable component in process

streams. Similarly, some of the polyphenols which are not high colored

can be tolerated in process streams and in epoxy grade BPA product.

The highly colored chromophoric impurities, however, must be removed

from the process irrespective of whether epoxy grade or polymer grade

BPA is being produced. This is achieved by absorbing these impurities

in a bed of sulfonated styrene-divinylbenzene cation exchange resin.

Once loaded, the absorption bed is regenerated by washing with wet

phenol (20% water) (354218, 354221).

In the production of high purity polymer grade BPA, higher poly-

phenols and tars must be removed to prevent buildup in recycle streams.

This is achieved by cleavage of the polyphenols at elevated temperature

and in the presence of an alkali catalyst to form phenol and para-

isopropenyl phenol (PIPH). The PIPH readily dimerize8 at room temper-

ature so that the cleavage product contains both PIPH monomer and dimer

(354251). The phenol, PIPH, and dimer are separated from heavy non-

cleavable tars and the alkali catalyst by distilling these products

overhead simultaneously with the cleavage reaction:

c, - CH, - k - CH,

EH 2 CH3

PIPH PIPH dimer


-

-

a -

0

Under acidic conditions the dimer readily reverts to the monomeric

form and, in addition, the monomer reacts with phenol to give nearly

quantitative yields of BPA:

6 + 6 - o”~rQo” 3

C - CH,

!H 2

phenol PIPH BPA

In the process evaluated here, this rearrangement occurs simul-

taneously with the removal of color bodies, the cation exchange resin

in the color absorption bed acting as the acid catalyst. This

rearrangement is an essential step after cleavage since PIPH Is itself

a major source of color bodies in BPA products. In the presence of

air, PIPH is readily oxidized to a peroxide which in turn further

reacts to form colored polymeric compounds (354222).

Another advantage of this rearrangement step is that it reduces

the equilibrium concentrations of 2,4’ -isomer and polyphenols in the

main condensation rea’ctor product (354041). Under anhydrous condf-

tions, the cation exchange resin acts to rearrange the 2,4’-bisphenol-.4

to the desired 4,4’-isomer. Since both water and acetone are essen-

tially absent in recycle streams, conditions favor this isomerization

reaction and the concentration of the 2,4’-isomer in the condensation

reactor feed is substantially lower. Surprisingly, this also reduces

the equilibrium concentration of 2,4 ‘-isomer in the condensation reac-

tion product stream by about half. Normally the steady state concen-

tration of by-products obtained in the reactor product stream when

recycling all by-products without rearrangement is about 40 wt% on a

9


phenol-free basis. When a rearrangement step is included in the recy-

cle circuit, this equilibrium by-product concentration falls to about

20 wt%. The net result is a marked reduction in by-product concentra-

tion throughout the system, improving the efficiency of bisphenol-A

separation and thus yielding a higher purity product.

The recovery of high purity BPA from the reaction mixture is

achieved by crystallization. BPA is relatively unstable at elevated

temperatures and should preferably be processed at less than 15OoC.

Low temperature crystallixation can be achieved by recovering a high

purity 1:l molar phenol/BPA crystal adduct which can subsequently be

split into its two components. Bisphenol-A and phenol form a peri-

tectic and eutectic system, as shown in Figure 1.1. The equimolar

adduct contains about 30 wt% phenol and has an incongruent freezing

point (354251). Crystallieation from mixtures containing S-58% BPA

yields the equimolar adduct, while crystallization from more concen-

trated mixtures yields BPA crystals. Because the reaction impurities

are all soluble in phenol at temperatures between 37 and 9S°C, high

purity adduct crystals can be obtained, given adequate crystal washing.

10


160

150

140

130

120

ou 110

Figure 1.1

MELTING POINT DIAGRAM FOR THE

PHENOL-BISPHENOL-A SYSTEM

Adduct +

bisphenol -A

cfystals

I I 0 10 20 30 40 50 60 70 80 90 loo

WEIGHT PERCENT BISPHENOL-A

Source: 354251.

11


Process Description

This evaluation is based on a unit designed to produce 120 million

lb/yr of bisphenol-A at an on-stream factor of 0.9 (7,884 hr/yr). The

product is a high purity BPA grade suitable for polymer production and

containing 99.7 wt% BPA. The unit would also be suitable for producing

a lower purity epoxy grade BPA with some savings in operating costs.

Key process conditions and design assumptions used as a basis for

design are summarized in Table 1.4. The process flow scheme is shown

in Figure 1.2 (foldout at end of this paper). Material flows of the

numbered streams In this flow diagram are given in Table 1.5 and major

process equipment is listed in Table 1.6.

Referring to Figure 1.2, phenol and acetone in a molar ratio of

1O:l are heated to the reaction temperature in preheater E-101 and sent

continuously to BPA condensation reactors R-lOlA,B, which are jacketed

vessels packed with ion exchange resin and which operate in parallel.

Phenol and acetone are reacted at 167Op (7SOc) and marginally super-

atmospheric pressure to produce BPA. The residence time is 1 hour and

the conversion to BPA, based on acetone, is about 50%. The reaction

temperature is maintained by circulating tempered cooling water through

the jacket.

The effluent stream from the reactor is pumped to concentrator

E-103, in which unreacted acetone, water, and some phenol are removed

at 284O~ (14OOC) and 200 mm Hg. The distillate from the concentrator

is sent to dehydration column C-101.

The concentrator bottoms, now consisting of BPA, phenol and reac-

tion by-products, are pumped to crystallizer V-102, where they are

cooled to llS°F (46OC) to produce a slurry of 1:l molar phenol/BPA

adduct crystals in mother liquor.

The slurry is sent to the first batch centrifuges, M-lOlA,B, and

the mother liquor is separated from the crystals. The separated mother

liquor is collected in buffer tank T-104. The crystals are washed with

12


Table 1.4

BISPHENOL-A FROM PHENOL AND ACETONE WITH AN ION EXCHANGE RESIN CATALYST

DESIGN BASES AND ASSUMPTIONS

BPA synthesis reaction

Reactor type

Reaction temperature

Reaction pressure

Residence time

Phenol/acetone feed mol ratio

Conversions per pass Phenol Ace tone

Selectivity to 4,4*-bisphenol-A

Cleavage reaction*

Reactor type


Reaction pressure

Residence timet

NaOH/BPAs feed weight ratio

Tar/product weight ratio

Tar/cleaved BPA's weight ratio

Weight ratio of BPA cleaved/isomer cleaved

Conversion per pass on total BPA's

Selectivity on total BPA's To phenol To PIPH To tar

Jacketed column packed with ion exchange resin

75oc (1670F)

Atmospheric

1 hr

lO.O:l

10.1% 50.5%

80.5%

Distillation column with 5 valve trays

160°c (320OF)

200-250 mm Hg

100 minutes

0.0038:1

0.065:1

0.15:1

1:l

80%

100% 74.4% 25.6%

13


Table 1.4 (Concluded)

BISPHENOL-A FROM PHENOL AND ACETONE WITH AN ION EXCHANGE RESIN CATALYST

DESIGN BASES AND ASSUMPTIONS

Rearrangement and decolorization

Reactor type


Reaction pressure

Residence time

Phenol/PIPH feed mol ratio

Conversion per pass Phenol PIPH

Selectivity to 4,4'-BPA

Regeneration of rearrangement reactor

Temperature

Pressure

Solvent for desorption

Elution rate

Operating rates Rearrangement and decolorization Desorption

*Including BPA and BPA isomer.

*Bottoms flow basis.

Jacketed column packed with ion exchange resin

70°C (158OF)

Atmospheric

20 minutes

56:l

1.8% 100%

100%

70°C (158OF)

Atmospheric

Phenol (80 wt%)/water (20 wt%) mixture

1 reactor volume/hr

24 hr/day 12 hr/day

14


phenol filtrate from the second centrifuge, M-102. The wash liquor is

combined with the separated mother liquor.

The crystals are discharged to V-103, where they are reslurried in

recycle phenol from phenol stripper E-106 and fresh phenol feed from

storage. The slurry of crystals in phenol is pumped to the second

centrifuge, M-102. The crystals are centrifuged, washed with pure

phenol from the phenol stripper and sent to melter V-104. The sepa-

rated phenol is collected in buffer tank T-103 to be used as wash

liquor for the first centrifuge.

The phenol/BPA adduce crystals are melted in V-104 at 2660~

(13O'C) to give a solution of BPA in pure phenol. This is pumped to

phenol stripper E-106, in which phenol is evaporated from the product

at 392'F (2OOOC) and 5 mm Hg. The evaporated phenol is recycled to

adduct reslurry vessel V-103 and the second centrifuge, M-102. The

molten product from the phenol stripper is cooled and flaked for pack-

ing.

The distillate separated from the condensation reactor effluent in

concentrator E-103 is dehydrated in dehydration column C-101 by intro-

ducing a dry acetone vapor stream from acetone column C-102. An essen-

tially anhydrous mixture of phenol and acetone is obtained at the

bottom of column C-101. This mixture is cooled and returned to the

reactor feed tank.

The distillate from the dehydration column, consisting of acetone

and water with only traces of phenol, enters the acetone column, in

which acetone and water are separated-water leaving the column as bot-

toms and dry acetone leaving the column as distillate. The water is

sent to waste treatment for removal of trace quantities of phenol.

A portion of the dry acetone is recycled to the reactor feed tank.

The remainder is vaporieed and returned to the bottom of the dehydra-

tion column as a drying agent.

A portion of the recycle mixture of mother liquor and wash liquor

from the crystal separation step is sent to cleavage column C-103, in

15


which both BPA and its 2,4'-isomer contained in the mixture are cleaved

to p-isopropenyl phenol (PIPH) and phenol at 320O~ (16OOC) and 200-250

mm Hg, in the presence of an alkali catalyst. Most of the phenol and

PIPH formed are simultaneously distilled from the reaction mixture.

The remaining phenol and PIPH, and unreacted BPA and its isomer are

separated from the residual tars in BPA evaporator E-118, at 4500F

(2320C) and 15 mm Hg. The distillates from the cleavage column and the

evaporator are combined with the remainder of the recycle mixture and

sent to rearrangement reactors R-102A,B,C. These are jacketed vessels

packed with cation exchange resin and operating slightly above atmo-

spheric pressure. TWO reactors are on-line while the third is being

regenerated.

The rearrangement between the cleavage products to form the

desired BPA takes place at 158OF (7OOC). The effluent stream from the

reactor is recycled directly to condensation reactor feed tank T-101.

A small amount of highly colored compounds, which are by-products

of bisphenol-A reaction and are present in the mother liquor recycle

stream, are adsorbed by the ion exchange resin in the rearrangement

reactors. These color bodies are desorbed from ion exchange resin by a

periodic wash with phenol/water mixture containing 20 wt% water. The

phenol/water mixture is then separated from the color bodies by evapora-

tion, and is reused.

16


Table 1.5

BISPHENOL-A PROM PHENOL AND ACETONE WITH AN ION EXCHANGE RESIN CATALYST

Plant Capacity: 120 Million lb/yr (7.884 hr/yr)

at 0.9 Stream Factor

Mol Stream Plowe (lb/hr) weoxLooooooo

Phenol Acetone Water BPA BPA isomer BPA/phenol adduct Tar pIsopropenylpheno1 Sodium hydroxide

Total lb/hr lb mol/he

94.1 13,368 - 139,633 38.1 - 4,381 8,383 18.0 -- 198 228.3 - - 10,046 228.3 -- -- 9,041 322.4 - I

134.2 40.0

- I

- -

---

13,368 4,381 167.303 142 73 1,726

123,388 4,247 1,341 23,738 12.370

97,814

Trace 23,738 12,370

27,773 4,202 1,341

--

4,710,483 -

Trace --

--

91,496 --

Trace 8,431 12,370 21,643

--

2,848 -- -- -- 30

21.643 -- --

-- - -- -

167,304 133,942 33,316 4,710,483 133,942 24,323 1,632 1,198 433 30,038 1,131 98

Mol Stream Plowa (lb/hr) Wt (10) (11) (12) (13) (14) (13) (16) (17) (18) (19)

Phenol Acetone Water BPA BPA Isomer BPAiphenol adduct Tar p-I~opropenylphenol Sodim hydroxide

Total lb/hr lb mol/hr

94.1 38.1 18.0 228.3 228.3 322.4

134.2 40.0

110,720 12,921 20.013 2,848 22,072 9,166 13 9,131 15 4.903 -- -- -- -- -- -- -- --

Trace -- -- -- -- -- -- -- -- -- 8,431 - -- - -- 13,327 13,327 - 13,173 - 12,339 -- 30 30 -- 30 30 -- 30 --

-- 21,643 21,643 - - -- -- -- -- -- -- -- -- -- -- -- --

-- -- - -- -- -_ - -

---------- 131,490 12,921 42,690 24,323 22,072 24,323 13,372 9,131 13,220 4,903 1,268 137 280 98 233 163 67 97 67 52

%l Stream Flows (lb/hr) Wt (20) (21) (22) (23) (24) (23) (26) (27) (28) (29)

Phenol 94.1 Acetone 38.1 Water 18.0 BPA 228.3 BPA isomer 228.3 BPA/phenol adduct 322.4 Tar -- p-18opropenylphenol 134.2 Sodium hydroxide 40.0

Total lb/hr lb mol/hr

4,246 - 27,773 - - - -- - 42,967 2,099 42,967 - 2,104 40,863 - 1.390 137 216 1,373 11 203

-- -- -- we - -- -- --

43,939 --

Trace 3,346 4,900

--

- -- 31 -- --

--

66.781 --

Trace 5,083 7,440

-- --

- --

-- -- 31

41.068 32,185 79.306 713 303 765

-- - -- -- -- -- - -- -- - --

- ---- 4,246 44,337 30,029 43.183

43 828 340 732

-- - - - --

1,373 2,113 76 37

-- 31

62 2

Stream Flowa (lb/hr) (30) (31) (32) (33) (34) (33) (36)*

Phenol 94.1 Acetone 38.1 Water 18.0 SPA 228.3 BPA isomer 228.3 BPAjpheool adduct 322.4 Tar p-Isopropenylphenol 134.2 Sodiom hydroxide 40.0

Total lb/hr lb wl/hr

*Operation 12 hr/day.

44.792

31

1,866 -

47 1,601

-- - 269 -

- 113.438 -- 31

- 5,132 - 9,041 - -- 989 -

111,414 -- 31

10.046 9.041

-- -- --

16.800

4.200

--

-- me

1.866 -- - 47

1.6OL -- 989 269 31

- - --

2,618 -- 2,887 31 - --

3,783 1,020 130.329 130,532 21,000 29 - 1,291 1,269 412

17


TABLE 1.6

BISPHENOL-A FROM PHENOL AND ACETONE WITH ION EXCHANGE RESIN CATALYST

MAJOR EQUIPMENT

CAPACITY: 120 MILLION LB (54,000 METRIC TONS)/YR POLYMER GRADE BISPHENOL-A AT 0.90 STREAM FACTOR

E;;;;::“’ NAME

--------_ --__--_-_____-_-_--_--

REACTORS

R-18lA.B EPA REACTOR

I-l#ZA-C REARRAWCPWEWT REACTOR

NEAT EXCHANCERS

E-181

E-I#2

E-113

E-I#4

E- 106

E-116

E-187

E-106

E-109

E-Ill

E-Ill

E-112

E-113

E-114

E-115

E-116

E-117

E-115

E-119

E-129

E-121

E-122

E-123

E-I24A.I

E-125

E-126

E-127

E-120

E-129

E-138

E-131

E-132

E-133

lURltACES

N-ill

REACTOR PREHEATER 95 1

CONCENTRATOR PREHEATER 761 18.9

CONCENTRATOR 439 6.4

E-103 CONDENSER 1.26# 11.3

CRYSTALLIZER COOLER 9.1## 13.6

PHENOL STRIPPER 43# 3.6

PHENOL CONDENSER 1.99n 3.3

C-ill PREHEATER 241 1.6

C-181 CONDENSER 5.771 20.9

C-ill REBOILER 2.7611 13.2

ACETONE VAPORIZER 63# 9.4

C-Ill SOTTON COOLER 26P 3

C-112 CONDENSER 9.661 36.5

C-II2 RESOILER 3.661 36.5

C-l#S PREHEATER 361 5.1

t-1#3 CONDENSER 771 12.6

C-II3 REBOILER 2.611 12.9

SPA EVAPORATOR 261 1.1

E-116 COIIDENSER 91 1.3

REARRANPEWENT COOLER 21# 1.2

E-122 PREHEATER 171 2

PNENOL/NZ# VAPORIZER 431 7.52

E-122 CONDENSER 1.2#8 9.6

TEMPERED WATER COOLER 6.141 EA 9.5 EA

T-I#1 BASE NEATER 5# 1.14

T-112 BASE HEATER 51 #.I4

T-I#3 5ASE NEATER 51 #.#4

T-1#4 5ASE NEATER 58 1.14

T-185 9ASE NEATER 51 1.14

T-I#6 6ASE NEATER !w #.#I

T-I#7 5ASE NEATER 61 1.84

T-151 BASE NEATER 5# l .14

T-153 BASE HEATER 68 #.#I

DOUTNERH HEATER

SIZE

DlAr 9.0 FT EA HEIGHT: 39 FT

DIAI NElCNT:

5;: ;; EA

AREA NEAT LOAD (SO FT) tNl4 BTU/HI) ------- -----------

NEAT LOAD tWM ITUINR) --------m-v

16.5

MATERIAL OF CONSTRUCTION REMARKS _--___--__---_-----_------ -____--___--_-----_------------

316 ss CLAD PACKING: RESINS

316 ss CLAD PACKlNfi: RESINS

SNELL TUBES __----__-- m----m----

CARBON STL

CARBON STL

CARBON STL

CARBON STL

CARBON STL

CARBON STL

CARBON STL

CARBON STL

CARBON STL

CAR6ON STL

CARBON STL

CARBON STL

CARBON STL

CARBON ST1

CARBON STL

CARBON STL

CARSON STL

CARBON STL

CARBON STL

CARBON STL

CARBON STL

CARION STL

CARBON STL

CARBON STL

CARBON STL

CARBON STL

CARBON STL

CARBON STL

CARION STL

316 ss

316 SS

316 ss

316 ss

316 ss

316 ss

316 ss

316 ss

CARSON STL

316 ss

CARBON STL

316 ss

CARllON STL

CARBON STL

316 ss

316 so

316 ss

316 ss

316 ss

316 ss

316 SS

316 ss

316 SE

CARION STL

CARlON STL

CARBON STL

CARlON STL

CAROON STL

CARSON STL

CAR5ON STL

CARBON STL

CARBON STL

CARBON STL

CAR5ON STL

27 FT OF 1ON EXCHANGE RESIN PACKING

17 FT OF ION EXCNAN6E RESIN PACKING

18


TABLE 1.6 (CONTINUED)


MAJOR EPUIPNENT

CAPACITY: 120 MILLION LB (54.110 METRIC TONSFIVR POLYMER GRADE BISPHENOL-A AT 1.91 STREAM FACTOR

EWl:W::NT MATERIAL OF CONSTRUCTION REMARKS

________--____--___------------

2#,###

2.#U#

2,511

lO.NB8

lO.BN

3#.###

a#, 0n0

16#.##~

O#.##B

3.631

66 .##I

304 ss

304 ss

304 ss

314 ss

314 ss

314 ss

SP4 SE

314 ss

CARBON STL

RUBBER LND

CARBON STL

EPUIPPED WITH HEATING COIL

EOUIPPED WITH NEATINC COIL

EPUIPPED WITH HEATING COIL

EQUIPPED WITH HEATIN COIL

EOUIPPED WITH HEATING COIL

EDUIPPED WITH HEATING COIL

EQUIPPED UITN HEATING COIL

EOUIPPED WITH HEATING COIL

EPUIPPED UITN‘NEATINC COIL

VOLUME (CAL1 -------mm-mm

701

lO.I##

2.001

O.#N

2##

2.210

4.3##

1.111

211

IOU

b.##H

11,000

316 SS CLAD

316 ss CLAD

316 SS CLAD AGITATED AND JACKETED.

316 SO CLAD AGITATED AND JACKETED.

316 SS CLAD

CARSON STL

CAROON STL

316 SS CLAD

316 SS CLAD

316 SS CLAD

CAROON STL

CARSON STL

DIAMETER NE IENT (FT) (FT)

_------- ______

6.H 2#

CLEAVAGE COLUMN 6.0 21

TRAYS/ SNELL PACKING

------mm_______ __-__-----

316 SE CLAD 316 SS

CARBON STL CARBON STL

316 SS CLAD 316 SS

6 VALVE TRAYS. 24 IN SPACING

ACETONE COLUMN 6.5 1#5 47 VALVE TRAYS. 24 IN SPACING

5 VALVE TRAVS. 24 IN SPACING

EA

EA

316 ES 40 IN BASKET. OBNP

316 15 32 IN OASKET, SUNP

3#4 OS 125 So FT SINGLE DRUM FLAKER

CARBON STL 58 HP

CAROON STL S# HP

CARBON STL 6# HP

CARBON STL 21 HP

CAROON STL On HP

TANKS

T-101

T-102

T-IS3

T-104

T-105

T-106

T-107

T-151

T-152

T-153

T-154

VESSELS

v-101

v-102

v-103

V-104

v-105

V-106

V-107

V-108

v-109

V-l 10

v-111

v-112

COLUMNS

C-101

c-102

c-103

REACTOR FEED TANK

SISPNENOL A STORAGE

NO.1 BUFFER TANK

NO.2 5UFFER TANK

NO.3 SUFFER TANK

NO.1 PHENOL/N20 TANK

NO.2 PHENOL/N20 TANK

PHENOL STORAGE

ACETONE STORAGE

ALKALI TANK

TEMPERED WATER TANK

. CONCENTRATOR RECEIVER

CRYSTALLIZER

ADDUCT RESLURRV TANK

MELTER

PHENOL RECEIVER

C-101 REFLUX DRUM

C-W2 REFLUX DRUM

C-113 REFLUX DRUM

E-116 RECEIVER

PWENOLINLO RECEIVER

JACKET FOR V-113

JACKET FOR V-114

DEHYDRATION COLUMN

MISCELLANEOUS EOUIPMENT

M-I#lA,B NO.1 CENTRIFUGE

H-112 NO.2 CENTRlFUfE

W-IBSA-C FLAYER

n-104 E-183 VACUUM PUMP

M-186 E-l@6 VACUUM PUMP

M-106 C-In3 VACUUM PUMP

N-107 E-116 VACUUM PUMP

N-106 E-122 VACUUM PUMP

PUMPS

100 SECTION: 65. INCLUDING 34 OPEMTINC. 31 SPARES: 766 OPERATING BNP

19


Process Discussion

The ion exchange resin catalyst used for both condensation and

rearrangement is a sulfonated styrene-divinylbenzene cation exchange

resin. Specific grades cited in patent references are:

Amberlite ICE-100 Rohm and Haas Co.

Dowex so-x-4 Dow Chemical Co.

Permutit QH Permutit-Boby Co.

Chempro C-20 Chemical Process Co.

The resins used today are probably developments of these original

grades and they may therefore be sold under different grade designa-

tions. Although the resin is theoretically not consllmed in the reac-

tions, some degradation or gradual loss of activity can be expected.

Such deactivation may be partially reversible by suitable acid treat-

ment but it is likely that periodic replacement of the resin beds will

be required. The frequency of replacement is not known but the average

replacement rate is claimed to be less than 2 lb/metric ton of BPA prod-

uct (354251). For this evaluation we have assumed a nominal replace-

ment rate of 1.5 lblmetric ton at a resin value of $3/lb, or 0.2c/lb

BPA.

Although the process as designed is expected to yield polymer

grade product, much depends on the efficiency of crystal washing in the

adduct separation step and on the efficiency of separation of tars from

process eitreanus h the cleavage system. The efficiency of the crystal

separation step could be improved vlth a more complex wash system or

with an additional reslurry/centrifuge step. The efficiency of tar

eeparation could be improved by operating with a lower tar content in

the evaporator bottoms. stream (stream 33.). The direct result of this

mu&d be an increase In pheaol usage. If the total flow rate of this

rtream were increased by about 50x,, specific phenol consumption would

r&m ft- 0.878 lb/lb BPA to 0.91 lb/lb (see reference 354252).

20


For the production of less pure, epoxy grade BPA, the unit could

be operated with the cleavage system on bypass. This would result in

savings of about 5 gal/lb BPA of cooling water circulation and 1.4

lb/lb BPA of steam. The net savings in utilities costs would be lc/lb

BPA. Raw materials efficiencies would be virtually unaffected since

the condensation reactor effluent is an equilibrium mixture. By-

products removed with the BPA product would be replaced by increased

by-product production in the reactor. Phenol loss in the tars blowdown

stream would be unaltered.

Process Economics

Investment costs have been estimated at a PEP Cost Index of 425

(1958 - loo), corresponding approximately to 2nd quarter 1982 U.S. Gulf

Coast conditions. Table 1.7 shows the investment cost required for a

unit to produce 120 million lb/yr of high purity bisphenol-A.

Table 1.8 shows the production cost estimate for this plant based

on a 0.9 stream factor (7,884 hr/yr). The indicated raw material effi-

ciencies are 93.9% of theoretical on phenol and 88.3% on acetone. Raw

materials are valued at approximate May 1982 list prices. Raw materi-

als costs alone account for 61% of product value and fixed costs for

30%. In this table, credit is taken for the by-product tar stream at a

nominal fuel value of 11,000 Btu/lb, on the assumption that it is in-

cinerated in the utility boilers or Dowtherm@ heating system. The

indicated net production cost of 60&b and product value of 70c/lb

compare with current list prices of 66c/lb for polymer grade

bisphenol-A and 62c/lb for epoxy grade.

The variation of BPA production cost with plant capacity and with

plant operating level is shown in Figure 1.3.

21


TABLE 1.7


TOTAL CAPITAL INVESTMENT

CAPACITY: 120 MILLION LB (54,000 METRIC TONS)/YR POLYMER GRADE BISPHENOL-A AT 0.90 STREAM FACTOR

PEP COST INDEX: 425

BATTERY LIMITS EQUIPMENT, F.O.B. REACTORS COLUMNS VESSELS + TANKS EXCHANGERS FURNACES MISCELLANEOUS EQUIPMENT PUMPS

TOTAL

BATTERY LIMITS EOUIPMENT INSTALLED

CONTINGENCY, 25.0 PERCENT

8ATTERY LIMITS INVESTMENT

OFF-SITES, INSTALLED COOLING TOWER STEAM GENERATION INERT 6AS TANKAGE WAREHOUSE FACILITIES

UTILITIES + STORAGE

GENERAL SERVICE FACILITIES WASTE TREATMENT

TOTAL

CONTINGENCY, 25.0 PERCENT

OFF-SITES INVESTMENT

TOTAL FIXED CAPITAL

CAPACITY EXPONENT

---m------

COST UP DOWN ------------ ---s -s--

S 1,717,200 0.79 346,600 0.68

1.191,800 0.57 3;641,400 0.84

225.200 0.79 896;i00 0.63 609,600 0.59

----------- 9 8,629.000 0.75

0.74 0.53 0.46 0.75 0.79 0.49 0.44

0.65

8 23,540,000

5,885,000 -----------

S 29,425,000 0.68 0.58

S 2,010,800 0.96 0.77 2,324,000 0.82 0.62

199,800 0.50 0.55 657,700 0.68 0.59

1.377.000 0.90 0.90 -----------

S 6.569.000 0.86 0.79

6,022,000 1.5059000

s--m------- 8 14,096,000

39524,000 ---s---w---

S 17.620.000 0.79 0.70

8 47.0459000 0.72 0.62

22


TABLE 1.8


PRODUCTION COSTS

PEP COST INDEX: 425

VARIABLE COSTS UNIT COST CONSUMPTION/LB C/LB

------------- -------------- ------

RAW MATERIALS

PHENOL ACETONE CAUSTIC SODA (50 X1 CATALYST MAKEUP

GROSS RAW MATERIALS

BY-PRODUCTS

TAR AT FUEL VALUE

38 C/ LB 0.878 LB 33.36 32 C/ LB 0.288 LB 9.22

6.75 C/ LB 0.0041 LB 0.03 3.00 S/ LB 0.00068 LB 0.20

------ 42.81

4.6 Cl LB -0.067 LB -0.31

UNIT COST CONSUMPTION/LB CONSUMPTION/KG ---------s------ -------------- --------------

UTILITIES

COOLING WATER 5.4 C/l,000 GAL 53 GAL 442 LITERS 0.29 STEAM 7.00 S/l,000 LB 6.76 LB 6.76 KG 4.73 ELECTRICITY 3.6 C/KWH 0.075 KWH 0.166 KWH 0.27 NATURAL GAS 4.17 S/MM BTU 1.549 BTU 860 KCAL 0.65 INERT GAS, LO P 73 C/1,000 SCF 0.657 SCF 38.8 LITERS 0.05

-e---m TOTAL UTILITIES 5.99

23


TABLE 1.8 (CONTINUED)


PRODUCTION COSTS

PEP COST INDEX: 425

(1) (2) CAPACITY (MILLION LB/YR) 60 120 240

------ ------ ------ INVESTMENT (S MILLION)

BATTERY LIMITS 19.7 29.4 47.3 OFF-SITES 10.9 17.6 30.4

----mm ------ ------ TOTAL FIXED CAPITAL 30.6 47.0 77.7

SCALING EXPONENTS 0.62 0.72

PRODUCTION COSTS (C/LB)

RAW MATERIALS BY-PRODUCTS UTILITIES

VARIABLE COSTS (3)

42.81 42.81 42.81 -0.31 -0.31 -0.31

5.99 5.99 5.99 ---w-- ------ ------

48.49 48.49 48.49

OPERATING LABOR, 7/SHIFT, S17.50/HR MAINTENANCE LABOR, 2 PCT/YR OF BL INV CONTROL LAB LABOR, 20 PCT OF OP LABOR

LABOR COSTS

1.79 0.89 0.45 0.66 0.49 0.39 0.36 0.18 0.09

------ -s---- ------ 2.81 1.56 0.93

MAINTENANCE MATERIALS, 2 PCT/YR OF BL INV 0.66 0.49 0.39 OPERATING SUPPLIES, 10 PCT OF OP LABOR 0.18 0.09 0.04

TOTAL DIRECT COSTS 52.14 50.63 49.85

PLANT OVERHEAD, 80 PCT OF LABOR COSTS TAXES AND INSURANCE, 2 PCT/YR OF TFC DEPRECIATION, 10 PCT/YR OF TFC

PLANT GATE COST (3)

2.24 1.25 0.74 1.02 0.78 0.65 5.10 3.92 3.24

------ -mm--- ------ 60.50 56.58 54s.48

G+A, SALES, RESEARCH. 5 PCT OF SALES

NET PRODUCTION COST

3.30 3.30 ------ -s----

63.80 59.88

3.30 -e-e--

57.78

ROI BEFORE TAXES, 25 PCT/YR OF TFC 12.75 9.79 8.09 ------ m----- ------

PRODUCT VALUE 76.55 69.67 65.87

---------_---------------------------------------------------- (1) OF POLYMER GRADE BISPHENOL-A (2) BASE CASE (3) FOR BASE CASE ONLY: MAY BE DIFFERENT FOR OTHER CAPACITIES.

24


l

Figure 1.3

BISPHENOL-A FROM PHENOL AND ACETONE

WITH AN ION EXCHANGE RESIN CATALYST

Effect of Operating Level and Plant Capacity on Production Cost

I I I I

PEP Cod Index: 425

I I I I

0.5 0.6 0.7 0.8 0.9 1.0

OPERATING LEVEL, fmction of design capacity

25


Comparison with-the HC1 Catalyzed Process

Three significant differences are immediately apparent in compar-

ing this evaluation with our 1972 evaluation of the HCl catalyzed pro-

cess. Firstly, the 1972 evaluation is based on a phenol-to-acetone

reactor feed ratio of 4: 1 and an acetone conversion of 99% per pass,

compared with the 1O:l ratio a.nd 50.5% conversion used in this evalua-

tion. Thus the reactor feed flowrate in this evaluation is 4.4 times

greater per pound of BPA than in the earlier evaluation and utilities

consumptions (particularly that of steam) are correspondingly higher.

The phenol-to-acetone feed ratio affects the overall yield of BPA and

the formation of by-products, Including color forming compounds. Low

phenol-to-acetone ratios lead to low BPA yields and high by-product

formation. Industry reviewers of the 1972 evaluation indicated that a

feed ratio of 8:l might be necessary in order to achieve the required

EPA yields and low color body formation. We believe that an 8:l ratio

would be used in designing an HCl catalyzed process to meet 1982 raw

material costs and product color standards.

The second significant difference between the two evaluations

appears in the final BPA crystal recovery operation. The present eval-

uation uses a two stage countercurrent centrifuggtion and washing

scheme for crystal recovery, whereas the 1972 evaluation uses a single

ccntrifugation st8ge. It is certain that the product color standard

achievable with the 1972 scheme is significantly poorer than that

achievable with the present scheme, particularly in combination with

the 10-r phenol to acetone reactor feed ratio. The product purity and

color standards achievable with the single centrifugation stage are

probably comparable with those achievable in the resin catalyzed pro-

cess operating with the cleavage system on bypass, i.e., producing BPA

of higher purity than is uslrally required for epoxy resin applications

but not sufficiently high to meet current polycarbonate grade specifl-

C8tiOU8. We believe th8t 8n HCl catalyxed process designed to produce

pdyccrrbonate grad8 BPA of a standard equivalent to that produced by

th8 resin catalyzed process would require significantly improved sol-

vent purification facilities and a second crystal washing and centri-

fugation stage. 26


The third significant difference between the two evaluatfons

appears in the environmental aspects of the two processes. The resin

catalyeed process is essentially nonpolluting. The scheme used to

separate reaction water from the process gives an effluent with a low

phenol content; it is suitable for feeding directly to biological treat-

ment . The HCl catalyzed process as evaluated in 1972 would not meet

today’s environmental standards without significantly higher investment

in effluent treatment to control phenol release in aqueous streams and

benzene vapors in operating areas.

To compare the two processes on a common basis, we have updated

the 1972 investment and operating cost estimates as shown in Table 1.9.

The comparison assumes production of a good epoxy grade product. The

resin catalyzed process operates with the cleavage system on bypass.

The HCl catalyzed process operates with an 8:l phenol-to-acetone feed

ratio and requires increased investment for environmental control.

The table shows that, although the HCl catalyzed process has a higher

investment cost, net production costs are essentially the same for

epoxy grade BPA. For a given product value, the percentage return on

investment is higher for the resin catalyzed process.

For polycarbonate grade BPA, production costs for the resin cata-

lyzed process are lc/lb higher than shown in Table 9.1 but we estimate

the increase for the HCl catalyzed process to be about 2-3c/lb in pro-

duction cost and 4-5c/lb in product value.

27


Table 1.9

PROCESS COMPARISON FOR EPOXY GRADE BPA (120 Million lb/yr)

Battery limits investment, $ million

Offsites investment, $ million

Total fixed capital, $ million

Production costs, c/lb

Raw materials Utilities

29.4 37.7

17.6 17.3

47.0 55.0

42.8 43.5 4.7 2.4

Variable costs 48.5 45.9

Fixed costs 1.1 2.4

TOTAL DIRECT COSTS 49.6 48.3

Overhead, depreciation, etc. 9.3 10.2

NRT PRODUCTION COST 58.9 58.5

R.O.I. at 25% of T.F.C. 9.8 11.5

PRODUCT VALUE 68.7 70.0

Resin Catalyzed

HCl Catalyzed

28


FIGURE 1.2 (sheet 1 of 2)

BISPHENOL-A FROM PHENOL AND ACETONE WITH ION EXCHANGE RESIN CATALYST (UNION CARBIDE PROCESS)

T-151 Plmol Stamp

- .

LP stm VlpXirCr To T-101

(

15 PI0 ll5'F I 1

Wohr to WDlh

M E-114

Rdoiler

R-101 AM

BPA kaaon

ConoMtmtor E-104 ikhmlkl Tb

: v-w

i

:Aota i

: :

i

3

2 :

i

: T-152

i Amtax Staqc

: Tallk

i

E-102

Cryxtdlinr Crpr)allizer r-,--

T(

M E-102

cmuntmtm M v-101

GnantnJtol bainr

ToV-102

Fmm V-IOJ z&l M-102

No.2 Centrifuge

: P

i 8 : Bi+wt&A :

PdcingordStomga i

i

cw : : : :

i 3 MPStm 15Opig

M-102 AthwC

Fldr*n

LP stm I5 pig

T-102

Rirphenol A j . TOM-lOl*ap &7'F s

: TcmpredCmling Wdcr

E-l&

Phenol Striper

.- Phenol Receiver

FIGURE 1.2 (shot 2 of 2)

BISPHENOL-A FROM PHENOL AND ACETONE WITH ION EXCHANGE RESIN CATALYST (UNION CARBIDE PROCESS)

I

: : i T-101 c-103 E-116 : :

No. 2 Cl*waBa cad-r

i BuffirTd C&M

i V-MB

Rdluxbnm Fmm M-101 MB

E-119 T-IQ E-120

thdemmr ND.3 RuffwTd

E-12,

PdWbr

E-122

ckndmmer

T-1W ;

No. 2 ;

PhB -.- -

ToVocwn

1-b

R

3agF v-110

MP% MQnm He 150 mla

~l/n)O Td i

: : :

;

:

i

i

: :

l

:

i

ii3 ; : : : i i

LP5tm .

I5 pig i : . : : .

ll5'F 175'F

t

i ; I T-153 E-115 E-117

Tarto Fuel

To T-101

E-118 v-109

---1 I

F

lSe"F 8

i ;

Alkali Tack Pmhmter Rekoiler BPAEvapomtw Rmminr

i

: : : . - . : i

Tempd Coding Water

@I 15B'F

ci

lLPShn ; I5 pig T

D&m10-A 300°F

~-------A x T~toDinpd

R-103 A rhw C T-106

Reanolgcn*nt No.1 Re~etm F'hcml/l120 Tank

E-17.2

P&d&O Vqmrlzer

v-110

Phmo1&0 Receiver

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