Manufacture of butyl rubber

The invention relates to a continuous process for the manufacture of butyl rubber in which the monomers are polymerized using a halogenated polymerization medium in a self-cleaning screw extruder at a constant pressure under boiling plug-flow conditions by means of a modified aluminium halide catalyst which operates at high temperature (relative to prior art very low temperatures of e.g. -100.degree. C.). In contrast to prior art slurry processes, the polymerization mixture forms a sticky, highly viscous mass which is conveyed by screw action to the extruder outlet as polymerization proceeds. The butyl rubber product is separated at a high concentration from vapors of unreacted monomers and polymerization medium which are recycled. Cooling is effected by vaporization of a portion of the mixture of monomers and polymerization medium in a flash tank before supply to the extruder and returning the vaporized fraction and recycle vapors to the flash tank through, in succession, a compressor, heat exchanger and throttle valve. By using the process of the invention, the total energy consumption in butyl rubber manufacture can be cut to about 20% of that of prior art processes.

This invention relates to the manufacture of butyl rubber. The expression 
"butyl rubber" is used in this specification in a wide sense to mean a 
copolymer of at least 90% by weight of isobutene, with a small amount of 
one or more conjugated diene comonomers, with optionally minor amounts of 
one or more comonomers. Typically, the conjugated diene is isoprene but 
the invention is not limited to copolymers of isobutene and isoprene 
alone. In particular, the invention relates to a continuous process for 
the manufacture of butyl rubber. 
Butyl rubber has been known for almost fifty years. Usually the rubber is 
produced commercially by a semi-batch copolymerisation of isobutene and 
conjugated diene comonomer carried out as a slurry process using methylene 
chloride reaction medium and a Lewis acid catalyst, typically aluminium 
trichloride. The reaction is carried out at extremely low temperatures 
(about -100.degree. C.) in order to obtain high molecular weight polymers. 
Details of some prior art processes to produce butyl rubber are given, for 
example, in Encyclopaedia of Chemical Technology (Kirk, R. E. and Othmer, 
D. F.) 3rd edn. John Wiley & Sons, New York (1978)--volume 8 pages 470-484 
(hereinafter "Kirk-Othmer"). Some prior art continuous processes for the 
manufacture of butyl rubber are discussed in "Vinyl and Related Polymers" 
Schildknecht, C. E., John Wiley & Sons, New York (1952) pages 576-580 
(hereinafter called "Schildknecht"). 
A major problem with the commercial manufacture of butyl rubber by prior 
art processes is the high capital and operating cost (mainly energy cost) 
of the refrigeration plant required to maintain the very low temperatures 
of polymerisation. For example, the cost of the refrigeration plant may 
represent about 35% of the capital cost and 75% of the energy cost of the 
entire plant. 
In recent years, catalysts have become available which enable the 
polymerisation reaction to be carried out at higher temperatures (e.g. 
-45.degree. C.). Such catalysts are known as "high" temperature catalysts 
and are described, for example, in "Copolymerisation of Isobutene and 
Isoprene at "High" Temperatures with Syncatalyst Systems Based on 
Aluminium Organic Compounds", Cesca, S., Bruzzone, M., Prioa, A., 
Ferraris, G. and Giusti, P., Rubb. Chem. Tech. 49 937(1976) and U.S. Pat. 
Nos. 4,154,916, 4,171,414, 4,146,692, 4,103,079 and 4,151,113 (Exxon 
Research & Engineering). The development of these "high" temperature 
catalysts has gone some way to alleviating the problem and expense of 
refrigeration but the temperatures involved are still well below ambient 
temperature. An associated problem, which exacerabates the problem of 
refrigeration,is that in prior art processes polymer concentration at the 
reactor outlet is relatively low, e.g. 20-30% by weight. This necessitates 
the removal, recycling and cooling of large quantities of reaction medium. 
Another major problem in the commercial manufacture of butyl rubber, which 
increases the difficulty and cost of production, is fouling of the 
reactor. Polymer builds up on the surfaces of the reactor and of equipment 
such as agitators, baffles and temperature probes inside the reactor. This 
build up of polymer also hinders heat transfer to the reactor coolant 
tubes and jacket. Frequent cleaning is necessary and there is consequent 
lost production (down) time. In commercial practice, it is usual to have 
two or three reactor lines so that whilst one line is in use for 
polymerisation, the others are cleaned and made ready for production (see 
"Kirk-Othmer" page 475). This greatly increases the capital cost of the 
plant. 
In prior art processes for the manufacture of butyl rubber, the reactors 
used are stirred cylindrical tank reactors having jackets and tubes for 
circulation of liquid coolant, such as liquid ethylene. See, for example 
"Kirk-Othmer" page 474, and U.S. Pat. Nos. 4,146,692 and 4,096,320. Quite 
complicated reactor designs have been evolved to reduce the problems of 
cooling and reactor fouling. For example, European Patent Application No. 
0,053,585 (Chemical Abstracts No. 97-P111113) describes a cylindrical 
reactor for the polymerisation of hydrocarbons in solution or suspension, 
which has an internal rotating cooling unit and scrapers to keep the 
reactor surfaces clean. One particular use for this reactor is in the 
manufacture of butyl rubber. This and other cylindrical reactors are also 
described in a recent paper (ICP 1984, 12 (10) 157-60-C.A. ref 104 
89162r). 
We have now found that a quite different apparatus can be used for the 
manufacture of butyl rubber provided that the apparatus is combined in a 
particular process arrangement and operated under specific process 
conditions. 
According to the present invention a continuous process for the manufacture 
of butyl rubber by the polymerisation of a monomer mixture comprising 
essentially at least 90% by weight of isobutene and 0.5 to 10%. by weight 
(based on the weight of mixture) of at least one conjugated diene 
comonomer by continuously supplying a feed stream of the monomer mixture 
and polymerisation medium together with catalyst to an extruder type 
reactor and polymerising therein, is characterised by the steps of (a) 
cooling the feed stream of monomers and halogenated hydrocarbon 
polymerisation medium to a polymerisation temperature of -70.degree. C. to 
15.degree. C. by vapourisation of a portion thereof under reduced 
pressure; (b) supplying the cooled feed stream to a reactor which is a 
self-cleaning screw extruder and polymerising therein under boiling 
plug-flow conditions at a constant pressure of 0.1 to 4 Bar using a "high" 
temperature aluminium halide catalyst and (c) removing at the reactor 
outlet polymer product at a concentration of at least 50% for recovery and 
a vapourised mixture of unreacted monomers and polymerisation medium, for 
recycle. 
Each one of the steps (a), (b) and (c) above is critical for the successful 
manufacture of butyl rubber by the continuous process of the present 
invention. A most important advantage is that by using the process of the 
present invention the total energy consumption in the manufacture of butyl 
rubber can be reduced to about 20% of that of prior art processes which 
use stirred tank type reactors. 
Self-cleaning screw extruders have been known for some years. They are 
useful for processing pasty, highly viscous materials where a mixing and 
kneading effect is required. Apart from mixing and kneading operations, 
typical applications described in the commercial literature include 
heating and cooling, melting, crystallization, sublimation, concentration 
and drying of viscous materials, solutions and melts and reactions with 
pasty-viscous intermediate phases. In the early 1970s, the use of 
self-cleaning two-shaft extruders was described for the continuous 
solvent-free production of block copolymers from 1,3 diolefins and vinyl 
aromatic compounds at 50.degree. to 150.degree. C. and the continuous 
homopolymerisation or copolymerisation of a 1,3 diene or a mixture of 1,3 
diene and at least one other polymerisable monomer, in the absence of 
polymerisation solvent at -20.degree. C. to +150.degree. C., generally at 
a screw polymerisation temperature of 10.degree. or 20.degree. C. (inlet) 
and 75.degree. C. or 80.degree. C. (outlet). See British Patent Nos. 
1,218,147 and 1,347,088. Their use has also been described in the 
preparation of diene polymers using rare earth type catalysts in the 
absence or substantial absence of solvent or diluent at 
50.degree.-60.degree. (inlet) and 80.degree.-100.degree. C. (outlet). See 
European Patent Application No. 0,127,236. As can be seen, the 
polymerisation temperatures in such processes is above ambient 
temperatures so that the problem of reactor cooling is minor, compared 
with the problem of operation at extremely low temperatures. Self-cleaning 
screw extruders have not been hitherto used or described for the 
manufacture of butyl rubber, because of the problem of maintaining the 
extremely low polymerisation temperatures (e.g. -100.degree. C.) required 
since there is substantially no heat transfer through the extruder walls. 
Butyl rubber and processes for its manufacture were extensively studied in 
the early years e.g. 1940-1945, and many attempts were made to devise 
improved manufacturing processes. In this context see, for example, 
"Schildknecht" pages 576 to 582 and the several patents referred to 
therein. In U.S. Pat. No. 2,494,588 (1944), which is not referred to by 
Schildknecht and which seems to have been largely ignored in the 
literature generally, there is a proposal for the production of solid 
polymers of isobutylene using an extruder-reactor of specific design. The 
apparatus described therein has a large diameter first portion and a small 
diameter second portion and a single extruder screw within the bore. The 
process described appears never to have been put into commercial operation 
and it does not appear to have been used as a basis for further research. 
In contrast to the process described in U.S. Pat. No. 2,494,588, the 
process of our invention has the following critical characteristics: 
1. The process carried out in a self-cleaning screw extruder under plug 
flow conditions. This is essential because at the polymerisation 
temperatures used, the butyl rubber swells in the reaction medium so that 
the reaction mixture is a sticky sludge as opposed to a polymer slurry of 
prior art processes. If some other type of continuous reactor is used, the 
residence time cannot be adequately controlled, side reactions occur 
leading to gel formation and poor quality product. 
2. The catalyst used is a "high" temperature aluminium halide catalyst and 
the polymerisation temperature is -70.degree. C. to +15.degree. C., 
preferably -50.degree. C. or above. By operating under boiling conditions 
and cooling the feed stream by vapourisation of a portion thereof under 
reduced pressure (step a), it is not possible to attain the extremely low 
polymerisation temperatures (e.g. -100.degree. C.) of prior art processes. 
3. The reaction mixture is cooled to the polymerisation temperature by 
vapourisation of a portion thereof under reduced pressure (step a) and 
maintained at the polymerisation temperature by polymerising under boiling 
plug flow conditions at a constant pressure of 0.1 to 4 Bar. No separate 
refrigerant, such as ethylene, is passed through the reactor and boiled 
off. As mentioned above, in screw type extruders there is substantially no 
heat transfer through the reactor walls so that the process arrangement 
described is critical to the attainment of the polymerisation temperature 
specified. 
4. The polymer recovered at the reactor outlet is at a concentration of at 
least 50% by weight. 
Another early prior art process is disclosed in U.K. Pat. No. 561,324 which 
relates particularly to means for separating the formed polymer from the 
recycle stream in which the polymerisation product, as it is formed, is 
immediately subjected to a kneading process combined with a movement 
whereby it is conveyed to an extruder, such operations being effected in a 
series of vessels closed to the exterior whereby the product of 
polymerisation is kneaded and extruded without any substantial loss of the 
diluent/refrigerant or reactant materials. One alternative embodiment of 
the device of the invention described in U.K. Pat. No. 561,324 bears some 
superficial resemblance to the apparatus used in the process of the 
present invention. However the apparatus disclosed is not a self-cleaning 
screw extruder, essential in the present invention, the cooling 
arrangements are quite different in that liquid ethylene is passed through 
the apparatus, the reaction temperature is the conventional extremely low 
temperature, preferably -40.degree. C. to -80.degree. C. and the catalyst 
described is boron trifluoride. 
The self-cleaning screw extruder used in the process of the present 
invention may be a twin-screw extruder or may be a specially designed 
apparatus, each capable of carrying a material of progressively increasing 
viscosity from inlet to outlet in a plug-flow type movement. Twin-screw 
extruders have two screws which intermesh with each other and consequently 
are almost completely self-cleaning. However one disadvantage of twin 
screw extruders is that the free volume of commercially available models 
is low, generally of only a few hundred litres, which is very small when 
compared with the commercial requirements of polymer production. Examples 
of twin-screw extruders are those available from such manufacturers as 
Werner and Pfleider of West Germany and Baker and Perkins of U.S.A. 
Screw-type self-cleaning extruders which are specially designed to handle 
viscous materials are available. Commercially available examples are the 
AP-CONTI and DISCOTHERM B type machines supplied by Dipl Inj. H. List of 
Switzerland. The former has two shafts running in a shaped casing. The 
main shaft with radially mounted, heatable plates joined by kneading bars 
is purged by the cleaning shaft and its kneading frames, which rotates 
faster than the main shaft (usually at four times the speed of the main 
shaft). The free volume of the biggest commercial machine is 2.6 m.sup.3. 
The latter type of self-cleaning extruder (i.e. DISCOTHERM B) has only one 
shaft which carries disc-like heatable elements with peripheral mixing 
bars, set at an angle, rotating in a cylindrical housing. The free volume 
of the biggest commercial machine of this type is 10 m.sup.3. Further 
details of these two types of specially designed apparatus are given in 
the literature available from the manufacturers. 
Reaction is carried out in halogenated hydrocarbon medium which may be a 
diluent or solvent for some or all of the reaction ingredients (but not 
the formed polymer). As is well known in the manufacture of butyl rubber, 
a preferred polymerisation medium is a halogenated aliphatic hydrocarbon, 
more particularly a halo alkane containing one or two carbon atoms per 
molecule. Methyl chloride is the one most usually employed. Other examples 
of halogenated hydrocarbons that may be used in the process of the 
invention are ethyl chloride and methylene chloride. In a preferred 
embodiment of our invention, the polymerisation medium comprises a mixture 
of a halogenated hydrocarbon and a non-halogenated hydrocarbon, the 
proportion of the latter being up to 50% of the total by volume, 
preferably 5 to 30% by volume, more preferably 5% to 20% by volume. Under 
such conditions the polymer formed in the polymerisation together with 
unreactd monomer and polymerisation medium (hereinafter called "solid 
phase") performs as a viscous swollen mass (i.e. it is not a free flowing 
slurry like sand, as in prior art processes). If the proportion of 
non-halogenated hydrocarbon is below about 5%, the solid phase in the 
reactor is not swollen enough with polymerisation medium and the particles 
of polymer remain discrete, i.e. as very small particles. At higher levels 
(e.g. up to 20%) the polymer particles become sticky enough to stick 
together forming a cohesive viscous mass. Furthermore, at these levels, 
the liquid phase (i.e. the monomer/polymerisation medium mixture in excess 
of that swelling the solid phase) does not dissolve excessive quantities 
of polymer, the viscosity is kept low and foam problems do not arise 
during a sudden boil off. Using the halogenated hydrocarbon nonhalogenated 
hydrocarbon mixture in this way enables the polymer concentration in the 
solid phase, and thus in the reactor output to be about 50-65% by weight 
(e.g. 60%), compared with a concentration of 10 to 20% weight in a free 
flowing slurry as discharged at a very low temperature from the reactor in 
conventional processes. 
The monomer mix which is polymerised in the process of the invention 
comprises essentially isobutene and at least one conjugated diene. The 
conjugated diene(s) employed may have from 4 to 14 carbon atoms (e.g. 4 to 
8 carbon atoms). Examples of such dienes are isoprene, 1,3 butadiene and 
2,3-dimethyl-1,3-butadiene, piperylene, cyclopentadiene, methyl 
cyclopentadiene and cyclohexadiene. Isoprene is preferred. The proportion 
of conjugated diene used is generally in the range of 0.5% to 10% by 
weight, preferably 0.5 to 5% by weight, e.g. 1 to 3%. Optionally other 
comonomer(s) may be included, if desired. The concentration of monomer in 
the polymerisation medium supplied to the reactor is high, e.g. 60% to 70% 
by weight, to keep the formed polymer concentration in the extruder at a 
high level. 
The catalyst used may be an aluminium halide dissolved in halogenated 
hydrocarbon, such as methyl chloride, but is preferably one of the "high" 
temperature catalysts which have been developed, such as those referred to 
above in this specification, dissolved in solvent. The catalyst may be an 
aluminium alkyl halide which has been reacted with a halogen or mixed 
halogen as described in British Pat. No. 1,362,295 or an aluminium alkyl, 
alkyl halide, alkoxy aluminium halide or a selected Lewis acid which has 
been reacted with one or more of a wide variety of other halogen 
containing compounds, such as described in British Pat. Nos. 1,407,414 to 
1,407,420 or an improved catalyst composition which is a hydrocarbon 
soluble catalyst formed by prereacting an alkyl aluminium halide with 
halogen, hydrogen halide or mixed halogen in critical mole ratios (such as 
described in U.S. Pat. No. 4,171,414) or a halogenated organo aluminium 
catalyst composition in which the mole ratio of halogen to alkyl aluminium 
halide is 0.3 to 0.8 (such as described in British Pat. No. 1,542,319). In 
principle however there appears to be no criticality upon the type of 
"high" temperature catalyst employed or the amount of catalyst used. 
The polymerisation is carried out at a constant pressure (i.e. is an 
isobaric polymerisation), the pressure corresponding to a boiling 
temperature of the mixture of monomer/polymerisation medium which is in 
the range -70.degree. C. to +15.degree. C. and preferably above 
-50.degree. C. In a preferred embodiment of the invention, the monomers 
and polymerisation medium are continuously supplied to the extruder via a 
flash drum maintained at reduced pressure so that a fraction of said 
monomers and polymerisation medium mixture is vapourised, cooling the 
remaining contents of the flash drum to the polymerisation temperature 
(which might be, for example, -45.degree. C.). Furthermore by arranging a 
compressor, a heat exchanger and throttle valve in the return line from 
the extruder, vapourised monomer-polymerisation medium mixture removed 
from the extruder under boiling conditions can be returned via the flash 
drum to the reactor as a liquid at the same temperature as the vapourised 
liquid (e.g. -45.degree. C.) thus extracting from the extruder the latent 
heat of vapourisation for each unit weight of mixture recycled. This 
arrangement reduces the amount of fluid refrigerant employed compared with 
prior art polymerisation processes and thus enables significant reductions 
to be made in energy consumption. 
Polymer formed in the polymerisation is conveyed through the extruder as a 
viscous mass by the extruder screw(s) to the outlet where the polymer is 
separated from residual monomers and polymerisation medium in separating 
means. Preferably the separating means comprises a vertically mounted 
discharge screw (mechanical filter) above a twin screw desolventizer so 
that the formed polymer can be squeezed free of liquid 
monomer-polymerisation medium mixture. Gaseous monomer/polymerisation 
medium mixture vents to a recycle line for recycle after compression, heat 
exchange and passage through a throttle valve to the flash drum used to 
supply the extruder. 
Residual monomer-polymerisation medium removed in the desolventizer is 
recovered and recycled, leaving polymer product ready for finishing, e.g. 
baling and wrapping.

Referring to the drawing, the extruder 1. is a self-cleaning screw-type 
extruder such as one of those described in this specification. The outlet 
is fitted with a twin discharge screw 2 mounted vertically across the 
outlet and connecting via valve 12 to a twin screw desolventiser 3. A 
mixture of fresh monomers and polymerisation medium (hereinafter called 
solvent) is supplied via line 10, heat exchanger 22, line 23 and throttle 
valve 21 to flash drum 8 together with recycled monomers and solvent 
supplied via line 24. In flash drum 8, a lowering of pressure occurs which 
involves vapourisation of a fraction of the mixture of monomers and 
solvent. This brings down the temperature of the liquid to the 
polymerisation temperature. The cooled stream of monomers and solvent are 
supplied to extruder 1, via line 16, the stream being split and supplied 
at points along the extruder 1, via lines 16a, 16b, 16c and 16d to ensure 
thorough distribution. Catalyst solution is supplied at the head of the 
extruder screw via line 11. As catalyst enters the extruder, it forms 
polymer which is insoluble in the monomer-solvent mixture and which 
entraps catalyst. The formed polymer becomes swollen with a certain amount 
of monomer-solvent mixture according to the equilibrium concentration at 
the polymerisation temperature. Thus separate phases are formed consisting 
of (a) a so-called solid phase, comprising polymer containing entrapped 
catalyst and equilibrium monomer-solvent mixture, (b) a so-called liquid 
phase comprising monomer-solvent mixture and traces of dissolved polymer 
and (c) a vapour phase. 
As the "solid phase" is conveyed towards the outlet by the screw action of 
the extruder shaft, the catalyst entrapped in the solid phase forms more 
polymer, consuming the monomers swelling said phase. The monomers consumed 
are continuously replaced by fresh monomers, coming from the liquid phase 
across the solid phase-liquid phase interface, which is being continuously 
renewed by the screw action. At the extruder outlet section, solid phase 
is forced into the twin discharge screw 2. In discharge screw 2, the solid 
phase is squeezed by the screws so that any entrapped liquid phase 
fraction exceeding the equilibrium concentration is separated and 
overflows back towards the extruder 1, whereas the gaseous monomer-solvent 
mixture vents though the vapour outlet line 14, for recycle. Any solid 
phase fraction entrained by the vapours at this part of the discharge 
screw is disengaged along the screw and pushed back towards the 
desolventiser. Under the screw action of discharge screw 2, the solid 
phase is forced through valve 12, the opening of which is controlled by 
pressure control, 13. The solid phase enters heated twin screw 
desolventiser, 3, to which catalyst deactivator and antioxidant are 
supplied via supply lines (not shown). In the heated desolventiser 3, 
liquid within and swelling the polymer is vapourised and supplied via line 
17 to a conventional recovery section (not shown). The polymer, 
substantially free of solvent,is supplied via line 18 to a finishing 
section for baling and packaging. 
As mentioned above, the extruder contains a vapour phase since the process 
of the invention is carried out under boiling conditions, albeit at 
temperatures between -70.degree. C. to +15.degree. C., usually above 
-50.degree. C. and therefore usually at below atmospheric pressure. For 
example, at -45.degree. C. the reactor is kept at the equilibrium pressure 
of about 0.3 bar absolute. Reaction heat is therefore removed as latent 
heat of evaporation of the liquid phase. Vapourised liquid phase leaving 
the extruder via line 14, passes via knock-out drum 4 and line 19 to 
compressor 5. In the compressor 5, the vapours are compressed and then are 
cooled by cooling water or, preferably refrigerant, condensed in heat 
exchanger 6 for return to the flash drum 8, via line 24 and throttle valve 
7 through which the pressure decreases to that of the extruder. In flash 
drum 8, a portion of the feed vapourises and passes to compressor 5, via 
line 20, drum 4 and line 19 and liquid feed in drum 8 cools to the 
temperature of extruder 1, before supply via line 16. Any non condensables 
in heat exchanger 6 are removed via an outlet situated at the coldest 
section of the heat exchanger towards a vacuum pump (not shown) via line 
15. 
By the above described arrangement of compressor, heat exchanger and 
throttle valve and associated equipment all the boiled-off monomer-solvent 
mixture vapour from the extruder is recycled to the extruder as a liquid 
at the same temperature as said vapour thus extracting from the extruder 
the latent heat of vapourisation for each unit weight of mixture recycled. 
The pressure into the extruder is kept constant by controlling through 
conventional means, the flow rate of the vapours to compressor 5. At 
constant pressure, the temperature of the boiling liquid is constant, once 
a steady state of liquid composition is reached. 
The temperature of the solid phase of reaction mixture in reactor 1 would 
tend to increase because of the heat of polymerisation developed therein. 
However this tendency is countered by the intimate contact between solid 
and liquid phases in reactor 1 enabling polymerisation heat to be 
transferred to the liquid phase and thence to the ambient through 
condenser 6. It will be appreciated that this arrangement reduces the need 
to use a fluid refrigerant at a temperature below that of the 
polymerisation enabling significant energy consumption economies to be 
made. 
As emphasised above, each one of the steps (a), (b) and (c) of the 
continuous process of the present invention is indispensable for the 
successful manufacture of butyl rubber. Firstly, if any one of steps (a), 
(b) or (c) is omitted, it is not possible to obtain adequate temperature 
control in the self-cleaning screw extruder so as to maintain a 
polymerisation temperature of -70.degree. C. to +15.degree. C. Secondly, 
if the screw extruder, used in the process, is replaced by, for example, a 
cylindrical type reactor, as used in prior art processes, fouling of the 
reactor becomes a serious problem because of the rheological 
characteristics of the polymer mass at the polymerisation temperatures 
used. At these temperatures, the polymer is swollen with reaction medium 
and becomes a viscous sticky mass, which is very difficult to handle in 
cylindrical agitated reactors. Thirdly, if the concentration of formed 
polymer is less than 50% at the reactor outlet, not only is temperature 
control more difficult, but much higher volumes of liquid have to be 
circulated. This negates one object of the invention which is to keep the 
energy consumption of the process as low as possible. By achieving the 
high concentration stated, an energy saving of about 50% over conventional 
processes may be obtained. For this reason, the concentration of monomers 
in the feed stream is preferably 60% to 70% by weight based on the weight 
of feed stream. The butyl rubber product has a high molecular weight (Mn) 
i.e. at least 50,000, normally 100,000 to 500,000. 
EXAMPLE 
One particular embodiment of the process of the present invention will now 
be described by way of example only. In this example, a polymerisation of 
isobutylene with isoprene was carried out in the apparatus described 
hereinbefore with reference to the schematic flow diagram shown in the 
drawing. The amounts of reaction ingredients used, expressed in kilograms 
per hour at steady state and the temperature in degrees centigrade at the 
point of measurement indicated, are shown in the table. For convenience, 
points of measurement referred to are underlined in the following 
description. 
A feed stream comprising a mixture of solvent, (a mixture of halogenated 
hydrocarbon and non-halogenated hydrocarbon) isoprene and isobutylene, in 
the relative amounts shown, was continuously supplied to flash drum 8 via 
line 10, heat exchanger 22, line 23 and throttle valve 21. Flash drum 8 
was maintained at reduced pressure so that a proportion of the feed stream 
vapourised and passed via line 20 to knock-out drum 4 and thence via line 
19 to compressor 5. After compression and heat exchange in heat exchanger 
6 a stream passed via line 24 and throttle valve 7 is returned to the 
flash drum 8. The contents of the flash drum 8 at steady state was thus 
cooled to -45.degree. C., for supply to the screw-type extruder 1. A feed 
stream was withdrawn from flash drum 8 via line 16 and supplied via a 
plurality of supply lines 16a, 16b, 16c and 16d to extruder 1. A catalyst 
stream comprising aluminium diethyl chloride and chlorine as 
initiator/coinitiator was supplied via line 11. The initiator consumption 
was 0.3 to 3.0 g per kilogram of finished product (i.e. butyl rubber) and 
the coinitiator consumption was 0.2 g per kilogram of finished product. 
Polymer formed in the extruder was conveyed as a solid phase to the outlet 
by the screw action of the screw shafts and forced into the twin discharge 
screw 2 and thence to heated screw desolventiser 3. Butyl rubber finished 
product having a polymerised isoprene content of 3% wt. was recovered via 
line 18 at the rate of 10 kilogram per hour. Residual monomers and 
polymerisation medium separated in the discharge screw 2 are recycled via 
line 14 to knock-out drum 4 (and thence to flash drum 8 by the compressor 
route described above). Final removal of monomers and polymerisation 
medium is achieved as a vapour in the twin screw desolventiser 3 already 
referred to and the vapour removed was recovered via line 17. 
TABLE 
__________________________________________________________________________ 
Stream 
Initial feed 
Vapourised 
Return 
Reactor feed 
Product 
Recycle 
Recovery 
Components 
23 20 24 16 18 14 17 
__________________________________________________________________________ 
Solvent 
4.9 4.9 
Isoprene 
0.3 11.8 41.8 47 30 0.1 
i-butylene 
11.8 2 
Butyl rubber 10 
Temp. .degree.C. 
0 -45 0 -45 -45 
__________________________________________________________________________ 
Units = kilograms per hour.