E N G I N E E R I N G STUDY OF C O N T I N U O U S P O L Y M E R I Z A T I O N OF A C R Y L I C M O N O M E R S J . F. T E R E N Z I AND H. F. C O S W A Y ' Plastics Division, American Cyanamid
co., Stamford,
Conn. 06904
Continuous free radical bulk and solution polymerization of acrylic monomers, methyl methacrylate in particular, i s described. Consideration i s given to appropriate kinetic models, and the experimental data from a pilot plant scale, two-stage continuous reactor system are presented. Agreement between the experimental and calculated results is reasonably good. The study is further amplified by a simple analog computer simulation of the kinetic network superimposed on the conservation equations. Trial and error computer adjustments of the kinetic constants provided much clearer insight into the more complicated kinetic systems.
PRACTICAL continuous bulk polymerization of acrylic monomers, particularly to form copolymers of methyl methacrylate and ethyl acrylate, has long been sought. Certain patents (Bild and Jukes, 1966; Gorham and Brandon, 1962; Terenzi and Schmitt, 1966) and literature (Swedlow, 1968) relate to continuous processes for producing cast sheet or optical quality materials as well as molding and extrusion compounds. The advantages of continuous polymerization have been pointed out (Denbigh, 1947; Jenkins, 1960). The continuous or capacity flow method, from a fundamental standpoint, enables one to obtain velocity or over-all constants from the steady state. Furthermore, if reasonable amounts of data are available and the system is programmed on an analog computer, the constants can be easily matched in both the transient and steady states by trial and error. The feasibility of employing bulk or concentrated solutions depends upon the heat of polymerization of the monomers in question as well as the entire kinetic regime. I n an engineering sense, one must also consider the prospects of reaching economic conversion levels, the reliability or reproducibility of the process under practical conditions, and the feasibility of mixing the viscous solutions t o effect adequate heat transfer (Schildknecht, 1956). For the free-radical bulk polymerization of methyl methacrylate one normally expects common first-order kinetics up to about 15 to 20y0 conversion, followed by a rapid autoaccelerating increase in reaction rate nith a simultaneous increase in molecular weight. This rate increase is attributed to the Trommsdorf or gel effect or reduced termination step. Other contributors have shown (Fujui et al., 1956; North and Reed, 1961) that the initial decrease in rate, followed by onset of the autocatalytic effect after a certain conversion, can be explained in terms of a two-stage diffusive termination process. In any event it is possible to adjust over-all kinetics empirically, if the effect is not so sensitive as to cause poor control and gross nonreproducibility. The magnitude of this effect is clearly demonstrated in the suspension polymerization of methyl methacrylate, where a considerable temperature rise occurs even though full cooling is applied a t the onset. This paper describes a preliminary study of a two-stage continuous partial conversion polymerization of methyl methacrylate. Only kinetic effects are considered along with analog computer simulation of transients. Present address, Polaroid Corp., Cambridge, Mass.
Preliminary Considerations
For bulk polymerization of methyl methacrylate, in the presence of a chain regulator such as dodecyl mercaptan, the kinetic curves can show reasonably slow reaction rates and isothermal polymerization can prevail up to a t least 50% conversion. The acceleration effect is also considerably depressed. A typical conversion curve is shown in Figure 1. h'eglecting chain transfer to monomer and using literature values of kinetic constants (Tobolsky and Xesrobian, 1954) and the typical kinetic expressions for homogeneous free radical polymerization:
it appears that approximately normal bulk kinetics prevail a t low conversions, and even beyond 20% conversion when mercaptan is present. In order t o apply such systems to continuous reactors one must consider the problems of heat removal and adequate mixing of the viscous media a t higher conversions and determine appropriate conservation equations. A steady-state heat balance around one step can be written as:
Taking values of U , AT (Colgan, 1960)) AH,, C, (Riddle, 1954), R, (Perry, 1967), and X as 5.0, 20, 230, 0.5) 30. and 155, respectively, Equation 4 reduces to
Heat mechanically introduced through agitation can be significant for extremely viscous solutions, but its effect is not considered here. Varying the parameters in Equation 5 to simulate practical conditions, one can calculate allowable conversions of 25 to 50% per hour. The large beneficial effect of cold monomer feed does not change with reactor geometry. For practicality-Le., in removing and treating unreacted monomers-it would be preferable to obtain a t least 6001, VOL.
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0.266% butyl mercaptan, Temperature. 100' C.
0.141 % lerf-butyl perbenzoate
I. 1% H z 0 II. 0.031% H z 0
were placed in a thermostated oil bath a t the desired temperature, taken out after suitable intervals, and put into a Dewar flask containing dry ice and acetone to arrest the reaction. The contents of the tubes were analyzed by measuring the refractive index or by reacting the excess monomer with mercaptan and titrating excess mercaptan iodometrically. Some runs were conducted in a recording dilatometer. Agreement among all of the methods was satisfactory. For simulation of practical operations with industrial raw materials and equipment, trace amounts of various suspects such as water were added separately and the kinetic deviation was determined. Reactors. A sketch of the continuous reactors is shown in Figure 2. The first-stage reactor was a 5-gallon stainless steel jacketed vessel containing two baffles with a dished bottom and flat top. The agitator was a four-bladed fantype impeller The reactor was also equipped with a temperature recorder, nitrogen purge lines, reflux condenser, and vent line. Controls were available for use of water, high pressure steam, and atmospheric steam in the jacket of the reactor. The second-stage screw reactor was a 3-gallon vessel with a vertical screw with a ribbon attached to the periphery of the screw and pitched in the opposite direction. There was provision for a vapor spare above the flight with two portholes to observe level and boiling. The vessel was designed for 150 p.s.i.g. Pumps and Feed Tanks. RIonomer feed tanks were simple cylindrical stainless steel vessels and appropriate pumps were provided. Continuous Operation. The premixed monomer, initiator, and transfer agent were continuously fed from the nitrogenpurged feed tanks into the first stage. The system could be started with monomer in the vessels or actually from any condition of conversion. Temperature was usually maintained at 100' C. and the vessels were maintained a t constant level by the overflow in stage 1 and by manipulation of the effluent valve in stage 2. To maintain observable good mixing, the speeds of the first and second stage were generally 100 and 50 r.p.ni., respectively. Polymer concentration was not allowed to exceed about 40% in the first and 60% in the second. Sormally the system was thoroughly cleaned out with solvent between runs.
TIME. (HOURS)
Figure 4.
Effect of oxygen on polymerization rate
0.0097% dicumyl peroxide and 0,6% dodecyl mercaptan Temperature. 100" C.
Results and Discussion
To apply the polymerization batch kinetics properly to the continuous reactors it was necessary t o simulate practical conditions as closely as possible-for example, Figure 3 shows the effect of small amounts of water in the monomer. Figure 4 shows the large effect of the presence of oxygen. Oxygen dissolved in the monomer acts as a polymerization inhibitor a t room temperature, whereas a t 110' C. and in the presence of monomer it causes a considerable increase in reaction rate. For most of the laboratory curves, the polymerization was generally first-order with respect to monomer. For continuous reactor simulation, however, for some work a straight-line approximation could be used along with the first model-i.e., Equation 6. This was true of cases in which the initiator had a long half life compared t o the reactor residence times and particularly between 20 and 60% conversion. Probably diffusional effects as well as imperfect mixing are interfering with the kinetics. Assuming perfect mixing and negligible volume change, the analytical solutions of Equations 6 and VOL.
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Figure 5.
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TIME, (HOURS)
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Continuous single-stage poly(methy1 methacrylate) polymerization
Temperature. 100’ C. Catalyst. 0.14% ferf-butyl perbenzoate up to 32 hours, at which point 0.0108% dicumyl peroxide was substituted 10)
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EXPERIEEWTAL ‘2nd STAGE EXPERIMENTAL 1st STAGE CILCULATED CALCULATED
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Figure 6. Continuous two-stage poly(methy1 methacrylate) polymerization Temperature. 100’ C. 0.14% fed-butyl perbenzoate and 0.60% dodecyl mercaptan 191 = 0% = 55 minutes
7 are straightforward and are given in Equations 11 and 12 in terms of polymer concentration.
of a runaway reaction, the run was terminated shortly after
4 hours. After examining the interior of the second-stage reaction, we found small portions of reaction products which were considerably higher in conversion than the experimental points show. Slower rates of 5 to 6% per hour, shown in the batch curve in Figure 7 , were used as a basis for the continuous run shown in Figure 8. Rates of 5 to 6% per hour shown check closely with the batch kinetics. The run was terminated after it was considered to have reached the steady state. Figure 9 shows a two-stage polymerization with 2,5-dimethyl2,5-di- (tert-butylperoxy )hexane as the initiator in the presence of xylene, a reasonably inert solvent. In this case the initial transients were exceedingly fast and the models do not fit. Analog Computer Simulation
The equipment used was the Pace TR-10 with readout directly plotted. In some of the more complicated situations, two computers were tied together. The equations t o be solved, including both first-order monomer dependencies and first-order catalyst dissipation kinetics, under isothermal conditions are: - P1F r1Ti71= TT‘ldPl/dt (131
+ P1F - P2F + ~2T1’2 = TT‘zdPp/dt
CFF -
C1F
CIF - C2F
- kl’C1Ii-1 = Ti’ldCl/dt
(15)
- k2’C2Tt2
(16)
= Tl‘zdCs/dt
- PI)(Ci)’” rz = Kz(l - P2) (Cz)l’z TI
Figure 5 shows experimentation with only the first single stage. The model follows the experimental data closely. Toward the end of the run shown in Figure 5, the feed stream was switched to monomer containing dicumyl peroxide rather than tert-butyl perbenzoate. Here a calculated first-order model would probably have fitted the experimental data better. Figures 6 and 7 show the two-stage operation. The operation shown in Figure 6 was conducted a t conversion rates of 237, per hour, indicating feasibility according to earlier predictions but not necessarily practicability. In Figure 6, the second-state experimental points become erratic a t the 50 to 60% conversion level. Because of the possibility 202
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FUNDAMENTALS
(14)
= Ki(1
(17) (18)
Residence times 81 and 82 have been defined as TT’I/F and W2/F, respectively. For convenience, temperatures were maintained as close to 100’ C. as possible in the runs described here. From other experimental data, both continuous and batch, the temperature effect of the over-all polymerization process could be described as:
K
= 0.7 exp [12,570 (1/373
- 1/’K)]
(19)
The data shown in Figure 9-i.e., the rapid transientswere subsequently reproduced a number of times. It was first thought that excess initiator charge was used, that the reaction dependences were of a much higher order, or that other more subtle, unknown reactions took place. This latter
Figure 7.
Conditions for slower polymerization rate
0.0108% dicumyl peroxide, 0.266% butyl mercaptan and 0.051% HzO Temperature.
100' C.
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8. Continuous two-stage poly(methy1 methacrylate) polymerization Temperature. Y100' C. 0.1 08% dicumyl peroxide, 0.60% dodecyl mercaptan 01 = 5.0 hours 02 = 5.5 hours
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trations would be approximately 1270 in the firsb stage and 28% in the second stage. The initial transients are not greatly affected by the half life of the initiator employed when half lives are long, but for less than 10 hours the effect becomes very pronounced. In Figure 11, data from one run are plotted on the computer traces. Use of a 3-hour half life causes actual and simulated curves to correlate well. To elucidate the true kinetic order, Equations 13 to 18 were solved for both zero-order and first-order kinetics. To simplify the system and partially eliminate the effect of initiator, the time interval for transient analyses is not taken
0.38
from the beginning of the run-i.e., from actual time, t = 0since the effect of initial catalyst concentration is considerably damped after 3 or 4 residence times. The step function change in the feed rate which produces the transient operations is not large enough to change the equilibrium catalyst level significantly in either vessel. The data from the run in Figure 9 are shown in Figure 12 together with the computer solutions starting a t 40 hours, a t which point the residence times in stages 1 and 2 were changed from 3.40 to 4.17 hours and from 4.50 to 5.60 hours, respectively. Both stages were sampled every 30 minutes and the samples analyzed by refractometer. The first-stage transient response is sufficiently large and correlates reasonably well with the firstorder rate of reaction. To elucidate the initiator decomposition kinetics further, batch tube polymerizations were performed and an attempt was made to analyze residual catalyst with a special vapor phase chromatography instrument. Initial results with a reported & l o % accuracy are shown in Figure 13. The decomposition curve is one that might be expected for an initiator with two peroxide groups and extrapolation of initial and final slopes indicates that half lives on the order of 1/2 hour and 4 hours might be expected, together with a descriptive expression such as: C1 = C F [ A exp (-2.236)
f B exp
Figure 12. Analog computer traces superimposed on experimental data from Figure 9 commencing at 40 hours 204
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FUNDAMENTALS
(-0.17t)l
(20)
Nomenclature
A‘
= constant composed of individual velocity
A,, A,
= surface area
constants
Figure 13.
Catalyst decomposition kinetics at 100” C.
Solvent system similar to feed stream in run described in Figure
9
Subsequent experimentation and simulation of the initiator decomposition kinetics (Repetti, 1967), using initial concentration values of about 0.05’%, did not verify the analytical accuracy given above, but averaged about &20%. The analytical problem is associated with the very small catalyst concentrations being used. For a catalyst having two peroxide residuals, one can assume a consecutive, two-step decomposition such as:
where A1,2,3,4 refer to the particular species and q 2 , to the velocity constants. These result in the following rate expressions : dAJdt = --1Al dAz/dt = alA1 - azAz
dAa/dt = up41
+a 4 2
(22 1
rlAd/dt = a2A2
Solution of these equations for the rate of initiation and substitution into the usual free radical scheme enables a more detailed simulation for the coiitinuous reactor. This work was done (Repetti, 1967) and essentially verifies the results herein reported, particularly when a2 is much greater than al. This appears to be the case despite the poor experimental accuracy of the data reported in Figure 13. Acknowledgment
The authors are indebted to the Research Service Department, Central Research Division, and to L. 0. Oldsberg, J. C. Barone, and F. H. H. Geurtsen for experimental work on the laboratory polymerizations and pilot plant equipment operation. They also thank R. V. Repetti for his help in preparing the manuscript.
of reactor walls, etc., surface area of condensor, respectively, sq. ft. CP = specific heat of reactor ingredients, B.t.u./lb. O F . Ct = chain transfer constant D, = degree of polymerkation Dno = degree of polymerization without chain transfer x = latent heat of vaporization, B.t.u./lb. F = weight throughput, lb./hr. f (t) = residence time distribution AHP = heat of polymerization, B.t.u./lb. klJ k 2 , k3 = specific polymerization rate constants for initiation, propagation and termination steps = rate constants for catalyst dissipation kl‘, k2I K = over-all polymerization rate constant [ M ] , [C], [PI = concentration of monomer, catalyst and dead polymer (subscripts F , 1, 2 refer to feed, 1st and 2nd stage concn.) M (t) = batch kinetic function for monomer disappearance PO,B = (in Equations 11 and 12). Initial polymer concentration and B = 91- 92/91 RP = rate of polymerization, % conversion/hr. ‘1, r2 = individual reaction rates in first and second stage CSHI = concentration of chain transfer agent R, = vaporization rate, lb./sq. ft. hr. AT = over-all temperature difference between reactants and jacket T?J Tj = temperature, O F . , of reactor ingredients and reactant feed temp. t = time 412 = catalyst half-life, hr. 01, e2 = first- and second-stage residence times (Wi/F, Tv2/F), hr. U = over-all heat transfer coefficient, B.t.u./hr. sq. ft. O F . W = reactor holdup, lb. literature Cited
Bild, F., Jukes, A. W. (to Imperial Chemical Industries), U. S. Patent 3,234,303 (Feb. 8, 1966). Colgan, J. D., American Cyanamid Co., Stamford, Conn., unpublished communications, 1960. Denbigh, K. G., Trans. Faraday SOC.43, 648 (1947). Fujui, S., et al., J . Polymer Sci. 20, No. 96, 584 (1956). Gorham, W. F., Brandon, D. F. (to Union Carbide Corp.), U. S. Patent 3,026,307 (March 20, 1962). Jenkins, A. D., Polymer 1, 79-89 (1960). North, A. M., Reed, G. A., Trans. Faraday SOC.67, 859 (1961). Perry, J. H., ed., “Chemical Engineer’s Handbook,” 2nd ed., McGraw-Hill, New York, 1967. Repetti K. V., American Cyanamid Co., Stamford, Conn., un ublished communications, 1967. Riddre, E. H., “Monomeric Acrylic Esters,” Reinhold, New York, 1954. Schildknecht, C. E., “Polymer Processes,” Interscience, New York. 1956. I
Swedlow Inc., Prospectus Common Stock, C. E. Unterberg, Towbin Co., Oct. 3, 1968. Tradmor, J. A., Biesenberger, J. A., IND.ENG.CHEM.FUNDAMENTALS 6. 236 (1966). Tadmor, J. A., Biesenberger, J. A., Polymer Eng. Sci. 8, No. 4, 299 (1966b). Terensi, J. F., Princeton University, unpublished lectures, 1962196.1.
Teienii, J. F., Schmitt, J. M. (to American Cyanamid Co.), U. S. Patent 3,262,960 (May 24, 1966). Tobolsks, A. V., Mesrobian. R. B.., “Organic Peroxides.” Interscience; New Pork, 1954. ’ Zeman, R., Amundson, N. R., A.I.Ch.E. J. 9, No. B, 297 (1963). RECEIVED for review November 19, 1968 ACCEPTEDMarch 3, 1969 VOL.
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