Anionic Styrene Polymerization in a Continuous Stirred-Tank Reactor

14 Anionic Styrene Polymerization in a Continuous StirredTank Reactor MICHAEL N. TREYBIG1 and RAYFORD G. ANTHONY

Downloaded by UNIV OF SYDNEY on May 3, 2015 | http://pubs.acs.org Publication Date: July 31, 1979 | doi: 10.1021/bk-1979-0104.ch014

Department of Chemical Engineering, Texas A&M College Station, TX 77843

University,

In the design, o p t i m i z a t i o n , or c o n t r o l of a p o l y m e r i z a t i o n r e a c t o r , a mathematical model which adequately represents the process i s d e s i r a b l e . In the f o r m u l a t i o n of such a model, i n f o r mation i s r e q u i r e d on both the k i n e t i c s of the s p e c i f i c r e a c t i o n and the mixing p a t t e r n of the r e a c t i o n v e s s e l used. For c o n t i n uous s t i r r e d tank r e a c t o r s , the assumption of p e r f e c t or micromixing i s f r e q u e n t l y made and the corresponding design equations used to estimate the r e a c t o r ' s performance. However, i n many l a r g e s c a l e i n d u s t r i a l p o l y m e r i z a t i o n processes the occurrence of imperfect mixing or segregation is more probable. In the case of a segregated p o l y m e r i z a t i o n r e a c t o r , design equations are r e q u i r e d which g i v e a d i f f e r e n t molecular weight d i s t r i b u t i o n from that obtained f o r the micro-mixed case. Since the p r o c e s s i b i l i t y and mechanical p r o p e r t i e s of a polymer f r a c t i o n are s t r o n g l y dependent on the shape of the molecular weight d i s t r i b u t i o n , i t is important to know the e f f e c t s of imperfect mixing on the shape of the molec u l a r weight d i s t r i b u t i o n and the degree of imperfect mixing occurring i n a reactor. Scope and O b j e c t i v e s The o b j e c t i v e s of t h i s work were: to study the e f f e c t of segregated mixing i n a s t i r r e d tank flow r e a c t o r on the molecular weight d i s t r i b u t i o n of p o l y s t y r e n e ; to determine the degree of segregation, i f any, o c c u r r i n g i n a bench s c a l e l a b o r a t o r y reactor; and to evaluate the usefulness of r e a c t o r flow models based on micro- and macro-mixing i n a constant-flow, s t i r r e d - t a n k r e a c t o r . Styrene was polymerized i n a bench s c a l e l a b o r a t o r y r e a c t o r w i t h p o l y s t y r y l l i t h i u m seed i n benzene s o l v e n t . A seeded polymerizat i o n system was chosen to s i m p l i f y the k i n e t i c d e s c r i p t i o n of the process compared with a system i n v o l v i n g simultaneous i n i t i a t i o n and propagation r e a c t i o n s . Mathematical models based on concepts of micro- and macro-mixing i n a s t i r r e d tank r e a c t o r were de1

Current address: S h e l l Development Company, Westhollow Research Center, Houston, Texas 0-8412-0506-x/79/47-104-295$08.00/0 © 1979 American Chemical Society In Polymerization Reactors and Processes; Henderson, J., et al.; ACS Symposium Series; American Chemical Society: Washington, DC, 1979.

POLYMERIZATION REACTORS AND PROCESSES

296

veloped. These models u t i l i z e k i n e t i c d e s c r i p t i o n s of t h i s p o l y mer system from previous s t u d i e s of the system, as w e l l as data obtained i n t h i s i n v e s t i g a t i o n . R e s u l t s from the l a b o r a t o r y experimentation and mathematical s i m u l a t i o n were compared. The comparison was used to determine the s u i t a b i l i t y of the mathematic a l s i m u l a t i o n f o r modeling the p o l y m e r i z a t i o n process.


Theory Reaction Mechanism. The r e a c t i o n mechanism of the a n i o n i c s o l u t i o n p o l y m e r i z a t i o n of styrene monomer using n - b u t y l l i t h i u m i n i t i a t o r has been the subject of considerable experimental and t h e o r e t i c a l i n v e s t i g a t i o n (1-8). The p o l y m e r i z a t i o n process occurs as the a l k y l l i t h i u m a t t a c k s monomeric styrene to i n i t i a t e a c t i v e s p e c i e s , which, i n turn, grow by a stepwise propagation r e a c t i o n . T h i s p o l y m e r i z a t i o n r e a c t i o n i s c h a r a c t e r i z e d by the production of s t r a i g h t c h a i n a c t i v e polymer molecules ( " l i v i n g " polymer) without termination, branching, or t r a n s f e r r e a c t i o n s . The stoichiometry of the p o l y m e r i z a t i o n process may be represented by the simple r e a c t i o n scheme: I + M ->P

(1)

1

P

x

P

+ M + P

2

+ M + P

j + 1

(2) j = 2, oo

(3)

However, the mechanisms by which the i n i t i a t i o n and propagat i o n r e a c t i o n s occur are f a r more complex. Dimeric a s s o c i a t i o n of p o l y s t y r y l l i t h i u m i s reported by Morton, et a l . (9) and i t i s g e n e r a l l y accepted that the r e a c t i o n s are f i r s t order with respect to monomer c o n c e n t r a t i o n . U n f o r t u n a t e l y , the existence of a s s o c i ated complexes of i n i t i a t o r and p o l y s t y r y l l i t h i u m as w e l l as p o s s i b l e cross a s s o c i a t i o n between the two species have negated the determination of the exact p o l y m e r i z a t i o n mechanisms (8, 10, 11, 12, 13). I t i s t h i s high degree of complexity which n e c e s s i t a t e s the use of e m p i r i c a l r a t e equations. One such e m p i r i c a l r a t e expression f o r the a u t o - c a t a l y t i c i n i t i a t i o n r e a c t i o n f o r the a n i o n i c p o l y m e r i z a t i o n of styrene i n benzene solvent as reported by Tanlak (14) i s given by: Rj « k

I M (1 + * P

x

3 T

)

(4)

Tanlak found the f o l l o w i n g r e l a t i o n s f o r the propagation r e a c t i o n s and monomer consumption: R . = a(P. - - P.)M J J-l J Rp = aP M p

(5)

p

T

In Polymerization Reactors and Processes; Henderson, J., et al.; ACS Symposium Series; American Chemical Society: Washington, DC, 1979.

(6)

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Styrene

Polymerization

297

where:

a


h + Vh

+ 2Kj,-i

— P

(8) T

S i m i l a r r e s u l t s f o r the propagation r e a c t i o n s were obtained by Timm and Kubicek (15). In t h i s work, the c h a r a c t e r i s t i c " l i v i n g " polymer phenomenon was u t i l i z e d by preparing a seed polymer i n a batch r e a c t o r . The seed polymer and styrene were then fed to a constant flow s t i r r e d tank r e a c t o r . T h i s procedure allowed use of the lumped parameter r a t e expression given by Equations (5) through (8) to d e s c r i b e the p o l y m e r i z a t i o n r e a c t i o n , and eliminated complications i n v o l v e d i n d e s c r i b i n g simultaneous i n i t i a t i o n and propagation r e a c t i o n s . Mixing Models. The assumption of p e r f e c t or micro-mixing i s f r e q u e n t l y made f o r continuous s t i r r e d tank r e a c t o r s and the ensuing r e a c t o r model used f o r design and o p t i m i z a t i o n s t u d i e s . For w e l l - a g i t a t e d r e a c t o r s with moderate r e a c t i o n r a t e s and f o r r e a c t i o n media which are not too v i s c o u s , t h i s model i s o f t e n j u s t i f i e d . Micro-mixed r e a c t o r s are c h a r a c t e r i z e d by uniform concent r a t i o n s throughout the r e a c t o r and an exponential residence time d i s t r i b u t i o n function. The concept of a w e l l - s t i r r e d segregated r e a c t o r which a l s o has an exponential residence time d i s t r i b u t i o n f u n c t i o n was i n t r o duced by Dankwerts (16, 17) and was elaborated upon by Zweitering (18). In a t o t a l l y segregated, s t i r r e d tank r e a c t o r , the feed stream i s envisioned to enter the r e a c t o r i n the form of macromolecular capsules which do not exchange t h e i r contents with other capsules i n the feed stream or i n the r e a c t o r volume. The capsules act as batch r e a c t o r s with r e a c t i o n times equal to t h e i r residence time i n the r e a c t o r . The r e a c t o r product i s thus found by c a l c u l a t i n g the weighted sum of a s e r i e s of batch r e a c t o r products with r e a c t i o n times from zero to i n f i n i t y . The weighting f a c t o r i s determined by the residence time d i s t r i b u t i o n f u n c t i o n of the constant flow s t i r r e d tank r e a c t o r . Many mixing models which u t i l i z e the s i m p l i f i e d concepts of micro-mixing and segregation have been introduced. Most notable of these are the two-environment models of Chen and Fan (19), Kearns and Manning (20), and others (21, 22), and the d i s p e r s i o n models of Spielman and L e v e n s p i e l (23), and Kattan and A d l e r (24). Since p o l y m e r i z a t i o n r e a c t i o n s i n continuous s t i r r e d tank r e a c t o r s are o f t e n c a r r i e d out under c o n d i t i o n s of high v i s c o s i t y not conducive to micro-mixing, t h e o r e t i c a l and experimental i n v e s t i g a t i o n s have been made to determine the e f f e c t s of segregat i o n on the molecular weight d i s t r i b u t i o n f o r v a r i o u s polymer systems (25, _26, 27, 28). Ahmad (27) studied the e f f e c t of mixing on the molecular weight d i s t r i b u t i o n of p o l y i s o p r e n e and Tadmor and Biesenberger (28) studied the e f f e c t of segregation on




298

molecular weight d i s t r i b u t i o n s . For a stepwise a d d i t i o n nonterminating p o l y m e r i z a t i o n i n a segregated constant flow s t i r r e d tank r e a c t o r , these authors found that a polymer would be produced with a molecular weight d i s t r i b u t i o n that i s broader than that of a batch r e a c t o r but more narrow than that of a micro-mixed reactor. The design equations f o r the mixing c o n d i t i o n s considered i n t h i s paper a r e presented i n Table I . The micro-mixed model was modified to i n c l u d e the e f f e c t of i n a c t i v e or dead polymer and the e f f e c t of part of the r e a c t a n t s passing through the r e a c t i o n zone without r e a c t i n g . The equations f o r the w e l l - s t i r r e d segregated r e a c t o r and f o r the batch r e a c t o r a r e a l s o presented i n Table I . F i g u r e 1 i l l u s t r a t e s the growth c h a r a c t e r i s t i c s of polymer chains i n micro-mixed and segregated w e l l - s t i r r e d reactors. For the micro-mixed CFSTR the growth l i n e s from a seed polymer are l i n e a r , w h i l e from the segregated CFSTR they e x h i b i t curvature due to the change of monomer concentration as a segregated lump passes through the r e a c t o r . Experimental Reactor Design. The continuous polymerization r e a c t i o n s i n t h i s i n v e s t i g a t i o n were performed i n a 50 ml pyrex g l a s s r e a c t o r . The mixing mechanism u t i l i z e d two mixing i m p e l l e r s and a Chemco magnet-drive mechanism. The g l a s s r e a c t o r , shown i n F i g u r e 2, has s i n g l e i n l e t and o u t l e t ports and one thermocouple p o r t . The r e a c t o r s h e l l i s made from a s e c t i o n of pyrex tubing 4.4 cm OD and 4.0 cm ID. The i n l e t and o u t l e t ports a r e made from 1/4 i n OD x 1.0 mm ID c a p i l l a r y tube. The thermocouple port i s made from 1/4 i n OD x 5/32 i n ID g l a s s tubing. Glass to s t a i n l e s s connections a r e made using 1/4 i n s t a i n l e s s Swagelok f i t t i n g s w i t h T e f l o n f r o n t f e r r u l e s and a 1/4 i n x 0.065 i n v i t o n 0 - r i n g . The s t a i n l e s s f i t t i n g s used at the i n l e t and thermocouple p o r t s are 1/4 i n to 1/16 i n reducers. The i n l e t port f i t t i n g i s connected by a short s e c t i o n of 1/16 i n tubing to a 1/16 i n s t a i n l e s s t e e which i s used to pre-mix the monomer and living-polymer feed streams. The thermocouple port allows entrance of a type "T" thermocouple i n a 1/16 i n s t a i n l e s s sheath. To f i l l the v o i d between the thermocouple and the tubing w a l l of the thermocouple p o r t , a plug of T e f l o n 5/32 i n OD with a 1/16 i n ID a x i a l hole i s placed i n the thermocouple p o r t . The f i t t i n g used at the r e a c t o r o u t l e t port i s a 1/4 i n to 1/8 i n reducer and i s connected to the r e a c t o r e f f l u e n t l i n e of 1/8 i n t e f l o n tubing. Two i m p e l l e r s a r e included i n the r e a c t o r c o n f i g u r a t i o n shown i n F i g u r e 2. A three-bladed t u r b i n e with 45° p i t c h and blades 1/8 i n x 5/8 i n i s mounted on the i m p e l l e r s h a f t a t the top of the r e a c t o r . A three-bladed p r o p e l l e r w i t h 45° p i t c h i s mounted a t the bottom of the i m p e l l e r s h a f t a t approximately two-thirds of the r e a c t o r depth. 1

!



14.

TREYBIG AND ANTHONY

Anionic

Styrene

Polymerization

Figure 1. Growth characteristics for seed polymer in CFSTR environments: growth characteristics for polymer chains in a micro-mixed environment; growth characteristics for polymer chains in a segregated environment


299

(a) (b)


J

=

(

D = w

D° n°

D° D ° w n

+

r

min

+ aM©

aM

J

= 1 - exp

-T)

j ( t )

-01n(l

D = D° n n

T.

t

j ( T )

exp P. = P. J min rP P~ = p>o

^rnin 1 + aM0

M 1 + aOP,,

min

M =

Micro-mixed

D°

n

D °

aM©

T

]

(T)dT

1 + 2D n

P°

f Jm

CSTR

Aj

j

Y

=

D = D w w

D

n

fo°

^

n D >

T

1

+ aMO*

aM A

3

A

= P e x p [ - t . /0] + Aj min

A

min 1 + aMQA A

T

= P° + P . Dj Aj

. mm

J

A

M 1 + a0 P

A

1 + 2

P

j

t )

e

X

p

[

-

t

/

]

d

t

+ 2D

G

CFSTR

D - D n n

(

(Dead) P o l y m e r i n M i c r o - m i x e d u

P° = P° Aj ( t ) (j+aM* t)

P

P

j

M =

Inactive

The d e r i v a t i o n s o f t h e e q u a t i o n s a r e g i v e n b y T r e y b i g ( 3 2 )

TABLE I . REACTOR MIXING MODELS


U


1

n

+

P

R T

T

B

JLE.

R

jm-

n

A

w

Equations f o r D , D and D /D r. ¥ , are same as f o r equations f o r dead polymer with replaced by .

J ( t ) = j + aM t R J - j min

T

K

1 + aM 0

a 0

R

M°(l + aO P ) pressure regulator; (Q) check valve; (—•) thermocouple; (1) compressed helium; (2, 3) molecular sieve columns; (4, S) benzene (solvent) tank; (5,M) styrene (monomer) tank; (6,Sd) seed polymer tank; (7,8) rotameters; (9) teflon tubing; (10) premixing tee; (11, R) reactor; (12) temperature bath; (13) heating coil; (14) thermometer probe; (15, TI) temperature indicator (16, TC) temperature controller


14.

TREYBIG AND ANTHONY

TABLE I I I .


Run No.

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15

Styrene

311

Polymerization

OPERATING CONDITIONS FOR CFSTR RUNS 1-15

Feed Concent r a t i o n of Styrene, M° ( gm-mo l e / liter)

0.856 0.685 0.729 3.583 2.309 2.309 2.308 3.456 3.458 4.147 4.157 2.678 2.714 2.656 2.675

Anionic

Feed ConcenMean Mixing t r a t i o n of Residence Speed Seed Polymer, Time (min) (RPM) P ( gm-mo 1 e / liter)

Reaction Temperature (°C)

T

0.0081 0.0075 0.0075 0.0034 0.0052 0.0052 0.0052 0.0035 0.0035 0.0024 0.0024 0.0067 0.0066 0.0066 0.0066

25.4 31.6 31.6 22.4 19.5 19.9 19.6 15.3 15.4 15.3 14.6 21.9 22.2 21.7 21.3

1000 1000 500 1000 1000 500 750 1000 500 1000 500 1000 500 300 250

24.8±0.25 23±2 22±2 26.7±0.25 22±2 20±2 20±2 25.7±0.25 26.2±0.25 25.6±0.25 26.410.25 3611 3711 3811 3510.25

Figure 7. Theoretical polymer distributions, based on kinetic description of Tanlak (14) for micro-mixed and totally segregated CFSTRS with polymer feed (CFSTRS: CMO = 0.5M; PT = 0.01M; 6 = 20.0 min; XM = 0.70)



312

l a t e r f i l t e r e d and then analyzed u s i n g the GPC. A f t e r the polymer samples were taken, the i m p e l l e r was stopped to allow e s t i m a t i o n of the volume of gas which c o l l e c t e d i n the r e a c t o r (due to degassing of helium from the feed stream) during the run. T h i s r e d u c t i o n i n the e f f e c t i v e r e a c t i o n volume of the r e a c t o r was noted and the gas was removed from the r e a c t o r through the e x i t port by t i l t i n g the r e a c t o r . Subsequent runs were then made by a d j u s t i n g the feed f l o w r a t e s and then the mixing speed w i t h the r e a c t o r i n i t i a l l y f i l l e d the r e a c t i o n medium from the previous run.


Results The experimental monomer conversion and degrees of polymeriz a t i o n f o r the continuous r e a c t o r runs are given i n Table IV. Experimental Values f o r the Lumped-Parameter Propagation Rate Constant. The experimental values f o r the lumped-parameter propagation r a t e constant were determined assuming a micro-mixed r e a c t o r , styrene c o n c e n t r a t i o n and s o l v i n g f o r a . The r e s u l t s f o r Runs 1-15 are included i n Table V. The value f o r the propagation constant based on a segregated model are the same as that f o r a micro-mixed model. For the case of a micro-mixed r e a c t o r with dead-polymer or a micro-mixed by-pass r e a c t o r , the true v a l u e of a would be l a r g e r than the value reported f o r the micro-mixed case by f a c t o r s of 1/D and 1/[1 - a0P /O ], r e s p e c t i v e l y . This would compensate f o r the decrease i n monomer conversion a s s o c i a t e d with dead-polymer and by-passing. T

B

R

C a l c u l a t e d Degrees of P o l y m e r i z a t i o n . The c a l c u l a t e d degrees of p o l y m e r i z a t i o n f o r the micro-mixed, segregated, and micro-mixed r e a c t o r w i t h dead-polymer models are given i n Table VI. Values f o r the lumped parameter propagation r a t e constant used i n the simulations were c a l c u l a t e d such that the monomer conversions f o r the models would be the same as that f o r the l a b o r a t o r y r e a c t o r . Therefore, the number average degrees of p o l y m e r i z a t i o n f o r each model i s equal to the experimentally observed number average. For the micro-mixed r e a c t o r w i t h dead-polymer model, average values of the f r a c t i o n dead-polymer, ^DAvg' * f o r each of the d i f f e r e n t seed mixtures. (Note that = 1 - fy^) The average values of -^ f o r each seed were determined by averaging the values of cf> r e q u i r e d to match the experimental number and weight average degrees of p o l y m e r i z a t i o n . The value of f o r each run was found by s o l v i n g the equation f o r D /D i n Table I f o r and s u b s t i t u t i n g the experimental values f o r the average degrees of p o l y m e r i z a t i o n . The values p c a l c u l a t e d f o r each run are given i n Table V I I . In the c a l c u l a t i o n of DAvg Seed I I the values of f o r Runs 3 and 4 were not used. The values f o r DACT g i n Table V I I I . w

e

r

e

u s e c

D

D

D

w

n

D

f

D

a

r

e

i

v

e

n

V


o

r

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313

TABLE IV. EXPERIMENTAL MONOMER CONVERSIONS AND DEGREES OF POLYMERIZATION FOR CFSTR RUNS 1-15


Run Number 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15

rpm

X m

1000 1000 500 1000 1000 500 750 1000 500 1000 500 1000 500 300 250

0.524 0.586 0.411 0.585 0.446 0.446 0.428 0.377 0.423 0.311 0.333 0.726 0.704 0.698 0.630

n

D w

206 200 187 768 346 346 338 524 568 586 726 424 421 412 387

289 267 263 1411 579 547 558 995 1061 1598 1701 691 698 677 609

D

D /D w

n

1.41 1.34 1.41 1.84 1.67 1.58 1.65 1.90 1.87 2.33 2.34 1.63 1.66 1.64 1.57

TABLE V. EXPERIMENTAL VALUES FOR LUMPED PARAMETER PROPAGATION RATE CONSTANT BASED ON MICRO-MIXED CFSTR MODEL

Run Number

Rate Constant, a

1 2 3 4 5 6 7 8

5.31 5.98 2.95 18.65 7.97 7.82 7.40 11.46

Run Number

9 10 11 12 13 14 15

Rate Constant* a 13.74 12.33 14.28 18.08 16.18 15.91 12.09

a i s a f u n c t i o n of temperature and polymer c o n c e n t r a t i o n as given by Equation 3.



u

n

n

206 200 187 768 346 346 338 524 568 686 726 424 421 412 387

D

w

289 267 263 1411 579 547 558 995 1061 1598 1701 691 698 677 609

w

D

1.41 1.34 1.41 1.84 1.67 1.58 1.65 1.90 1.87 2.33 2.34 1.63 1.66 1.64 1.57

D /D n

Experimental

w 267 253 237 1281 484 484 470 811 895 1122 1199 636 630 614 567

D

CFSTR

f o r a l l runs,

1.30 1.27 1.27 1.67 1.40 1.40 1.39 1.55 1.58 1.64 1.65 1.50 1.50 1.49 1.47

D /D w n

Micro-mixed

w 256 242 233 910 412 412 404 661 708 919 967 467 468 458 438

D

1.24 1.21 1.24 1.18 1.19 1.19 1.20 1.26 1.25 1.34 1.33 1.10 1.11 1.11 1.13

D /D w n

Segregated CFSTR

289 264 244 1687 576 576 556 1030 1147 1464 1573 696 690 671 617

1.40 1.32 1.31 2.20 1.67 1.67 1.65 1.97 2.02 2.13 2.17 1.64 1.64 1.63 1.60

Micro-mixed CFSTR With Dead Polymer F r a c t i o n D D /D w w n

COMPARISON OF EXPERIMENTAL AND CALCULATED DEGREES OF POLYMERIZATION

(D ) - - ^ , = (D ) . , n calculated n experimental

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15

, Number

R

TABLE V I .


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Styrene

315

Polymerization


TABLE V I I . FRACTION DEAD POLYMER REQUIRED TO MATCH EXPERIMENTAL DEGREES OF POLYMERIZATION USING A MICRO-MIXED REACTION WITH DEAD POLYMER

Run Number

F r a c t i o n Dead Polymer, 4>

1

0.427 0.328 0.599 0.115 0.292 0.217 0.291 0.253

D

2 3 4 5 6 7 8

TABLE V I I I .

Seed Number

Run Number

9 10 11 12 13 14 15

F r a c t i o n Dead Polymer, D

0.211 0.360 0.352 0.121 0.146 0.144 0.112

AVERAGE FRACTION DEAD POLYMER FOR SEED MIXTURES

CFSTR Run Number

Average F r a c t i o n Dead Polymer, ^D^vg

I II III

1 2-11 12 - 15

0.427 0.288 0.131


POLYMERIZATION

316

REACTORS AND PROCESSES


C a l c u l a t e d Molecular Weight D i s t r i b u t i o n s . The c a l c u l a t e d weight f r a c t i o n d i s t r i b u t i o n s f o r the micro-mixed, segregated, and micro-mixed r e a c t o r with dead-polymer models f o r Runs 2, 5, 8, 10 and 12 a r e shown along with the experimental d i s t r i b u t i o n s i n Figures 8 through 12. These f i g u r e s i l l u s t r a t e the e f f e c t s of micro-mixing and segregation on the weight f r a c t i o n d i s t r i b u t i o n as w e l l as the a b i l i t y of the models to simulate the experimental d i s t r i b u t i o n s a t d i f f e r e n t degrees of p o l y m e r i z t i o n . The c a l c u l a t e d mole f r a c t i o n d i s t r i b u t i o n s f o r Runs 8 and 12 a r e shown with the experimental d i s t r i b u t i o n s i n Figure 13 and 14. Streaking Observed i n Reaction Medium During Continuous Polymerizations. Non-uniformities i n the r e a c t o r contents i n the form of streaks were observed during continuous polymerizations at mixing speeds of l e s s than 1,000 rpm. At mixing speeds of 1,000 rpm, the r e a c t o r appeared to be d i v i d e d i n t o two homogeneous mixing zones: one occupying the upper h a l f of the r e a c t o r and the other occupying the lower h a l f . At lower mixing speeds f o r Runs 3 and 7, a feed stream or s t r e a k was observed to pass from the r e a c t o r feed port over the blades of the lower i m p e l l e r and down i n t o the center of the lower i m p e l l e r . Durings Runs 6, 8 and 13, a d d i t i o n a l streaks were observed i n the lower mixing zone. During Runs 9, 11, 14 and 15, c o n s i d e r a b l e s t r e a k i n g was observed i n both the upper and lower mixing zones. I t should a l s o be pointed out that during Run 11 (at low rpm and high degrees of polymerization) i n which an a p p r e c i a b l e amount of degassing occurred, small bubbles were o c c a s s i o n a l l y observed to t r a v e l from the top of the r e a c t o r i n t o the lower mixing r e g i o n . This i s an i n d i c a t i o n that a well-mixed c o n d i t i o n was achieved to a t l e a s t a macroscopic l e v e l . Discussion The 50 ml g l a s s r e a c t o r proved to be w e l l - s u i t e d f o r the procedures implemented i n t h i s i n v e s t i g a t i o n . The small s i z e of the r e a c t o r allowed e f f i c i e n t use of the m a t e r i a l s r e q u i r e d f o r both the residence time d i s t r i b u t i o n s t u d i e s and f o r the continuous p o l y m e r i z a t i o n experiments. V i s u a l i n s p e c t i o n of the r e a c t o r contents during o p e r a t i o n proved v a l u a b l e i n determining imperfections i n the mixing p a t t e r n during RTD s t u d i e s as w e l l as f o r observing s t r e a k i n g and bubble formation i n the r e a c t o r during p o l y m e r i z a t i o n s . The only disadvantage a s s o c i a t e d with the small g l a s s r e a c t o r i s the increased care i n handling r e q u i r e d over that o f a small s t a i n l e s s s t e e l r e a c t o r . E r r o r s i n the Temperature Measurement During Polymerizations Runs. The i n t e r n a l r e a c t o r temperatures measured during Runs 2, 3, 5, 6 and 7 were found to be i n c o n s i s t e n t when a heat balance was made on the r e a c t o r . E r r o r s i n these temperature measurements may be due to increased r e s i s t a n c e o f the r e a c t o r thermocouple


14.

TREYBIG AND ANTHONY

Anionic

Styrene

Polymerization


5.0

Chain Length, j

Figure

8. Comparison of experimental and calculated weight fraction distributions for Run 2((%) Exp; ( ) Micro-W; ( ) Micro; ( ) Seg)

Chain Length, j

Figure

9. Comparison of experimental and calculated weight fraction distributions for Run S((%) Exp; ( ) Micro-*D; ( ) Micro; ( ) Seg)




318


TREYBIG AND ANTHONY

Anionic

Styrene

Polymerization


14.


319



320

Chain Length, j

Figure

14. Comparison of experimental and calculated mole fraction distributions for Run 12 ((%) Exp; ( ) Micro-W; ( ) Micro; ( ) Seg)


14.

TREYBIG AND ANTHONY

Anionic

Styrene

Polymerization

321


caused be bending the thermocouple sheath when handling the r e a c t o r between runs. Because of these e r r o r s the r e a c t i o n temperatures f o r these runs are not known b e t t e r than ±2°C. Reactor Runs 12, 13 and 14 were intended t o be c a r r i e d out a t 35°C, but i t was found that the temperature i n the bath f o r the r e f e r e n c e thermocouple had r i s e n from 0.6°C t o 4.5°C during the e i g h t hour reaction p e r i o d r e q u i r e d to complete the runs. The r e a c t i o n temperatures for Runs 12, 13 and 14 were c o r r e c t e d f o r t h i s o v e r s i g h t and a r e b e l i e v e d t o be accurate to ±1.0°C. The temperatures recorded f o r the remainder of the r e a c t i o n runs a r e accurate to ±0.25°C. E f f e c t of Mixing Speed on Monomer Conversion and Molecular Weight D i s t r i b u t i o n . The monomer conversion obtained i n the seeded p o l y m e r i z a t i o n of styrene i n a CFSTR w i l l be independent of the degree of segregation and thus the mixing speed as long as an exponential residence time d i s t r i b u t i o n i s maintained. Theref o r e , a dependence of monomer c o n v e r s i o n on mixing speed f o r runs with the same feed c o n c e n t r a t i o n s , average residence time, and r e a c t i o n temperature would i n d i c a t e n o n - i d e a l mixing a t the lower mixing speeds. The experimental monomer conversions obtained f o r runs a t s i m i l a r feed c o n d i t i o n s and average residence times but d i f f e r e n t mixing speeds are g i v e n i n Table IV. I n i t i a l comparison of CFSTR runs w i t h s i m i l a r feed c o n d i t i o n s i n d i c a t e s c o n d i t i o n s f o r which the monomer conversion may be dependent on mixing speed. However, when the e f f e c t s of e x p e r i mental e r r o r i n monomer c o n v e r s i o n and d i f f e r e n c e s i n r e a c t i o n temperature a r e considered, the monomer conversion i s seen t o be r e l a t i v e l y independent of mixing speed f o r rpm equal to or greater than 500. Comparing Run 14 w i t h Run 12 r e v e a l s a small decrease i n monomer c o n v e r s i o n i n s p i t e of a r i s e i n r e a c t o r temperature of 2°C. T h i s i n d i c a t e d the presence of a small amount of bypassing or dead volume a t the lower mixing speed. T h i s imperfect mixing p a t t e r n would a l s o be present i n Run 15. The experimental molecular weight d i s t r i b u t i o n s given i n F i g u r e s 8 through 14 i l l u s t r a t e l i t t l e o r no s i g n i f i c a n t e f f e c t s on the shape of the molecular weight d i s t r i b u t i o n s d i r e c t l y a t t r i butable t o the mixing speed. Thus no e f f e c t s of increased segreg a t i o n w i t h decrease i n mixing speed were observed on the molecular weight d i s t r i b u t i o n s . E v a l u a t i o n of Mixing Models. The micro-mixed r e a c t o r w i l l produce polymer d i s t t f i b u t i o n s w i t h i n c r e a s i n g amounts of h i g h molecular weight t a i l as the degree o f p o l y m e r i z a t i o n of the p o l y mer product i n c r e a s e s over that of the o r i g i n a l seed polymer. T h i s trend i s i l l u s t r a t e d by the curves f o r the micro-mixed reactor i n F i g u r e s 8 through 14. A l s o c h a r a c t e r i s t i c of the seeded, micromixed r e a c t o r i s the convergence o f the p o l y d i s p e r s i t y index to 2 for a h i g h degree of p o l y m e r i z a t i o n . T h i s trend i s i l l u s t r a t e d to some extent i n Table VI which presents the c a l c u l a t e d degrees of polymerizations.




322

The micro-mixed r e a c t o r model was not able to simulate adequately the e x p e r i m e n t a l l y observed weight average degrees of p o l y m e r i z a t i o n s or molecular weight d i s t r i b u t i o n . These f a c t s are i l l u s t r a t e d i n Table VI and F i g u r e s 8 through 14. In g e n e r a l , the weight average degrees of p o l y m e r i z a t i o n c a l c u l a t e d f o r the micromixed r e a c t o r were smaller than those observed experimentally. T h i s i s due to the more narrow polymer d i s t r i b u t i o n p r e d i c t e d by the micro-mixed model as shown i n F i g u r e s 8 through 14. The micro-mixed r e a c t o r w i t h dead-polymer model was developed to account f o r the l a r g e values of the p o l y d i s p e r s i t y index observed e x p e r i m e n t a l l y . The e f f e c t of i n c r e a s i n g the f r a c t i o n of dead-polymer i n the r e a c t o r feed while m a i n t a i n i n g the same monomer conversion i s to broaden the product polymer d i s t r i b u t i o n and t h e r e f o r e to i n c r e a s e the p o l y d i s p e r s i t y index. As i l l u s t r a t e d i n Table V, t h i s model, w i t h i t s a d j u s t a b l e parameter, £, can e x a c t l y match experiment average molecular weights and e a s i l y account f o r v a l u e s of the p o l y d i s p e r s i t y index s i g n i f i c a n t l y g r e a t e r than 2. The f a i r degree of c o n s i s t e n c y observed i n the values of j) f o r Seeds I I and I I I and the e x c e l l e n t agreement between the experimental molecular weight d i s t r i b u t i o n and those c a l c u l a t e d with (f> , lends c r e d i b i l i t y to the dead-polymer model. The agreement between experimental and c a l c u l a t e d d i s t r i b u t i o n at i n c r e a s i n g degrees of p o l y m e r i z a t i o n are given i n F i g u r e s 8 through 14. The bimodal weight f r a c t i o n d i s t r i b u t i o n s c a l c u l a t e d f o r Runs 8 and 10, which are shown i n F i g u r e s 10 and 11 are of p a r t i c u l a r interest. There i s good agreement between experiment and theory i n s p i t e of l i m i t a t i o n s i n the a b i l i t y of the GPC data r e d u c t i o n r o u t i n e to handle bimodal d i s t r i b u t i o n s . To d i f f e r e n t i a t e between the micro-mixed r e a c t o r w i t h deadpolymer and the by-pass r e a c t o r models i n t h i s i n v e s t i g a t i o n , the e f f e c t of mixing speed on the value of "(J) was observed. As i l l u s t r a t e d i n Table V, the value "" i s not observed to i n c r e a s e with decreasing mixing speed as would be expected f o r a by-pass r e a c t o r . T h i s r u l e s out the p o s s i b i l i t y of a by-pass model and f u r t h e r s u b s t a n t i a t e s the dead-polymer model. The w e l l - s t i r r e d segregated r e a c t o r w i l l produce polymer d i s t r i b u t i o n s w i t h low molecular weight t a i l s and sharp t r u n c a t i o n s a t the h i g h molecular weight ends at i n c r e a s e d degrees of p o l y m e r i z a t i o n of the polymer product. This i s i l l u s t r a t e d i n F i g u r e s 8 through 14. The v a l u e of the p o l y d i s p e r s i t y index f o r the segregated r e a c t o r product w i l l always be l e s s than that of the micro-mixed r e a c t o r (assuming no dead-polmer) as i l l u s t r a t e d i n Table VI. The segregated model was not a b l e to simulate the e x p e r i mentally observed degrees of p o l y m e r i z a t i o n on the molecular weight d i s t r i b u t i o n s . As shown i n F i g u r e s 8 through 14, the segregated d i s t r i b u t i o n s were i n general too narrow and e x h i b i t e d peaks i n the mole f r a c t i o n and weight f r a c t i o n curves which f a r exceeded those observed e x p e r i m e n t a l l y . D

g

11



14.

TREYBIG AND ANTHONY

Anionic

Styrene

Polymerization

323

S i g n f i c a n c e of Streaking i n Reaction Medium. Streaks i n the r e a c t i o n medium observed d u r i n g most continuous p o l y m e r i z a t i o n runs i n d i c a t e the presence of some degree of segregation or i n complete micro-mixing. But, as i n d i c a t e d i n F i g u r e s 8 through 14 for the comparison of experimental and c a l c u l a t e d d i s t r i b u t i o n s , no s i g n i f i c a n t i n f l u e n c e of segregation on the shape of the d i s t r i b u t i o n s was observed. In f a c t , the product d i s t r i b u t i o n i s simulated w e l l u s i n g a micro-mixed model with dead-polymer. T h i s anomaly may be explained i n p a r t by arguments due to P a t t e r s o n (33) based on s t u d i e s of a CFSTR using Monte C a r l o techniques. The e f f e c t s of micro-mixing on the molecular weight d i s t r i b u t i o n are much more pronounced than those of segregation. According to P a t t e r s o n (33) only a small i n c r e a s e i n micro-mixing over that of t o t a l segregation w i l l y i e l d a polymer d i s t r i b u t i o n very s i m i l a r to that of micro-mixed r e a c t o r . Conclusions The most s i g n i f i c a n t r e s u l t s and c o n c l u s i o n s are summarized below: 1. The monomer conversion i n t h i s seeded p o l y m e r i z a t i o n system i s independent of the degree of segregation as long as an exponential residence time d i s t r i b u t i o n funct i o n i s maintained. 2. The mixing speed had l i t t l e or no s i g n f i c a n t e f f e c t on the monomer conversions or the shape of the molecular weight d i s t r i b u t i o n s f o r mixing speeds of 500 rpm or greater. 3. A micro-mixed, seeded r e a c t o r w i l l produce a broad polymer d i s t r i b u t i o n w i t h a high molecular weight t a i l and p o l y d i s p e r s i t y index that approaches 2 at l a r g e degrees of p o l y m e r i z a t i o n . 4. The e f f e c t of dead-polymer and by-passing on the micromixed r e a c t o r f o r the same degree of monomer conversion i s to broaden the product polymer d i s t r i b u t i o n and thus allow v a l u e s of the p o l y d i s p e r s i t y index much l a r g e r than 2. 5. A w e l l - s t i r r e d segregated r e a c t o r would produce a product polymer with a low molecular weight t a i l and a sharp t r u n c a t i o n at the high molecular weight end f o r l a r g e degrees of product polymer p o l y m e r i z a t i o n . At equal monomer conversions the weight average degrees of p o l y m e r i z a t i o n w i l l be l e s s f o r a t o t a l l y segregated r e a c t o r than f o r a micro-mixed r e a c t o r . 6. The micro-mixed r e a c t o r with dead-polymer model simul a t e d the product of the l a b o r a t o r y r e a c t o r w e l l w i t h i n experimental accuracy. 7. In s p i t e of v i s u a l i n d i c a t i o n s of at l e a s t p a r t i a l segreg a t i o n , the concept of micro-mixing proved to be most u s e f u l i n modeling the l a b o r a t o r y r e a c t o r .


324

POLYMERIZATION

REACTORS AND PROCESSES

Symbols CFSTR D Dw E(t) GPC I J,J n


Imax Jmin

h M MF.

%

MWD

Pj"

P

Bj

?

RJ

R?D t

x

m a A 0 °R *0

A

Constant flow s t i r r e d tank r e a c t o r Number average degree of p o l y m e r i z a t i o n Weight average degree of p o l y m e r i z a t i o n E x i t age d i s t r i b u t i o n , dE/dt = 0 ~ e x p ( - t / 0 ) d t Gel Permeation Chromatograph I n i t i a t o r concentration Polymer c h a i n length Largest polymer chain length i n polymer d i s t r i b u t i o n Smallest polymer c h a i n l e n g t h i n seed d i s t r i b u t i o n and reactor effluent I n i t i a t i o n r a t e constant Propagation r a t e constant E q u i l i b r i u m constant f o r polymer a s s o c i a t i o n Monomer c o n c e n t r a t i o n Mole f r a c t i o n polymer of length j Number average molecular weight Weight average molecular weight Molecular weight d i s t r i b u t i o n Concentration of polymer of chain length j i n r e a c t o r feed Concentration of polymer of chain length j Concentration of polymer of chain l e n g t h J Concentration of a c t i v e polymer of chain l e n g t h j Concentration of polymer of chain length j i n the by-pass stream of a by-pass CFSTR Concentration of dead-polymer of c h a i n l e n g t h j Concentration of polymer of chain l e n g t h j i n the r e a c t o r zone of a by-pass CFSTR Concentration of t o t a l polymer Rate of i n i t i a t i o n Rate of monomer consumption Rate of propagation Residence time d i s t r i b u t i o n Time Time r e q u i r e d f o r smallest polymer molecule to grow to length j Weight f r a c t i o n of polymer of l e n g t h j Monomer conversion Lumped parameter propagation f u n c t i o n Denoting d i f f e r e n c e Average residence time Average residence time i n r e a c t i o n zone of a by-pass CFSTR Zeroth moment of polymer d i s t r i b u t i o n F i r s t moment of polymer d i s t r i b u t i o n Second moment of polymer d i s t r i b u t i o n Dimensionless time defined as [1 - e x p ( t / 0 ) ] A u t o c a t a l y t i c r a t e constant f o r i n i t i a t i o n F r a c t i o n a c t i v e polymer i n CFSTR with dead polymer F r a c t i o n by-pass i n by-pass CFSTR 1


14.

TREYBIG AND ANTHONY

< j ) R ° D

Anionic Styrene Polymerization

325

F r a c t i o n dead polymer i n CFSTR with dead-polymer F r a c t i o n passing through r e a c t i o n zone i n by-pass reactor Superscript denotes feed stream of seed polymer

Acknowledgments The authors appreciate the encouragement and support of t h i s work by the Department of Chemical Engineering, the Texas Engi neering Experiment Station, and Ε. I. duPont deNemours & Company. L i t e r a t u r e Cited


1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14. 15. 16. 17. 18. 19. 20. 21.

Worsfold, D. J . and Bywater, S., Makromol. Chem., 65, 245 (1963). Bywater, S. and Worsofld, D. J . , Can. J . Chem., 40, 1564 (1962). F e t t e r s , L. J . , Ph. D. D i s s e r t a t i o n , U n i v e r s i t y of Akron, Ohio (1962). Hsieh, H. L., J . Polymer S c i . , A3, 153 (1965). Hooke, R. and Jeeves, Τ. Α., Assoc. Comp. Cach., 8, 2, 212 229 (1961). O ' D r i s c o l l , K. F. and Tobolsky, Α. V., J. Polymer S c i . , 35, 259 (1959). Johnson, A. F. and Worsfold, D. J . , J . Polymer S c i . , A3, 449 (1965). Bywater, S. and Worsfold, D. J . , Advan. i n Chem. Ser., 52, 36 (1969). Morton, M. and F e t t e r s , L. S., J. Polymer S c i . , A2, 3311 (1964). Worsfold, D. J . and Bywater, S., Can. J . Chem., 38, 1891 (1960). Margerison, D. and Newport, J . P., Trans. Faraday Soc., 59, 1891 (1963). Hsieh, H. L. and Glaze, W. H., Rubber Chem. Technol., 43, 22 (1970). Lenz, R. W., Organic Chemistry of Synethic High Polymers Interscience, New York (1967). Tanlak, T., M. S. Thesis, Texas A&M U n i v e r s i t y , College Station, Texas (1975). Timm, D. C. and Kubicek, L. F., Chem. Eng. S c i . , 29, 2145 (1974). Danckwerts, P. V., Chem. Eng. S c i . , 2, 1 (1953). Danckwerts, P. V., Chemical Reaction Engineering, 12th Meet ing, Eur. Fed. Chem. Eng., Amsterdam (1957). Zwietering, T. N., Chem. Eng. S c i . , 11, 1 (1959). Chen, M. S. K. and Fan, L. T., Can. J . Chem. Eng., 49, 704 (1971). Keairns, D. L. and Manning, F. S., AIChE J . , 15, 660 (1969). Ng, D. Y. C. and Rippin, D. W. T., Third Eur. Symp. Chem. Reaction Eng., Amsterdam, Pergamon Press, Oxford 161 (1965).



326

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33.

Goto, S. and Matsubara, Μ., Chem. Eng. S c i . , 30, 61 (1975). Spielman, L. A. and Levenspiel, O., Chem. Eng. S c i . , 20, 247 (1965). Kattan, A. and Adler, R. J . , Chem. Eng. S c i . , 27, 1013 (1972) Nagasubramanian, K. and Graessley, W. W., Chem. Eng. S c i . , 25, 1559 (1970). Rao, D. P. and Edwards, L. L., Chem. Eng. S c i . , 28, 1179 (1973). Ahmad, Α., Ph. D. D i s s e r t a t i o n , Texas A&M University, College Station, Texas (1975). Tadmor, Z. and Biesenberger, J . Α., I&EC Fundamentals, 5, 336 (1966). Uraneck, C. Α., Burleigh, J . E. and Cleary, J . W., Anal. Chem., 40, 327 (1968). Kolthoff, J . M. and Harris, W. E., Ind. Eng. Chem. Anal. Ed., 18, 161 (1946). Lewin, S. Ζ., Chem. Educ., 43, Reprint (1966). Treybig, Μ. Ν., " E f f e c t of Mixing on Polymerization of Styrene," Master of Science Thesis, Texas A&M University, College Station, Texas (1977). Patterson, G. Κ., Personal Communication (1977).

RECEIVED January 15, 1979.


Anionic Styrene Polymerization in a Continuous Stirred-Tank Reactor

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