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THE COPOLYMERIZATION OF METHACRYLATES IN A COUNTER-ROTATING TWIN-SCREW EXTRUDERJongbloed, H.A.; Mulder, R.K.S.; Janssen, L.P.B.M.

Published in:Polymer Engineering and Science

DOI:10.1002/pen.760350705

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Publication date:1995

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Citation for published version (APA):Jongbloed, H. A., Mulder, R. K. S., & Janssen, L. P. B. M. (1995). THE COPOLYMERIZATION OFMETHACRYLATES IN A COUNTER-ROTATING TWIN-SCREW EXTRUDER. Polymer Engineering andScience, 35(7), 587 - 597. DOI: 10.1002/pen.760350705

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The Copolymerization of Methacrylates in aCounter-Rotating Twin-Screw Extruder

H. A. JONGBLOED, R. K. S. MULDER, and L. P. B. M. JANSSEN

Department of Chemical EngineeringUniversity of Groningen

9747 AG Groningen, The Netherlands

The copolymerization of n-butylmethacrylate with 2-hydroxypropylmethacrylatewas studied in a closely intermeshing counterrotating twin-screw extruder. Theaverage molecular weight of the product can be increased by increasing the screwrotation rate or the die resistance or by decreasing the throughput or the barreltemperature. The conversion can be improved by decreasing the throughput,increasing the die resistance, and (within limits) increasing the barrel tempera-ture, as well as through post-initiation. Compared with various classical polymer-ization processes, this situation requires that particular attention be paid to theoccurrence of a gel effect, the existence of a thermodynamic ceiling temperature,and the reactivity ratio of the monomers used.

INTRODUCTION

Traditionally, the extruder was primarily used as aplasticating apparatus and in food technology.

Recentlythere has been growing interest in the use ofthe extruder as a reactor for synthesizing or modifY-ing polymers. The main advantage of the extruder asa polymerization reactor with regard to the conven-tional polymerization reactors is the ability of theextruder to work without solvents. In order to achievegood mixing and heat transfer in conventional reac-tors, the increasing viscosity during the polymeriza-tion reaction necessitates the use of solvents. In ex-truders however, the mechanical action of the screwsprovides mixing and heat transfer even at high vis-cosities.

In recent years, various investigations have beenreported in which reactions were performed success-fully in twin-screw extruders (1-3), indicating thatreactive extrusion is a rapidly developing field. How-ever, most of these processes concerned homopoly-merizations from monomers or prepolymers or modi-fication and grafting reactions. In the present paper,the bulk copolymerization of n-butylmethacrylatewith 2-hydroxypropylmethacrylate in a closely inter-meshing counterrotating twin-screw extruder is ex-amined.

THECOUNTERROTATINGTWIN-SCREW EXTRUDER

The counterrotating twin-screw extruder used inthis work is closely intermeshing. This implies thatthe channel of one screw is blocked by the flight ofthe other screw. The extruder can therefore be con-sidered as a series of C-shaped chambers (Fig. 1), in

which material is transported towards the die. Themaximum volumetric displacement rate or theoreti-cal throughput Qth equals the number of C-shapedchambers transported per unit time multiplied by thechamber volume:

Qth = 2mNVc (1)

where m is the number of thread starts, N is thescrew rotation rate, and Vc is the volume of a cham-ber.

Because of mechanical clearances, however, thechambers are not completely closed. Leakage flowscause interaction between the chambers. The realvolumetric throughput Q is therefore given by:

Q= Qth - Ql (2)

in which Ql is the sum of all leakage flows over across section of the extruder. Four kinds of leakagegaps can be distinguished (Fig. 2) (4):

• The flight gap is a clearance between the barrel andthe flight of the screws.

• The tetrahedron gap originates from a gap betweenthe flight walls. It has approximately the shape of atetrahedron. This is the only gap that connectschambers on the opposite screws.

• The calender gap is formed by the clearance be-tween the flight of one screw and the bottom of thechannel of the other screw, and resembles a calen-der;

• The side gap is a clearance between the flight wallsof the screws, perpendicular to the plane throughthe screw axes.

POLYMER ENGINEERING AND SCIENCE, MID-APRIL 1995, Vol. 35, No.7 587

H. A. Jongbloed, R. K. S. Mulder, and L. P. B. M. Janssen

Fig. 1. C-shaped chamber.

The leakage flows are driven by inter-chamber pres-sure differences D Pc and conveying of the movingsurfaces:

DPcQl=AN+B- (3)

h

where A and B are geometrical constants and h isthe viscosity.

If a liquid feed, such as a monomer or a prepoly-mer, is fed to the extruder, two zones can be distin-guished in the extruder (Fig. 3):

• a partially filled zone, where the chambers are notyet completely filled with material;

Fig. 2. Leakage gaps in a closely intermeshing cOunterrotat-ing twin-screw extruder.

• the pump zone or fully filled zone, where pressureis built up to equal the die pressure and where thechambers are fully filled.

In the partially filled zone, there is no axial pres-sure gradient. In this zone the monomer or prepoly-mer is heated to the temperature at which polymer-ization starts. The starting point of the reaction istherefore determined by the temperature of the bar-rel, the throughput, and the screw rotation rate. De-pending upon the actual value of these parameters,the reaction could start in the partially filled or in thefully filled zone.

When a polymerization is performed in an extruder,there are five externally adjustable parameters thatcan influence the reactive extrusion process. Theseparameters are the screw rotation rate, the through-put, the barrel temperature, the die resistance, andthe initiator radical concentration. A reactive extru-sion interaction diagram has been presented by Gan-zeveld and Janssen (5), from which the influence ofthe adjustable parameters on other reactive extru-sion features can be derived (Fig. 4). Although this

Fig. 3. Schematic representation oj the extruder.

588 POLYMER ENGINEERINGAND SCIENCE, MID-APRIL 1995, Vol. 35, No.7

The Copolymerization oj Methacrylates in a Counter-Rotating Twin-Screw Extruder

diagram has originally been derived for homopolymer-izations, it is also helpful for analyzing the machineperformance and the general tendencies of the reac-tion process for copolymerizations in extruders.

THE REACTION KINETICS

The bulk copolymerization of n-butylmethacryl-ate (BMA)with 2-hydroxypropylmethacrylate (HPMA),which was carried out in an extruder to investigatethe reactive extrusion process, is a free radical addi-tion polymerization. This implies that the polymerchains are formed in a very short time, after whichthey are excluded from further reaction. Every singlepolymer chain is initiated and terminated at differentpoints in time, which means that every chain has itsown reaction time.

The kinetic scheme for polymerization is:

decomposition initiator:kd

I-->2R'

kinitiation monomer: R·+ M ?i Mi

kppropagation: Mi'+M --> Mi·+1

termination:k,c

Mi'+ Mj --> Mi+j (combination)

Fig. 4. Reactive extrusion interaction diagram: + positiveinfluence; - negative influence.

k'dMi'+ Mj· --> Mi + Mj (disproportionation)

where kd, ki, kp' ktc, and ktd are the correspondingrate constants in every reaction step.

The theoretical polymerization rate up can be de-rived from this scheme and equals:

Up = kp V .fkd [I]05[ M] (4)kt

where [I] and [M] are the initiator and monomerconcentration, respectively.

In a concentrated system, however, (as in the bulkpolymerization in an extruder) deviations from thisequation occur. This is caused by the strong increasein viscosity during polymerization. The polymer chainradicals cannot diffuse easily through the highly vis-cous medium, so the termination rate decreases. Be-cause the diffusion of the small monomer moleculesis not restricted until very high conversions, if thepolymerization is carried out above the glass transi-tion temperature (6,7), the propagation rate does notdecrease, resulting in an increase of the polymeriza-tion rate. This phenomenon is known as the Tromms-dorff effect or gel effect (8). The gel effect results in ahigher molecular weight and higher conversion of thereaction than would be expected from conventionalkinetics.

In case of the polymerization of methylmethacry-late (MMA),a number of models describing the diffu-sion limitation of the chain radicals have been pre-sented in the literature (7,9, 10), but very often thesemodels contain parameters unknown for the copoly-merization of BMA with HPMA or for the individualhomopolymerizations. However, Malavasic et at (11)have shown for several acrylates and methacrylates,including MMAand BMA,that the intensity of the geleffect decreases and the conversion at the onset ofthe gel effect increases with increasing length of thealkyl sidegroup. These effects can be explained byconsidering the process that leads to the terminationreaction between two macroradicals. Among others,Benson et at (12) and Dionisio et at (13) describe thetermination reaction as a process involving bothtranslational diffusion of the polymer chains and seg-mental motion of those chains, leading to chemicalreaction. The segmental motion improves with in-creasing length of the alkyl group as a consequenceof the shielding of the carbonyl by the alkyl group,and with that the gel effect diminishes. At the sametime, the glass transition temperature Tg of the corre-sponding polymer decreases because the overall flexi-bility of the polymer chain influences Tg. For thehomopolymers of BMA and HPMA, the values of Tgare 20°C and 76°C, respectively (14). Therefore HPMAis expected to have a larger influence on the gel effectthan BMAduring the copolymerization.

In case of the copolymerization of two monomers A(Le. BMA)and B (Le. HPMA)the propagation step ofthe polymerization consists of four different reac-



tions:k11

~A'+A-> ~AA·

kl2

~A·+B-> ~AB'

k21~B'+A-> ~BA·

k22~B'+B-> ~BB'

where each reaction has its own propagation rateconstant,

If one assumes that the radical concentrations areconstant, the well-known copolymerization equationcan be derived from the mass balances over everycomponent, This equation reads:

d[A] [A] r1[A] + [B]d[B] = [B] r2[B] + [A] (5)

in which r1 and r2 are the reactivity ratios defined asusual:

r1 = k11lk12

r2 = k221k21 (6)

The values of r I and r2 for the copolymerization ofBMAwith HPMAcan be calculated using the Q- ande-values according to Alfrey and Price (14, 15), If BMAis taken as monomer A, this yields r I = 0.29 andr2 = 0.28, The ratio d[ A]I d[ B] signifies the ratio ofincorporation of monomers A and B, which is notnecessarily equal to the monomer feed ratio, Conse-quently, one of the monomers is consumed morerapidly than the other, and the composition of themonomer feed varies, leading to a continuous shift ofthe composition of the copolymer formed, The courseof the ratio of incorporation is a function of the mono-mer feed ratio. For the copolymerization of BMAwithHPMA,this is illustrated in Fig. 5, in which:

[A]fa= [A] + [B]

d[ A] (7)Fa= d[A] + d[B]

As can be seen in the Figure, this particular copoly-merization shows an azeotropic point, in which theincorporation sequence of both monomers does equalthe monomer composition, so there is no preferentialreaction of one of the monomers. This implies that noshift in compositional distribution of the copolymerwill occur if the monomer feed ratio is set to thisazeotrope. For the copolymerization of BMA withHPMA,this feed ratio should be 50.3 mol% BMA to49.7 mol% HPMA.Because a deviation of the mono-mer feed ratio from this azeotrope will intensify itselfduring the polymerization, weighing the monomersmust be carried out with accuracy,

Most studies in the literature on the kinetics ofpolymerization reactions have been limited to lowdegrees of conversion and low reaction temperatures.

Fig, 5. Relation between composition Fa of the copolymerformed and the monomer feed fa.

At high temperatures, however, another complicatingfactor arises. Dainton and Ivin (16) have establishedthat for certain systems a depropagation step has tobe taken into account, leading to an equilibrium ofpropagation and depropagation:

kpMi·+ M ~ Mi'+I (8)

kdp

in which kdp is the depropagation rate constant. Thisimplies that above a certain temperature no net poly-merization occurs. This so-called ceiling temperatureTc satisfies the following equation:

T. = DHpc DSp+Rln[M] (9)

Thus, the higher the conversion, the lower themonomer concentration and consequently the lowerthe ceiling temperature.

For pure MMA,the ceiling temperature has a valueof 220°C (17), but, for the polymerizations of BMAand HPMA,no data are known concerning the ther-modynamic equilibrium. Elias (18), however, statesthat the standard molar entropy for the polymeriza-tion of liquid monomer to condensed amorphous pol-ymer is practically independent of constitution in thecase of compounds with olefinic double bonds, Differ-ences in Tc are therefore caused by the polymeriza-tion enthalpy alone, The polymerization enthalpy ofthe homopolymerizations of BMAand HPMAas givenin the literature (14) are 57.5 10Imol for BMA at atemperature of 74,5°C and 50.5 kJlmol for HPMAat25°C. If the influence of temperature is neglected, anenthalpy of 54.0 kJ Imol can be expected for thecopolymerization with azeotropic monomer feed ratio,

590 POLYMER ENGINEERINGAND SCIENCE, MID-APRIL 1995, Vol. 35, NO.7


If this value is compared with the polymerizationenthalpy of MMA,which is 55.5 10Imol at 74.5°C, itcan be concluded that the occurrence of a ceilingtemperature under the working conditions used inreactive extrusion is probable.

Up to now, there have been no kinetic investiga-tions concerning the copolymerization of BMA withHPMAat high temperatures and conversions. Be-cause the temperature of the extruder barrel rangesfrom 115 to 145°C, resulting in an even higher prod-uct temperature due to the heat generated by thereaction, and conversions up to 92% have been ob-tained, additional experiments were necessary to in-vestigate the course of the copolymerization at hightemperatures. The polymerization of HPMA and thecopolymerization of BMAwith HPMA(both inhibitedand with a combination of two peroxide initiatorswith a total initiator concentration of 0.07 mol/I)were studied using isothermal differential scanningcalorimetry (DSC) (11, 19) at temperatures of 110,120, 130, and 140°C. Although this method gives agood picture of the course of the reaction qualita-tively, the variance in the conversions obtained isabout 5%. To simulate the situation in the extruder,the inhibitor was not removed from the monomer.The results will be compared with the results of mea-surements with BMA(5).

To be able to interpret the results adequately andto validate the possible occurrence of a ceiling tem-perature, the polymerization enthalpy of the copoly-merization of BMAwith HPMAwas determined usingDSCand compared with the values for the homopoly-merizations of BMAand HPMAgiven in literature.

RESULTS

The polymerization enthalpy of the copolymeriza-tion obtained with DSC equals A Hp = 60.0 ± 0.5 101mol.This value shows a significant discrepancy fromthe expected value of 54.0 10Imol. This differencecan be explained by the fact that the polymerizationenthalpy of HPMA,as reported in the literature, wasdetermined during emulsion polymerization, in whichsolvent effects can occur.

The results of the isothermal experiments areshown in Figs. 6 to 8. Looking at the conversion-timedata at low polymerization temperatures (100 and110°C), one can see a clear transition from con-ventional kinetics to the gel effect stage. At highertemperatures, a curved profile can be detected, prac-tically indicating an overall enhancement of the poly-merization. This curved profile has also been ob-served by Dube et at (20) in the polymerization ofbutylacrylate (BA).

The gel effect gets more pronounced in the se-quence BMA, BMAIHPMA, HPMA, and the onset ofthe gel effect occurs at a lower conversion in thisorder: 70%, 60% and 35% conversion, respectively.This is due to the higher viscosity of the copolymerand the HPMApolymer with regard to the BMApoly-mer.

In case of the copolymerization, a deviation from

Fig. 6. Conversion curve oj the homopolymerization oj n-BMA(from Ref 5): • T= ll0°C; • T= 120°C; D T= 130°C.

Fig. 7. Conversion curve oj the homopolymerization oj 2-HPMA: • T= 100oC; • T= 110°C; • T= 120°C; ... T=130°C.

the azeotropic monomer feed ratio would cause aconsecutive polymerization of the fast reacting HPMAand the slower BMA.This could result in a similarpicture of the reaction course as depicted in Fig. 8.However, in our case, a consecutive polymerization isvery unlikely because the slope of the conversion-timeplot of the copolymerization at the end of the reactionis much steeper than in the case of the homopolymer-ization of BMA, whereas at high temperatures (120and 130°C)no separate reaction regimes can be dis-tinguished. Both observations corroborate the expec-tation of azeotropic polymerization.

THE COPOLYMERIZATION OF BMA WITHHPMA IN AN EXTRUDER

In order to investigate the influence of the ad-justable reaction parameters and extrusion parame-



Fig. 8. Conversion curve oj the copolymerization oj n-BMAwith 2-HPMA: • T= 100°C; _ T= 110°C; • T= 120°C; A

T= 130°C.

ters on the product and therewith on the process, thecopolymerization of BMA with HPMA was carriedout in a closely intermeshing counterrotating twin-screw extruder. The rotation rate of the screws, thethroughput, the temperature, and the die resistancewere varied, the latter having a distinct influence onthe fully filled length and therewith on the averageresidence time in the extruder.

Experimental Setup

The two monomers BMA and HPMA (both inhib-ited) were mixed in a ratio of 50.3 mol% BMA and49.7 mol% HPMA, corresponding to the monomerfeed ratio in the azeotropic point. Instead of a singleinitiator, a combination of two peroxide initiators withdifferent half-life values was added to the monomersto prevent a deficiency of radicals towards the end ofthe reactive extrusion process, where the tempera-ture has increased significantly. The total initiatorconcentration was 0.07 mol/I. The mixture of mono-mers and initiators was continuously flushed withnitrogen to prevent further inhibition of the reactionby oxygen and was fed to the extruder at room tem-perature.

The reactive extrusion took place in a 40 mm Rolle-paal closely intermeshing counterrotating twin-screwextruder (L/D= 15). The extruder has five heatingzones, which can be used to create a temperatureprofile over the extruder. The barrel wall of the ex-truder was kept at a uniform temperature in the lastfour zones ranging from 115°Cto 145°Cin the differ-ent experiments. The first zone (closest to the feedport) was always kept 10°Cbelow the temperature ofthe other zones. The output of the extruder rangedfrom 1.2 to 3.6 kg/h. The die resistance could bevaried by changing the size of the die outlet. Thepressure profile over the extruder was measured withthree pressure transducers. The die pressure varied

between 0.1 and 27.5 bar. Samples were collectedwhen the extruder reached a steady state after everyadjustment of the reaction and extrusion parameters.In order to stop the reaction immediately, the sam-ples were directly frozen in liquid nitrogen at theoutlet of the extruder.

Analysis

The conversion of the polymerization reaction wasdetermined gravimetrically. About 1.5 g of the poly-mer-monomer mixture was weighed precisely anddissolved in acetone. A small amount of hydro-quinone was added to this solution to prevent furtherpolymerization on heating. The solution was pouredinto a petri dish. Acetone and monomer were evapo-rated in a vacuum stove at 150°C and 300 mbar. Ifone weighs the remaining polymer, the conversioncan be calculated:

weight of polymerconversion = x 100%

weight of mixture

The disadvantage to this method is that monomercould be prevented from evaporation by the formationof a polymer film. Therefore the film was kept verythin and reproducibility experiments were performed.These experiments showed a variance in conversionof 2.5%.

The average molecular weight of the reactive extru-sion product was determined by gel permeation chro-matography using chloroform as a solvent andpolystyrene as a standard. The Mark-Houwink con-stants for poly-BMAat 25°C are K= 4.37*10-5 dl/gand a = 0.80 (14). For poly-HPMA these values areunknown, so it was assumed that they are the sameas for poly-BMA.The deviation in molecular weightsobtained was about 500 g/mol, if the influence ofpossibly inaccurate Mark-Houwink constants wasneglected.

The composition of one series of samples was ana-lyzed using element analysis. The copolymer was pu-rified by removing the remaining monomer. For thispurpose, a concentrated solution of the copolymer inacetone was poured out in petroleum-ether (40/60).The precipitated copolymer was filtrated and dried ina vacuum stove.

RESULTS AND DISCUSSION

The properties of the reactive extrusion product arefixed by the reaction temperature, residence time,and mixing. These can be varied by changing theadjustable reaction and extrusion parameters. Thus,the analysis of the influence of each parameter on theproduct is an important resource in investigating thereactive extrusion process qualitatively.

The Screw Rotation Rate

The screw rotation rate was varied between 0.2 and0.8 rps at various values of the throughput. Figure 9shows the results of these experiments. Though anincrease in rotation rate involves a decrease in resi-



Fig. 9. Influence oj screw rotation rate on conversion at T=130"C and three different throughputs:. 19 g/min; • 41g/min; & 61 g/min.

dence time through which a decrease in conversion isexpected, it has no influence on the conversion of thepolymerization. In all probability this is due to bettermixing upon increasing the screw rotation rate, boththrough an increase in leakage flows and througha direct improvement, thereby increasing the heattransfer coefficient.

In the feed zone of the extruder, the monomer isheated by the extruder barrel to the temperature atwhich polymerization starts. At a high screw rotationrate, where the heat transfer coefficient is high, themonomer will reach the reaction temperature muchfaster than at a low screw rotation rate. Although anincrease in screw rotation rate causes a decrease inthe overall residence time, the time in which thematerial can react (the effective residence time) doesnot decrease accordingly, as a consequence of theearlier starting point with regard to low rotation rates.

After the polymerization has started, heat will beproduced by the exothermic reaction. When the heattransfer coefficient is high and the mixing is good, thetemperature of the material within the screw cham-bers is relatively low and any possible temperatureinhomogeneities are small. The low temperature willcause the material at the end of the extruder, wherethe conversion is high, to be very viscous, resulting ina diffusion limitation of the long polymer chain rad-icals. The gel effect will therefore appear, causinga higher conversion than expected. Both the earlierstarting point of the reaction as the appearance of thegel effect will counterbalance the reduction in trueresidence time.

These observations are in agreement with the re-sults obtained by analyzing the average molecul-ar weight shown in Fig. 10. Temperature inhomo-geneities at a low screw rotation rate cause thedecomposition rate of the initiator to increase, result-ing in a low molecular weight. At a high rotation rate,

Fig. 10. Influence oj screw rotation rate on number averagemolecular weight at T= 130"C and three different through-puts: • 199/min; • 41 g/min; D 61 g/min.

the gel effect will bring about an increase in molecu-lar weight.

The Throughput

From the results of the experiments, as shown inFig. 11, an increase in the throughput appears todecrease the conversion of the polymerization quitedramatically. Although an increase in the throughputdoes somewhat decrease the residence time, this ef-fect is not big enough to account completely for theconversion decrease. The same reasoning as with theinfluence of the screw rotation rate can be used.Along with the decrease in residence time, a decreasein the heat transfer coefficient and mixing occurs.These effects cause the monomer to reach its reac-tion temperature further down the extruder, so theeffective residence time decreases even more while

Fig. 11. Influence oj throughput on conversion at T= 130"Cand N= 0.5 1/ s.


H. A. Jongbloed, R. K. S. Mulder, and L. P. B. M. Janssenincreasing the throughput. The average temperatureof the material will be higher, so the polymerizationrate will increase through an increase in reaction rateconstants. At the same time, the viscosity of thematerial and therewith the intensity of the gel effectand the conversion decrease. The increase in thepolymerization rate obviously does not counterbal-ance the decrease in conversion consequent on thedecrease in the gel effect.

Through the effects mentioned above, the aver-age molecular weight decreases with increasing thethroughput (see Fig. 12).

The Barrel Temperature

On increasing the temperature of the extruder bar-rel the conversion initially increases, but above atemperature of about 125°C it decreases again (Fig.13).The conversion increase is caused by an increasein the polymerization rate. Because of the relativelylow temperature, the gel effect will still be present.The decrease in conversion above 125°C originatesfrom three effects. In the first place, an increase intemperature causes a decrease in viscosity towardsthe end of the extruder and in the die. Therefore thenumber of fully filled chambers falls off, causing adecrease in residence time. This effect was experi-mentally observed: above a temperature of 130°C, alarge pressure drop occurred. Stuber (21) derived amodel for the polymerization of MMA,in which thetemperature and conversion of the polymer-monomermixture are related to the viscosity. It can be con-cluded from this model that at high conversions thetemperature does have a fairly large influence on theviscosity. This affirms the above chain of reasoning.

The second cause for the decrease in conversion isa decrease in gel effect as a result of the high temper-ature and low viscosity. At the same time, a loweringof the initiator efficiency at a high temperature and apossible occurrence of a thermodynamic equilibrium

Fig. 12. Influence oj throughput on number average molecu-lar weight at T=130°C: • N=0.21/s; . N=0.51/s; &

N= 0.81/s.

Fig. 13. Influence oj temperature on conversion at N = 0.51/ sand Q = 41 g/ min.

could also bring about the conversion decrease. Thehigher the temperature, the higher the equilibriummonomer concentration and thus the lower the con-version.

Because a rise in the temperature causes an in-crease in the decomposition rate of the initiator, theaverage molecular weight of the polymer decreaseswith increasing temperature (Fig. 14). Both the de-crease in molecular weight and the temperature riseresult in a lower viscosity, so the gel effect reduces,intensifYing the effect of temperature on molecularweight.

The samples of this series of experiments wereused to determine the copolymer composition withelement analysis. Should a non-azeotropic copoly-merization occur after all, the composition of thesamples should be influenced by the differences intemperature and degrees of conversion. The results

Fig. 14. Influence oJ temperature on number average molecu-lar weight at N = 0.5 1/ sand Q = 41 g/ min.



visualized in Fig. 15 show that the composition isalmost constant and corresponds fairly well to theexpected monomer incorporation ratio.

The Die Resistance or Fully Filled Length

The die resistance can be varied by changing thesize of the die outlet. An increase in the die resistanceincreases the fully filled length and therewith theresidence time, while all other parameters remainconstant. The conversion therefore increases too (Fig.16). However, the maximum attainable conversionappears to be limited to a value of 92%, indicatingthat the residence time is not the only factor deter-mining the conversion. At high conversions, the de-propagation reaction might play an important role, sothe conversion is limited by the thermodynamic equi-

Fig. 15. 1nfluence oj temperature on fraction oj HPMA in thepolymer.

Fig. 16. 1nfluence oj die pressure on conversion at N = 0.8lis, T= 130°C, and throughputs of: • 51 glmin; D 61glmin.

Fig. 17. 1nfluence oj post-initiation on conversion at varyingthe main throughput at N = 0.5 1I sand T= 130°C: • with-out initiator side Jeed; • with initiator side Jeed.

librium. Moreover, the amount and efficiency of theinitiator could be a limiting factor.

To investigate the possibility of increasing the max-imum attainable conversion by enlarging the amountof initiator, a concentrated solution of 1.2 mol/l ini-tiator in monomer was continuously fed through aside feed to the third zone of the extruder during thereactive extrusion process. The amount of the sidefeed was adjusted so as to ensure a constant ratio ofmonomer and initiator within the extruder. Figure 17shows the results of these experiments. At a through-put of 20 g/min, the conversion was raised from 92%to 96%, which means that an extra reduction of 50%of the remaining monomer was attained with regardto the process without initiator side feed. The amountof initiator is therefore a significant parameter in thereactive extrusion process.

In the same way as the conversion, the averagemolecular weight increases initially and reaches amaximum value further on. In most polymerizationsystems, diffusion limitation of the monomer doesnot take place until very high conversions if the poly-merization is carried out above the glass transitiontemperature. Even when the amount of active initia-tor has reached zero at high conversions, the chancesare that there are still polymer chain radicals presentas a result of the diffusion limitation of these chains.The limitation in molecular weight therefore shouldnot occur as a result of lack of radicals or monomerdiffusion limitation. The most likely explanation,therefore, is the appearance of an equilibrium, whichtends to limit both the conversion and the averagemolecular weight (Fig. 18).

SUMMARY AND CONCLUSIONS

The use of an extruder as a polymerization reactoris not only restricted to homopolymerizations, suchas the polymerization of MMA(21) and BMA (5), but



Fig. 18. 1nfluence of die pressure on number average molecu-lar weight at N = 0.8 1/ s, T= 130"C, and throughputs oj: •51 g/min; D 61 g/min.

can be extended to copolymerizations as well. Thecopolymerization of BMAwith HPMAworks out verywell in the closely intermeshing counterrotating twin-screw extruder. Other radical addition copolymeriza-tions may be carried out in an extruder equally well.Care has to be taken, however, in choosing the typesof copolymerizations performed. The polymerizationtime needed to attain a certain required degree ofconversion should not exceed the economical resi-dence time of the material in the extruder. Besides,choosing any monomers with different reactivity ra-tios in a copolymerization could cause a drift in thecomposition distribution of the product.

The course of the polymerization and the propertiesof the product are strongly dependent on the adjust-ment of the reaction and extrusion parameters, suchas the screw rotation rate, throughput, temperature,and die resistance. An investigation was carried outwith the azeotropic copolymerization of BMA withHPMAto unravel the effects of these parameters onthe reactive extrusion product. The following resultswere obtained.

• The screw rotation rate has no influence on theconversion of the polymerization. On increasingthe rotation rate, the decrease in residence time iscounterbalanced by an increase in effective reac-tion time and gel effect.

• An increase in screw rotation rate causes an in-crease in average molecular weight, because of alessening of temperature inhomogeneities and anincrease in gel effect.

• The conversion of the polymerization decreases withincreasing throughput as a result of a shorter resi-dence time and effective reaction time.

• Temperature inhomogeneities on increasing thethroughput cause the average molecular weight todecrease.

• An increase in temperature initially causes an in-

crease in conversion as a result of an increase inreaction rate constants, then causes a decrease inconversion as a result of a decrease in filled lengthand gel effect and the appearance of a thermody-namic equilibrium.

• Owing to an increase in the decomposition rateof the initiator, the average molecular weight de-creases when increasing the temperature.

• An increase of the die resistance and therewith thefully filled length positively affects the conversionas a result of an increased residence time. Initiatordeficiency and a thermodynamic equilibrium limitthe conversion.

• The average molecular weight increases when in-creasing the fully filled length.

ACKNOWLEDGMENTThe authors wish to thank DSM Research (The

Netherlands) for making this research financiallypossible.

NOMENCLATUREA = Geometrical constant (m3).a =Mark-Houwink constant (-).B = Geometrical constant (m3).D = Diameter (m).D Hp = Polymerization enthalpy (JImol).I = InitiatorK =Mark-Houwink constant (-).kd = Initiator decomposition rate constant (lis).kdP = Depropagation rate constant O/s).ki = Initiation rate constant (m3/mol' s).kp = Propagation rate constant (m3/mol· s).ktc =Termination by combination rate constant

(m3/mol· s).ktd =Termination by disproportionation rate

constant (m3/mol· s)L = Extruder length (m).M =MonomerMi· =Growing radical of chain length im =Number of thread starts ( - ).N = Screw rotation rate O/s).D Pc = Inter-chamber pressure gradient (Pa).Q =Throughput (m3Is).Ql =Totalleakage flow (m3Is).Qth =Theoretical throughput (m3Is).R = Gas constant (J/mol· K).R' = Initiator radicalrI = Reactivity ratio ( - ).r2 = Reactivity ratio ( - ).D Sp = Polymerization entropy (J/mol' K).Tc = Ceiling temperature (K, °C).Vc = Chamber volume (m3).up = Polymerization rate (mol/m3. s).h =Viscosity (Pa· s).[] =Concentration (mol/m3).

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Received June 21, 1993


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