NORBERT PLATZER
Design of Continuous and Batch Polyethylene, PVC, and styrene resins, in particular, are discussed an-made polymers are applied as substitutes for
TABLE I. WORLD PRODUCTION IN 1968
Billion pounds Polyethylene
11.5
Poly(viny1 chloride)
10.5
4.5
Styrene resins Other thermoplastics and thermosets Total
TABLE I I .
16.0 ~
42.5
ADDITION POLYMERIZATION PROCESSES
Free Radical and Ionic Initiation Monomer in Liquid Phase 1. Bulk (Mass) Polymerization ( a ) Polymer soluble in monomer ( b ) Polymer insoluble in monomer (c) Polymer swollen bv monomer ( d ) Falling drop 2. (a, Solution Polymer soluble i n solvent ( b ) Solvent/nonsolvent Polymer insoluble in solvent 3. Siispension Initiator dissolved in monomer 4. Emulsion Initiator dissolved in dispersing medium 11. Monomer in Vapor Phase 111. Monomer in Solid Phase
I.
M metals, wood, stone, glass, paper, and natural
fibers, and have found novel uses. Polymer manufacturers offer a large variety of different grades of the same polymer, tailored either for certain applications with specific properties, such as transparency, toughness, flexibility, elasticity, nonflammability, insulation, heat, cold, or weather resistance, or for specific fabrication techniques, such as molding, extrusion, calendering, foaming, film casting, fiber spinning, or surface coating. Ten years ago the demand for each grade was small, between 1 and 40 million pounds. T o produce all these numerous grades, batch polymerization was generally preferred over continuous operation. Since then, the demand for a single commodity grade has grown to such an extent that the same polymerization reactor has to be used year-round for the production of a single grade. Today it appears desirable to produce these commodity grades on a continuous basis. Consumption has risen to 70 lb per year per head of U.S. population and to 15 lb per year of world population. Table I shows that world production passed the 40-billion-pound mark in 1968. Of these 42.5 billion pounds, 26.5 billion pounds were composed of the three polymers : polyethylene, PVC, and styrene resins. In deference to their commercial importance, most of this article is confined to the polymerization processes of these three polymers. Classification of polymerization processes
AUTHOR Norbert Platzer is Senior Polymer Specialist at Monsanto Co., 730 Worcester St., Indian Orchard, M a s s . 07057. Thts article is based on a paper given as the main lecture at the I U P A C International Symposium on Macromolecular Chemistry, Budapest, Hungary, on A u g . 26, 7969, and presented i n the German language a t Jour research laboratories of the W e s t German chemical industry. 6
I N D U S T R I A L A N D E N G I N E E R I N G CHEMISTRY
Addition polymerization processes may be classified according to the phase in which the monomer is present during polymerization : liquid, vapor, or solid phase, as shown in Table 11. In most commercial addition polymerization processes, the polymerization is started with the monomer in the liquid phase. Commercial processes are subdivided into bulk, solution, suspension, and emulsion polymerization. In bulk polymerization, we distinguish between polymers, which are soluble in their monomers, such as
Polymerization Processes
poly(viny1 acetate), polystyrene or poly(methy1 methacrylate), and those polymers, which are insoluble in their monomers, such as poly(viny1 chloride) or polyacrylonitrile. Between them lies the type of polymers, like polyethylene, which is swollen by its monomer a t specific conditions, such as elevated temperature and pressure. In the production of graft copolymers, two polymerization methods are frequently combined; for instance, bulk with suspension polymerization or emulsion with suspension polymerization. I n the production of block copolymers, anionic polymerization may be combined with free radical polymerization.
I
1
BATCH POLYMERIZATION CYCLE
Example reactor: 3700 U. S. gal-5300
lb/batch
Standard
Reduced
hours
hours
Cycle Materials charging and purging Heating Polymerization Unreacted monomer stripping Cooling Discharge Reactor flushing and cleaning Maintenance Inefficiency Total turnaround time Lb/operating hour
Reducing turnaround time of batch polymerization
Most of us who have scaled up a batch polymerization process from a laboratory flask or Coke or champagne bottle to commercial production, have faced the question: “How can we reduce the turnaround time?” A batch cycle might look like Table 111. By charging the materials hot and by discharging the polymeric batch hot into a cooling tank, we are able to cut down some of the turnaround time. Through the use of faster free-radical initiators, redox systems or ionic catalysts, we are able to reduce the polymerization time to a fraction of the overall time. As you see, we are able to double the output of a reactor easily. In the cases of suspension, emulsion, and solution polymerization, we can raise the output further by increasing the ratio of monomer to water or solvent. But, we are now faced with another problem if we want to carry out the process isothermally and not adiabatically-the removal of the heat of polymerization.
TABLE 1 1 1 .
1/2 1 8 1 1 1/2 1 1/2 1/2 14 380
1/2 0 4
1 /2 0
1/2 1,/4 1/2 1/4 6-1 /2 815
I
I
TABLE IV.
HEAT OF POLYMERIZATION
Kcal /mol, @ 25OC 21.2 19.5 17.6 16.7 22.9 18 .O 21.2 18.5 13.2 18.4 7.4
Ethylene Propylene Butadiene Styrene Vinyl chloride Vinylidene chloride Vinyl acetate Methyl acrylate Methyl methacrylate Acrylonitrile Formaldehyde
Removal of heat of polymerization
All common vinyl monomers generate about the same amount of heat during polymerization on a molar basis, as is shown in Table IV. Removal of this heat was easy in small reactors, but became a problem in larger ones. For example, the ratio of cooling jacket surface VOL. 6 2
NO.
1
JANUARY 1 9 7 0
7
1
TABLE V. JACKET SURFACE AREAS OF AGITATED POLYMERIZATION KETTLES
c
Vol, U S . gal 2,000 3,000 3,700 6,300 16,500
Cooling area, ft2 180 245 292 392 744
to volume of polymerization kettle becomes smaller with size as indicated in Table V. Large polymerization reactors require, in addition to cooling jackets, the installation of internal cooling baffles or bundles of cooling coils ( I ) (Figure 1) and frequently of reflux condensers. T h e presence of a diluent or solvent, which either boils lower than the monomer or forms a n azeotropic mixture, aids in heat removal through refluxing. Figure 2 shows the use of reflux condensers in batch (2) and continuous operations ( 3 ) . Heat removal is easier in semicontinuous operation where monomer and/or initiator is added gradually, and even easier in continuous operation. Heat removal through vaporization is also employed in the fallingdrop or spray method. Figure 3 illustrates two continuous operations employing ionic catalysts. According to a BASF patent ( 4 ) , styrene monomer is mixed with a solution of sodium naphthenate in tetrahydrofurane in a spray nozzle and ejected into a reactor. T h e heat of polymerization is removed by evaporating a portion of the monomer and the solvent. According to a Phillips Petroleum patent (~5),butadiene is sprayed with butyl lithium into the top of a tower reactor. O n the way down through the reactor, maintained a t 66°C and 330 psig, the droplets polymerize and the heat of polymerization is removed by partial vaporization. T h e evaporated butadiene is condensed a t 38°C and recycled. The polymer collected a t the bottom is discharged through a devolatilizing extruder. These processes are not yet used commercially.
Figure 1. Nonagitated reactor for batch isothermal polymerization ( I )
facture of PVC, hard spots means particles of closed surface unable to absorb plasticizer. They are particles left on the reactor’s ~ 7 a l from l the previous batch. T o avoid their formation, the kettles have to be cleaned frequently with pressure water, steam, or solvent. The possibility of hard spots formation is probably the reason why only batch suspension and batch bulk polymerization are employed for the production of dry-blend PVC resin. However, in the case of paste PVC resin, where a closed particle surface is needed, batch and continuous emulsion polymerizations have been practiced for more than 30 years ( 6 ) , and semicontinuous suspension polymerization has been practiced for 10 years (7). I n other cases, such as polystyrene or polyethylene, hard spots are either higher molecular weight or crosslinked polymer particles which will not melt as easily as the main portion of the batch. They can be formed by excessive stay of particles in the reactor, and may be reduced by frequent cleaning. They are also formed during extrusion or injection molding and may be reduced by hydrogenation of unsaturated double bonds (8)*
Residence time
One advantage of batch over continuous polymerization is that the polymerization time may be kept exactly uniform for the entire batch, resulting in a product of narrow specifications, such as molecular weight distribution, particle size, particle surface, composition, or configuration. The expression “hard spots” is familiar to all concerned with quality control. In the manu8
I N D U S T R I A L A N D E N G I N E E R I N G CHEMISTRY
Figure 2. Reflux condensers f o r the removal of heat of polymerization
Figure 3. Falling drop reactors
I n continuous processes, the polymerization time depends on the residence time. I t may be assumed that a single agitated kettle would be employed for a continuous suspension or emulsion polymerization and that “perfect” mixing takes place. The probability that a suspended droplet or particle remains in the kettle can be expressed by the forthcoming Equation 1
fl(.Z)
= Ze-z
(1)
where 2 = Qt/V [flow rate (gal/min) X time (min)]/ volume of kettle (gal). This equation (at the ratio of V’Q = 1) is plotted as Graph A in Figure 4. If we assume that the average residence time ( V / Q ) would be, for example, 4 hr, it can be calculated that 10% of the fed monomer would remain in the reactor less than 25 min, 63y0 for less than 4 hr, and 0.791, longer than 21 hr. For a stream of suspended droplets of particles through a cascade of perfectly agitated kettles, the probability that a droplet or particle remains in the cascade between t and t dt units of time is
+
(2) This equation is plotted as Graph B in Figure 4 for two equal volume reactors in cascade. Integration shows that only O.5y0 of the fed monomer is in the system less than 25 min and only 26,4y0 of the monomer is less than 4 hr. Equation 2 applies only for cascading kettles where no back-mixing occurs (q = 0). Backmixing is inevitable when the continuous polymerization takes place in a multistage reactor where the different sections are only separated by baffles. If a particle
Figure 4. Pulse response of perfectly agitated kettles and plug-Jow reactor VOL. 6 2
NO.
1
JANUARY 1970
9
backtracks once during its passage, it will, in effect, 2 stages, before it finally leaves the pass through n reactor. Likewise, if it backtracks twice, it will, in 4 stages during its passage, etc. effect, traverse n Therefore, the probability densities for these effected 2, n 4,n 6, etc., stages are: stages of tz
+
+ + + + Zn+l . . (n + l ) ! ( n + 3 ) ! ,-Z
Zn+3
,-Z
Z n + 5 . ,-2
(n
+ 5)!
(3)
O590
100
110
100
110
+
120 130 TEMPERATURE "C
140
150
140
150
For backmixing, Z becomes equal to (Q 2q) t/V instead of Qt/V (Equation 1) (9). Little backmixing occurs in plug-flow tubular reactors, for which the fgllowing equation for pipeline mixing may be applied :
I t can be seen that the degree of pipeline mixing is characterized by the parameter :
40
10
INDUSTRIAL AND E N G I N E E R I N G CHEMISTRY
130
PARTICLE SIZE GROWTH IN -0RlDE EMULSION
QL
Graph C of Figure 4 shows the pulse response of a plug-flow tubular reactor which simulates closest the uniformity of batch operations. Several polymer properties are dependent upon the difference in residence time. Figure 5 illustrates two examples: the effect of residence time on particle size and on molecular weight distribution. I n the presence of a water-soluble initiator, polymer particles grow during emulsion polymerization. Batch polymerization of vinyl chloride results either in a single particle size, or, when seeded, in two sizes-large and small-or only large (70). Continuous operation, on the other hand, results in a broad particle size distribution, dependent on the difference of residence time. Styrene is generally polymerized in a n adiabatic rather than isothermic procedure. Polymerization rate increases and molecular weight decreases with temperature. Dependent upon the residence time at the various temperature zones, the molecular weight distribution can be broadened or narrowed according to the reactor design. Both continuous cascading ( I I ) and plug-flow reactors (72) were designed to prevent backmixing in continuous suspension polymerization processes. Although not employed in suspension polymerization, cascading kettles are widely used in the manufacture of rubber latices by emulsion polymerization, and plug-flow tubular reactors in the manufacture of low-density polyethylene. Low-density polyethylene (0.91 5-0.935 g/cm3) is made in a high-pressure tubular reactor process (73), as illustrated in the flow diagram of Figure 6. To provide plug flow, the reactor tubes are characterized by a high length-to-diameter ratio which varies between
120
TEMPERATURE "C
08 06
04
02 0
Figure 5.
Figure 6.
0 IO I2 POLYMERIZATION TIME HOURS
2
4
6
8
Time-dependent polyme? characteristics
Tubular reactor for high-pressure polymerization of ethylene
250/1. to 12,000/1. Although comparable to bulk polymerization of a liquid monomer, the initiation step takes place in the vapor phase, since the polymerization temperature of 100-200" C is above the critical temperature of ethylene (9.9"C). I n the high-pressure process, purified ethylene is compressed to 15,000-50,000 psig and mixed with a free-radical initiator. The first section of reactor tubes has to be heated, because the linear velocity is very high (residence time as low as 20 sec), and the generated heat of polymerization is insufficient. Only the second section of tubes is cooled to provide the desired temperature profile. The reactor can be either a long single tube, a tube with multiple feed streams along its length, or a bundle of tubes, mounted vertically or horizontally in high-pressure cells or bays and equipped with relief valves, rupture disks, and retractable cleaning rods. The high velocity causes a pressure drop along the tube length and, with it, a change in propagation rate, which is pressure dependent. This results in a broader molecular weight distribution than by employing a stirred autoclave reactor. Operating tubular reactors a t pressures above 30,000 psig and temperatures above 160°C allows oxygen initiation. Lower operating temperatures are possible if free radical initiators in combination with chain-transfer agents are employed. Preferred initiators are peroxides with a 1-min half-life a t polymerization temperature (74). Chain transfer agents of various types, such as ketones, aldehydes, alcohols, hydrogen, or chlorinated compounds, are added to provide a narrower molecular weight distribution. Conversion varies below 2070, dependent on initiation and reactor design. T h e polymeric mixture is discharged into a separator maintained a t 200-250°C and reduced pressure of 1000-7500 psig (15). Lowering of the pressure reduces the solubility of ethylene in polyethylene and results in a separation of the two. A second pressure reduction releases more ethylene gas. T h e polymer melt is extruded into strands or ribbons, while the unreacted ethylene gas is purified from oils and waxes and returned to the compressors. T h e largest high-pressure polyethylene unit will be started up by Northern Petrochemical Co. a t Joliet, Ill., in 1971. I t will have a capacity of 450 million pounds per year and will use the BASF process. I n some instances, the same reactor design, such as cascading kettles or plug-flow tubular reactors, may be used for different monomers or different types of polymerization processes. I n most cases, however, the reactor and process have to be modified, according to the product. Various reactors and auxiliary equipment, designed for specific types of polymerization processes, are discussed in the following chapters. T h e application of the three principles: fast removal of heat of polymerization, uniform mixing, and trends to control residence time may be recognized.
Bulk polymerization of polymers soluble in their monomers
A tremendous rise in melt viscosity is experienced during bulk (or mass) polymerization of polymers, soluble in their monomers. I n the case of low-molecular-weight poly(viny1 acetate), one is able to complete polymerization in a n agitated kettle. However, when one polymerizes styrene in bulk in a n agitated kettle one stops a t 30-4070 conversion and transfers the viscous sirup to another type reactor. Already in the 1930's, 0. Roehm (76) completed the styrene bulk polymerization in a plate-and-frame press. As shown in Figure 7, the heat of polymerization is removed by circulating water through the plates. T h e polymer blocks, formed inside the frames, are cooled, removed, crushed, and granulated. Nine hundred million pounds per year of polystyrene are still made by this plate-and-frame press process. T h e capacity of one standard polypress is 10 to 13 million pounds per year. Also during the 1930's, Wulff, Dorrer, and Ohlinger (17) designed the first continuous bulk polymerization process for styrene a t Ludwigshafen. They used a n adiabatic tower as illustrated in Figure 8. After batch prepolymerization in one of the two agitated kettles, the sirup was fed continuously at the top of the tower. Its temperatures were controlled by cooling coils and jackets. With 100-140°C a t the top and 160-2OO0C a t the bottom, thermal polymerization could be completed to a conversion of 92--98y0 within 3-10 hr. The polymer melt was discharged a t the bottom through a short-barrel extruder furnished with a strand-die. T h e
Figure 7. Batch bulk polymerization of styrene (16) VOL. 6 2
NO.
1
JANUARY 1970
11
extruded strands were cooled and pelletized. The continuous tower process was also adopted for vinyl acetate polymerization. I n later years, the polystyrene process was modified (78)-e.g., monomeric styrene was removed as vapor a t the top of the tower, condensed, and recycled to the poll-merization process. Today, Ludwigshafen uses a continuous, isothermic, thermally initiated bulk polymerization process. I n 1953, Dow Chemical Co. filed a patent on a n agitated tower process (79). T o avoid channeling inside the towers and for a better heat transfer, DOWchanged to three towers in series with slow agitators and grids of cooling or heating pipes (Figure 9). Each tower is provided with three temperature zones in which the temperature is automatically controlled by circulating Dowtherm (a mixture of diphenyl and diphenyl oxide, bp 257OC). The polymeric melt is allowed to rise from 95" to 225OC. In order to reduce viscosity and to ease heat transfer, a solvent, such as ethyl benzene, may be added. Unreacted monomer and solvent are stripped off in a stationary vacuum devolatilizer (20). Into this, the reaction mixture is fed through a multiorifice die to produce continuous strands. This allows rapid diffusion and vaporization of monomer and solvent, which are recycled. The polymeric melt is extruded, cooled, and pelletized. Besides stationary vacuum devolatilizer and vented single or twin screw extruders, wiped film devolatilizers may be employed. There are numerous designs of wiped
Figure 8. Continuous bulk polymerization of styrene i n a singly agitated thermally controlled adiabatic tower (77)
Figure 9. Continuous bulk polymerization 12
INDUSTRIAL A N D ENGINEERING CHEMISTRY
of
styrene in a series of agitated towers (78-20)
Figure 10. Removal of unreacted monomer on vacuum two-roll mixing mills with smooth (27) and threaded rolls (22)
main-polymerizer, followed by a stationary devolatilizer (20). I n later years, several other companies patented the design of vertical and horizontal shear screw reactors (23) Linear polybutadiene or other elastomers may be dissolved in the styrene monomer before polymerization, resulting in a n impact graft copolymer. I n the continuous tower process (79) and in the Union Carbide process (22), a solvent, such as ethyl benzene, is generally added. One and four fifths billion pounds per year of general purpose and impact polystyrene are made by the tower process, and 250 million pounds per year are made by the Union Carbide process. T h e capacities of modern units are as high as 130 million pounds per year. T h e continuous bulk processes may also be used to produce random copolymers or to graft both monomers upon a n elastomer. For the copolymerization of monomers of different reactivities, such as styrene and acrylonitrile, Dow Chemical Co. recommends a continuous recycle copolymerization process (24,as illustrated in Figure 11. I t is said that a copolymer of the same composition is obtained as the monomers feed mixture because the system automatically adjusts the monomer mixture composition in the reaction zone to reach a steady state. BASF is employing continuous bulk polymerization, not only for its PSAN, but also for its acrylonitrile-butylacrylate-styrene graft copolymer. Bulk polymerization of polymers insoluble in their monomers
film devolatilizers, where the monomer and solvent are removed under vacuum while a serrated plow impeller or screw scrapes the walls from the polymeric melt and gives a forward motion to the melt. To produce a uniform high-molecular-weight polystyrene a t low-polymerization temperature, and to avoid the making of low-molecular-weight portions in the hot section of the adiabatic polymerization tower, H. Ohlinger suggested, in 1939, to carry out the bulk polymerization a t 8OoC to only 36-40%. According to his suggestion, the polymeric sirup from the prepolymerization kettles (see Figure 8) a t Ludwigshafen were fed to a vacuum two-roll mixing mill (27), and the nonreacted monomer was evaporated and recycled. I n contrast to this process, which was abandoned for economical reasons, Union Carbide Corp. polymerizes styrene in vacuum agitated kettles with reflux condensers and carries out the polymerization to almost completion. I t uses a vacuum two-roll mixing mill with threaded rolls (22) which combine the functions of devolatilization and conveying. Both vacuum mills with smooth and threaded rolls are illustrated in Figure 10. I t may be mentioned that, already in 1948, Dow Chemical Co. filed a patent on the use of a n agitated kettle as prepolymerizer, and a single-screw extruder as
I n contrast to polystyrene, poly(viny1 chloride) is insoluble in its monomer and precipitates from the monomer in bulk polymerization. T h e advantage of bulk polymerization over standard suspension polymerization is higher purity. T h e product is free of foreign materials, such as dispersing agents or salts, which would influence moisture absorption, trans-
Figure ? 1. Continuous recycle copolymerization process (24) VOL. 6 2
NO. 1 J A N U A R Y 1 9 7 0
13
Solvent-nonsolvent polymerization
Figure 12. Bulk polymerization of vinyl chloride one-step PechineySt. Gobain process (25)
parency, and electrical properties unfavorably. In the original one-step Pechiney-St. Gobain process which was practiced in France and Germany, a 3000-gal ball mill was used as reactor (25), as illustrated in Figure 12. T h e balls densified the precipitated fluffy PVC to a fine powder, similar to general-purpose grade, suitable for extrusion or calendering. I n 1960, PSG patented a continuous ball mill process (26) with a periodically opening discharge valve. This process did not become cominercial. Further development led PSG to its twostep bulk polymerization process (27). One line consists of one 2000-gal fast agitated prepolymerization kettle and three or four 4200-5300-gal batch ribbon blender-type autoclaves, as illustrated in Figure 13. Usually, half of the vinyl chloride is fed to the prepolymerizer, furnished with a reflux condenser, and polymerized to 7y0conversion by the addition of a fast initiator (28) in less than 1 hr. The speed of agitation determines the size of the polymer particles, which act as seed in the second step of polymerization. T h e second half of vinyl chloride is fed to the ribbon blender-type autoclave, and polymerization is completed in 5-9 hr. T h e reaction is simply stopped by evaporating the unreacted monomer, which is condensed and reused in the following batch. After applying a vacuum, the resin is discharged under agitation by pneumatic conveyors to screens for removing oversized particles. One line has a capacity of 35-45 million pounds per year dryblend resin. During the last three years, PSG has sold its two-stage bulk polymerization process to 12 companies in eight countries. The process is especially suitable for PVC of lower molecular weight because it can be operated twice as fast as suspension polymerization, resulting in a lower production cost and unit capital cost. PSG has also patented (29) a vertical reactor with two agitators-a slow turning top agitator and a faster turning bottom agitator, plus a reflux condenser. This reactor will be built as large as 10,000 gal and might replace the horizontal PSG reactors in the future. 14
INDUSTRIAL A N D ENGINEERING CHEMISTRY
Solvent-nonsolvent polymerization is one of the older methods for polymerizing vinyl chloride. The diluent is miscible (soluble) with the liquid monomer, but is a nonsolvent for the polymer, Hydrocarbons, alcohols, ethers, or esters are such diluents. During the polymerization, the polymer precipitates and can be separated by filtration or centrifuging. The advantage of this method over bulk polymerization is that conversion can be carried out almost to completion. Its disadvantage is the recovery of the solvent and the fact that diluents frequently act as chain transfer agent. Already in 1929, Imperial Chemical Industries obtained a patent (30) on the continuous solvent-nonsolvent polymerization of vinyl chloride. Union Carbide’s vinyl chloride/vinyl acetate copolymer (VYHH) is also made by a continuous solvent-nonsolvent process (37). Solution polymerization
Solution polymerization is frequently employed in copolymerization, where the copolymer formed is soluble in the solvent, and the polymeric solution is used directly in surface coating or as adhesive. Solution and solvent-nonsolvent polymerizations are employed in the low-pressure polymerization of olefins with Ziegler-Natta or Phillips catalysts. I n the solution process, the solvent may contain cycloparaffin (32).
Figure 13. Bulk polymerization of vinyl chloride two-step PechineySt. Gobain process (27)
Figure 74. Low-pressure polyolejn processes
Both monomer and polymer remain in solution during the reaction while the catalyst is maintained in suspension by agitation. I n the solvent-nonsolvent process, a n aliphatic paraffin is used as liquid dispersant in which the monomer is soluble, and the catalyst and the polymer are suspended. Reaction temperature may be as high as 125" to 175°C and a t 300-400 psig pressures in the solution process. But the reaction temperature has to be held below 110°C in the solvent-nonsolvent process to prevent dissolution of the polymer. The latter process has the advantage over the solution process that the viscosity is lower, because the polymer is precipitated as suspension. A flow diagram of solution and of solvent-nonsolvent polymerization to linear polyethylene (specific gravity, 0.94-0.97), polypropylene, and their copolymers is shown in Figure 14. T h e Phillips catalyst is a complex of chromium oxide with silica-alumina (32). The Ziegler-Natta catalyst is a complex of titanium chloride with aluminum alkyl, and is used in naphtha dispersion a t 60-80°C and slight pressure with an efficiency up to 1 : 100,000. The Ziegler-Natta catalyst has the advantage over the Phillips catalyst in that it can be operated a t lower temperature without dissolving the polymer. Lower molecular weights are obtainable through the addition of hydrogen or other chain transfer agents. T h e polymerization reactors are jacketed agitated kettles, up to 8500 gal, and are operated mostly con-
tinuously and sometimes batchwise. They represent only a small portion of the intricate plant, which consists of five sections : (a) the catalyst preparation and activation section; (b) the polymerization section with a monomer recycle system; (c) the catalyst deactivation and removal section; (d) the polymer separation, drying, and compounding section; and finally (e) the solvent, diluent, and precipitant fractionating and recovery section. T h e polymer is frequently compounded with antioxidants and colorants in twin-screw extruders. I n the solution process, the spent catalyst is filtered off from the polymeric solution. I n the solvent-nonsolvent process, the deactivated catalyst is generally dissolved and extracted from the polymeric slurry. If the amount of catalyst used is very small, it is allowed to remain as oxide in the polymer, eliminating the extraction step. This polymer is employed in all applications of linear polyolefins, with the exception of an electrical insulator. Although the low-pressure process requires a n extensive plant, it has the advantage over the high-pressure process that its conversion rate is significantly higher, requiring less monomer recycle. Further attempts have been made toward capital and production cost reductions. I n the novel solventless process, propylene is charged as a liquid and the polymerization is carried out with a Ziegler-Natta catalyst, activated with a minute quantity of water a t room temperature yielding a VOL. 6 2
NO.
1
JANUARY
1970
15
Figure 75. Suspension and bulk-suspension polymerization
product of 85y0 tacticity. The catalyst efficiency is three to six times higher than in the processes employing solvents. The catalyst is killed with superheated steam and again allowed to remain as oxide in the polymer. This process eliminates not only the step of catalyst removal but also that of solvent recovery. For the production of low-density polyethylene, a high pressure of 15,000 to 50,000 psig is employed. At its Houston plant, Phillips Petroleum Co. started a new low-density polyethylene unit operating only at 500-600 psig and using the same chrome oxide catalyst as for its high density polyethylene. Stereoregulated PVC and polystyrene are also made in solution or solvent-nonsolvent process with either a Ziegler-Natta catalyst or a boron fluoride cationic catalyst a t -20" to 0°C. Suspension polymerization
Suspension polymerization differs from solvent-nonsolvent polymerization insofar that both the monomer and the polymer are insoluble in the diluent, which is generally water. The monomer droplets are suspended with a suspending agent, such as poly(viny1 alcohol), calcium phosphate, colloidal clay, hydroxy ethyl cellulose, or acrylates in the water, and the initiator is dissolved in the monomer. Most dry blend and general-purpose PVC is produced by batch suspension polymerization. Developed in the thirties by B. F. Goodrich Co. and Wacker-Chemie, process and equipment designs have changed little 16
INDUSTRIAL A N D E N G I N E E R I N G CHEMISTRY
(Figure 15). Agitated, jacketed kettle, glass-lined, or stainless-steel clad is used as reactor, sometimes with a reflux condenser. Reactor size varies between 2C00 and 16,000 gal. Letdown tank, continuous centrifuge, and rotary dryer are the standard auxiliary equipment. Faster initiators, large reactors, and more automation are the major innovations. Acetyl cyclohexyl sulfonyl peroxide, diisopropyl peroxydicarbonate, tertiary butyl peroxy pivalate, and azo-bis-isobut>-ronitrile are faster initiators, generally used in combination with lauroyl peroxide. Recently, Goodyear (33) disclosed and licensed an initiator that is I t is diethyl peroxydicarbonate, formed in sztu. formed by adding ethyl chloroforinate with the monomer, and sodium bicarbonate and hydrogen peroxide in the water phase. A polymerization rate of 4-1 '2 hr and a cost savings of 0.2-0.5 C/kg of PVC are claimed. Cost reduction in the processing step is just as important as in the polymerization step. This is generally accomplished by combining flash with rotary drying. Another version, taken by Allied Chemical Corp. in its Painesville plant, is the combination of continuous flash dryers with batch-operated fluidized bed dryers ( 3 4 ) . Chemische Werke Huels installed a digital cornputer system (35) (Foxboro PCP-88 dual computer) for the automatic control of 10 mixing kettles and 80 batch autoclaves. It claims that performance is much better with direct digital control than with analog control systems. The computer is able to charge over 100 different formulations.
Batch suspension polymerization is also used in the production of beads or pearls of poly(viny1 acetate), poly(acry1ic esters), polystyrene, expandable polystyrene, and PSAN. Several designs for continuous suspension polymerization reactors have been patented (36). T h e reactors are either agitated kettles with extended bottom sections for settling, or towers with multiple blade agitators or compartments. The principle of the latter is the same as of cascading kettles to avoid backmixing. A semicontinuous suspension polymerization process (7) is employed by Wacker-Chemie and Tenneco for the production of PVC paste resin. Vinyl chloride, containing lauroyl peroxide, is homogenized to very small droplets in water with emulsifier by passing through a colloid mill. T h e homogenized monomer is charged a t the top of a simple flooded tower. The tower, 2.5 X 65 ft, is jacketed, stainless-steel clad, and without agitation. The flow of the suspension is downward. At 50"C, conversion is brought to 90%. T h e PVC slurry is discharged a t the bottom, separated, and washed on a continuous pressure filter, followed by drying and grinding of agglomerated particles. Mixed with plasticizer, the viscosity of the plastisol is more sensitive than that of plastisol from emulsion polymer, requiring a slower casting speed. Suspension polymerization is frequently combined with bulk polymerization, such as in the manufacture of impact polystyrene and ABS. Polybutadiene or another elastomer is dissolved in styrene monomer. I n a n agitated batch kettle, styrene or styrene plus acrylonitrile is grafted upon 6-1 2y0 rubber backbone by a bulk prepolymerization process. After 3o-35Y0 conversion, the sirup is suspended in water and the polymerization is completed in a larger agitated batch kettle by suspension polymerization. Kettles u p to 10,000 gal are in use. Tertiary butyl perbenzoate, benzoyl peroxide, and other styrene-soluble initiators are used in combination with lauryl mercaptan or terpinolene as chain-transfer agent. The beads obtained are centrifuged and dried. I n 1965, Cosden obtained a patent (37) on a n intermittent bulk/suspension polymerization process for impact polystyrene and ABS. Thermal prepolymerization was carried out a t 100°C for 18 hr to a conversion of 30%. Every 6 hr, one third of the polymeric sirup was transferred to the suspension kettle. The polymeric sirup was mixed with an equal quantity of water which contained initiator and suspending agents. The polymerization was completed a t 100°C within 6 hr, yielding a product of uniform bead size. In a n older patent to Distillers (38), a continuous bulk/suspension polymerization for styrene is described. T h e monomer was mixed with benzoyl peroxide and prepolymerized in a tubular reactor a t 80°C to a conversion of 68-70%. T h e highly viscous sirup left the reactor through a face pelletizer which was attached to the bottom of a continuous flooded suspension tower maintained a t 9OoC. The cut pellets were suspended in
water and passed upward. The time of passage was adjusted so as to ensure complete polymerization in 3 hr. Emulsion polymerization
The difference between aqueous emulsion and suspension polymerization lies in the type of initiator used, not particle size. The initiator for emulsion polymerization is water soluble, whereas the initiator for suspension polymerization is monomer soluble. I n contrast to suspension polymerization, where the particle size is fixed by the size of the monomer droplet which contains the initiator, the particle in emulsion polymerization grows from a micelle or seed (39). T h e initiator is either a persulfate, a hydroperoxide, or a redox system (do), as in the manufacture of cold rubber. Protective colloids, surfactants, and buffers aid in the stability of the emulsions. Advantages of emulsion polymerization are: (a) effective temperature control during polymerization due to rapid heat transfer; (b) minimum tendency toward coalescence of polymer particles ; (c) the preparation of high solids content emulsions with low viscosity ; (d) the possibility of controlling particle size; and (e) good control of many copolymerizations which are difficult to control in bulk or suspension methods. Continuous, batch, and delayed monomer addition processes are commercially employed. Emulsion polymerization has found its largest application in the manufacture of synthetic rubber, such as polybutadiene, GRS, and other copolymer rubbers. Cascading kettles are used in their continuous polymerization processes. Poly(viny1 acetate) (47) and polyacrylate lattices for surface coating or as adhesives are more frequently made in single kettles, either in a direct batch operation or by adding monomer and initiator in portions. High-pressure emulsion polymerization of ethylene has been known for almost 30 years (42). A flow diagram of a more recent continuous process of Spencer (43) is illustrated in Figure 16. T h e polymerization is carried out a t 85°C and 3000 psig with potassium persulfate as initiator in a n autoclave a t vigorous agitation. T h e latex produced is concentrated from 22 to 40% in a vacuum wiped-film evaporator and screened. Distillate from the concentrator and ethylene from the defoam tank are recycled to the reaction process. Ethylene may also be copolymerized with vinyl chloride in aqueous emulsion a t pressures u p to 50,000 psig (44) with potassium persulfate and a redox system as initiators. Already in the forties, I. G. Farbenindustrie produced a polystyrene latex. Polymerization was carried out a t 70°C for 2 hr and a 90°C for 1-1/4 hr with potassium persulfate as initiator. The latex was either coagulated by electrolytes or dried on continuous vacuum mill rolls. At the same time, emulsion copolymerization of styrene with acrylonitrile, methacrylonitrile, and acrylates (45)was taken up. Twenty-five years ago, T. A. T e Grotenhuis observed that a rubber of superior flexibility and abrasion resisVOL. 6 2
NO.
1
JANUARY 1970
17
Figure 76. Continuous emulsion polymerizution of ethylene (44)
tance could be obtained from a heterogeneous polyblend of a hard unrnasticated rubber as disperse phase and an elastic rubber as continuous phase (46). I t was recognized that a heterogeneous blend was also necessary for making tough and flexible plastics. T o obtain an impact polystyrene, butadiene and styrene were polymerized in emulsion separately and the two latices were blended together (47). T o obtain a n ABS copolymer blend, a nitrile-rubber latex (48)and a PSAN latex were prepared separately and blended together, as patented by Uniroyal (49)and shown in Figure 17a. In 1952, the term “graft copolymer of styrene upon polybutadiene” was defined by IUPAC and the principle of grafting became known (50). Since then, numerous patents (57) have been disclosed on the graft copolymerization to produce impact polystyrene, ABS, similar ter- and tetrapolymers, as well as rigid PVC. Figure 17b shows one of the current processes for ABS graft polyhlends (52). Styrene and acrylonitrile are grafted upon a polybutadiene latex and blended with a SAN copolynier latex in the desired proportion. An antioxidant is added to the mixed latices which are coagulated by the addition of an electrolyte. Control of temperature, agitation, and dilution results in the formation of a crumb of desired size and density. Experiments have also been made to coagulate the latices by freezing (53) or electrodeposition. The crumb is washed, filtered, dried, and compounded either in Banbury mixers or twin screw extruders. I n the cornpounding step, other polymers made by bulk or suspension polymerization may be blended in. A great variety of multicomponent polymer systems is possible by combining the blending of latices with the compounding of their crumb. PVC resins, made by emulsion polymerization, are employed as diluent resin for rigid and plasticized applications where transparency is not required. Its main application is as paste resin. The polymer particles are extremely fine and relatively nonabsorptive toward plasticizers with which it forms a paste a t room temperature. Another application is as vinyl chloride copolymer latex with unsaturated comonomers for surface coating. In 1935, 1. G. Farbenindustrie started the first produc18
I N D U S T R I A L A N D E N G I N E E R I N G CHEMISTRY
tion employing a batch process. Polymerization was carried out in a rotating glass-lined 8000-gal autoclave followed by a coagulating kettle, centrifuge, and vacuum dryer. Paste resins of single or binodel particle size are manufactured in batch operation using standard agitated kettles with reflux condensers. Paste resin of broader particle size distribution (0.3-1.0 p ) is manufactured by a continuous process, as shown in Figure 18, practiced since 1937. The polymerization reactor is a tall, glasslined, jacketed kettle (generally 3500 gal) with a H/D ratio of 3-4: 1 (6). I t is furnished with a short paddle agitator in the top section but without a reflux condenser. There, mortomer and water, containing emulsifier and initiator, are fed a t the top continuously. T h e agitation is adjusted so that effective emulsification of the monomer in water takes place in the top section and sufficient motion and saturation are maintained at the bottom. T h e conversion is not completed and 10% of the monomer charged leaves the reactor with the latex. Frequently, the outflow of two reactors is combined and led to a specially designed degassing kettle, kept a t reduced pressure. The top section contains a spiral chute on which the latex flows down and degasses. The degassed latex is collected in the bottom section. T h e recovered monomer is scrubbed and returned for reuse. The PVC is dried in a spray dryer (up to 11,000 lb/hr), and the agglomerated particles are ground and screened. This continuous process is less expensive than the conventional batch emulsion polymerization process. However, the product is less heat-stable due to the difference in residence time in the continuous reactor. Various designs of continuous emulsion polymerization reactors have been patented. In an earlier patent (54,Shell described continuous emulsion polymerization of vinyl chloride in a compartmented reactor, whereas its later patents claim only undivided vertical reactors (55). Two agitator designs are patented by I.C.I. (56) to create inside a tower reactor a pattern of so-called “Taylor rings’’ or segmentation of the liquid into horizontal planes. Vapor phase polymerization
I n contrast to the commercial bulk polymerization processes, vapor phase polymerization has the disadvantage that the heat of polymerization cannot be removed by the latent heat of vaporization. There are few low-pressure vapor phase polymerization processes for olefin (57) known. Most patents state that diluents or solvents are desirable. Figure 19 illustrates a spray reactor (58)in which the olefin enters as gas and the Ziegler-Natta catalyst dispersed in a liquid diluent is sprayed in a t the top. Lupolen 5261-2, a high-density polyethylene, and Novolen, a polypropylene of BASF, are made by a vapor phase polymerization process. Last May, a 53-million-pound-per-year Novolen plant got into operation a t Wesseling. A Phillips catalyst is used because it is less sensitive to traces of impurities, such as acetylene or aldehydes. I t is charged without
Fagure 17. Batch polymerization processes for the manufacture of ABS ( a ) Blend of two random copolymers: A
+B
( b ) Blend of graft polymer with random copolymer: B f D
Figure 78. Continuous emulsion polymerization of vinyl chloride VOL. 6 2
NO.
1
J A N U A R Y 1970
19
Literature Cited (1) Dow Chemical Co., U.S. Patent 2,494,924 (1950).
Figure 19. Vapor phase polymerization of olejin gas (58)
solvent with purified olefin gas into a fluidized bed tower reactor. For polyethylene the reactor temperature is maintained at 100-110” C and the ethylene pressure at 500 psig. The heat of polymerization is removed by recycle expansion-compression cooling (59). The polymer is removed as a free-flowing powder. The catalyst is allowed to remain in the polymer, as its efficiency is very high (>10,000 parts of polymer/l part of catalyst). A novel approach of vapor phase polymerizing vinyl chloride has been patented by Solvay (60). A peroxide initiator is absorbed on porous PVC and placed in an autoclave. Vinyl chloride gas is fed into the reactor under agitation at 58-59OC and 114-psig pressure. The pressure is maintained constant during polymerization and more vinyl chloride is added. After 5-10 hr, the vinyl chloride addition is stopped, and a powdery PVC is obtained. Solid phase polymerization
Polymerization of monomers in the crystalline state is still in the research stage. I t is initiated either by high energy radiation (67) or ultraviolet light. Little is known about the nature and site of initiation or of the rate of crystal structure (62). Radiation induced polymerization of monomer crystals, such as hexamethyl cyclotrisiloxane, has been studied in view of their ease of polymerization and high vapor pressure (63). Photopolymerization of crystals of 2,5-distyrylpyrazine and of 1,4-bis(/3-pyridyl(2)vinyl)benzene was investigated by the Japanese Textile Institute (64). These new territories of polymerization will lead to completely new designs of process equipment. Concluding remarks
I hope that this survey leaves you with the impression that polymerization process technology has outgrown the state of empirical scaling-up from laboratory glassware and has reached the state of designing and building commercial plants which operate economically in the production of uniform polymers of improved properties. If one of you should have to select the best process, his choice will be guided by many factors-for the most part by the final application of the polymer. 20
INDUSTRIAL AND ENGINEERING CHEMISTRY
(2) Farbwerke Hoechst A.-G., British Patent 723,991 (1955). (3) Shell Oil Co., U.S. Patent 3,349,070 (1967). (4) BASF, German Patent 1,139,975 (1965). (5) Phillips Petroleum Co., U.S. Patent 3,182,050 (1965). (6) I. G. Farbenindustrie A.G., U.S. Patent 2,363,951 (1944); Jacobi, B., Angew. Chcm., 64, 539 (1952). (7) Wacker-Chemie, U.S. Patent 2,981,722 (1961). (8) Witt, D. R., Hogan, 3 . P., A.C.S. Meeting, Spring 1969. 4, 88 (1965). (9) Retallik, W. B., I N D .ENC.CAEM., FUNDAM., (10) B. F Goodrich Co U.S. Patent 2 520 959 (1950)’ Koch H Sommer W., Kunststke, 47, 153 (195 (i946). (43) Spencer Chem. Co., Belgian Patent 621,386 (1963). (44) Monsanto Co., U. S. Patent 3,403,137 (1968). (45) I. G. Farbenindustrie A,-G., U.S. Patent 1,933,052 (1933), 2,102,179 (1937), 2, 140,048 (1938). (46) T e Grotenhuis, T. A,, U.S. Patent 2,457,097 (1948). (47) Dow Chemical Co., U.S. Patent 2,460,300 (1949); Goodrich Co., U.S. Patent 2,614,089 (1952); General Tire & Rubber Co., U.S. Patent 2,745,818 (1956). (48) I. G. Farbenindustrie, U.S. Patent 2,140,048 (1936). (49) Uniroyal, U.S. Patent 2,439,202 (1948), 2,498,652 (1950), 2,505,549 (1950), 2,698,313 (1953). (50) IUPAC, J . Poiym. Sci., 8,260. U.S. Patent 2,754,282 (19561, 2,948,703 (1960); Fire(51) Union Carbide Cor stone, U.S. Patent 2,7fi1270 (l956), 2,802,808 (1957) ; Goodyear, U S . Patent 2,835,645 (1958); Monsanto, U.S. Patent 2,863,849 (1958); B.P. 778,267; U.S.P. 2,550,139 (1951). (52) Uniro a1 U.S. Patent 2,820,773 (1958); Borg-Warner, U.S. Patent 3,010,936 (1961); nion Carbide Corp., U.S. Patent 3,168,593 (1965). (53) Brit. Geon Ltd., Belgian Patent 654,722 (1965). (54) Shell Dev. Co., U.S. Patent 2,475,016 (1949). (55) Shell Dev. Co., U.S. Patent 2,618,626 (1952), 2,813,850 (1957). (56) I.C.I., U.S. Patent 3,003,98G (1961). (57) Ruhrchemie A.G U S. Patent 2,924,591 (1960); Wisseroth, K., Agnew. Makromol. Ckem., 8 , 21 (1369). (58) Dow Chemical Co., U.S. Patent 2,906,742 (1959). (59) B.4SF, G.P. 1,013,870 (1956). (60) Solvay et Cie, Belgian Patent 686,088; Netherlands Patent 6,711,784 (1967). (Gl) Schmitz, J. W., Lawton, F. J.,Srieno, 113, 718 (1951); Fox, D., eta[., “Physics Chemistry of the OrganicSolid State,” Inlerrcience, Ed. 1, 4 (1963). ( 6 2 ) Okamura, S., J. Polyrn. Sci., 58, 925 (1962); C 4, 827 (1964). ( 6 3 ) Chawla, A. S., St. Pierre, E., Ad”. Chem. Ser., 91 (1969). (64) Hasegawa, M., Jup. Chem. Quart., 5-1,45 (1969).
.
I _
&
‘
I
