WO2015010943A1

WO2015010943A1 - Method for producing high-molecular-weight isobutene homopolymers

Info

Publication number: WO2015010943A1
Application number: PCT/EP2014/065066
Authority: WO
Inventors: Rosa Corberan Roc; Robert Haschick; Marco-Christian Volland; Alexander GEHR; Andreas Daiss; Klaus MÜHLBACH; Isa Alexandra QUEIROZ DA FONSECA; Christiane Lange; Oliver Soetje; Markus Brym; Thomas Wettling; Joachim Eichhorn; Stefan Hirsch; Jens Hoferichter; Dirk Opfermann; Eckard Schauss; Markus Weis
Original assignee: Basf Se
Priority date: 2013-07-25
Filing date: 2014-07-15
Publication date: 2015-01-29

Abstract

The invention relates to the production of isobutene homopolymers having a weight-average molecular weight of 75,000 to 10,000,000 by polymerizing isobutene in a polymerization device in the liquid phase at temperatures of -80°C to -190°C in a suitable inert solvent and in the presence of a polymerization catalyst complex based on Lewis acids and co-catalysts, wherein isobutene, inert solvent, and polymerization catalyst complex are fed to the mixing chamber of a mixing device in at least three partial flows and are mixed there, said mixture is fed into the liquid phase space of the polymerization device, and the polymerization of the isobutene is allowed to proceed there.

Description

Description: The present invention relates to an improved process for the preparation of isobutene homopolymers having a weight-average molecular weight of 75,000 to 10,000,000 by polymerization of isobutene in a liquid phase polymerization apparatus at temperatures from -80 ° C to -190 ° C in a suitable inert solvent which is liquid under the polymerization conditions in the presence of a Lewis acid-based polymerization catalyst complex and co-catalysts.

Efficient and specification-specific production processes of higher molecular weight isobutene homopolymers generally require very low polymerization temperatures. A common method for producing such isobutene homopolymers is the so-called "BASF belt process", in which liquid isobutene is passed with boron trifluoride as the polymerization catalyst and a high excess of liquid ethene on an endless steel strip of 50 to 60 cm wide , which is designed trough-shaped by a suitable guide and is located in a gas-tight, cylindrical housing. By continuous evaporation of ethene at atmospheric pressure, a temperature of -104 ° C is established. The heat of polymerization is thereby completely dissipated. The vaporized ethene is collected, purified and recycled. The resulting polyisobutenes are freed from adhering ethene and residual monomers by degassing. The type of polymerization leads to a virtually complete isobutene conversion. In the BASF strip process, the polymerization temperature can be controlled simply and reliably owing to the boiling cooling, ie by formation of large vapor passages. A disadvantage of the BASF strip process, however, is that due to lack of movement of the reaction mixture on the belt no adequate mixing of the reaction mixture and thus no product surface renewal takes place, which can adversely affect the product properties. This leads, for example, to an inhomogeneous distribution of the ethene used for the boiling cooling and, associated therewith, to a local overheating of the reaction mixture as soon as the ethene has evaporated. In addition, an explosive boiling of the reaction mixture can occur when overheated areas and ethenreiche cold areas contact each other, which then leads to contamination of the reactor wall by entrainment of polymerizing reaction mixture. Another disadvantage is that the inhomogeneous temperature distribution causes an undesirable broadening of the molecular weight distribution of the polymer, which is associated with unfavorable product properties. Another disadvantage of the BASF belt process is that the steel strip is subject to wear and thus causes high maintenance costs. Furthermore, it is disadvantageous in the BASF strip process that the reactor walls and the product feed are not cooled in the downstream working part (usually an extruder); since polyisobutene is highly tacky above its glass transition temperature, this leads to a marked adhesion of the reactor walls with polymers, which makes an increased cleaning effort necessary. A further disadvantage of the BASF strip process is that boron trifluoride contained in the recycled ethene stream has a highly corrosive effect at relatively high temperatures, which causes a high maintenance outlay on the ethene processing cycle.

Another common method for the preparation of higher molecular weight isobutene homopolymers is the "ηοη-Breiverfahren" ("Exxon slurry process"), in which the polymerization is carried out in a cooling jacket, which is charged with liquid ethene, stirred at -80 to -85 ° C. As a catalyst system, anhydrous aluminum chloride in methyl chloride is used, and as a result of the very vigorous stirring, the polymer precipitates as a slurry consisting of small droplets, which overflows into a degassing vessel via an intermediate vessel, where the slurry is steamed and treated with hot water so that the volatiles (substantially unreacted isobutene and methyl chloride) can be removed and subjected to reprocessing The remaining aqueous slurry of polymer particles is worked up by removing catalyst residues, solvent residues and isobutene residues.

Although intensive mixing and product surface renewal take place in the Exxon Breeding process, the polymerization temperature is difficult to control solely by the jacket cooling. Since the adhesion of the polymer to the reactor and apparatus walls can not be completely prevented, the reactor and apparatus must be cleaned from time to time.

The BASF strip process and the Exxon Breeding process are further described in Ullmann's Encyclopedia of Industrial Chemistry, 5th Edition, Vol. A21, pages 555-561, under the heading "

Polyisobutylenes "described.

It is an object of the present invention to provide an easy to carry out, efficient and economical process for the preparation of relatively high molecular weight isobutene homopolymers which, with regard to the product parameters to be adjusted such as molecular weight, polydispersity and residual monomer content, permit reliable control of the polymerization and are easy to clean and easy to handle In particular, a non-adhesive product before work-up supplies. Because of the more thorough mixing of the reaction product, the polymerization should be carried out in a conventional closed reactor and in the disperse phase in a non-miscible fluid or homogeneously mixed in a miscible fluid, i. in a suitable solvent.

The polymerization of isobutene to higher molecular weight isobutene homopolymers in conventional closed reactors in solvents or diluents is also known from other documents in addition to the described Exxon-Breiverfahren. Thus, DE-OS 2 061 289 discloses an isobutene polymerization process in which isobutene between 0 ° C and -160 ° C in an inert diluent such as ethylene, methane, ethane or propane using boron trifluoride as Catalyst is polymerized in the presence of a solution of formaldehyde in an alcohol such as isobutanol as a molecular weight regulator in a reaction flask to higher molecular weight polyisobutene. In the monograph "Polymerization and Polycondensation Processes, Advances in Chemistry Series 34" (1961) JP Kennedy and RM Thomas describe in their article "Cationic Polymerization at Ultralow Temperatures" on page 1 1 1 -1 19 the polymerization of isobutene in a propane. Isopentane mixture in a cooled reactor at -30 ° C to -190 ° C by means of an aluminum trichloride catalyst to higher molecular weight polyisobutene. Aluminum trichloride has the disadvantage that, as a non-volatile catalyst, it hinders the subsequent purification of the polyisobutene. Reaction accelerators or chain length regulators are not used.

It is known from the literature article "Fundamental Studies on Cationic Polymerization IV - Homo- and Copolymerizations with Various Catalysts" by JP Kennedy and RG Squires in Polymer 6, pages 579-587, 1965, that isobutene catalyses boron trifluoride in alkyl chloride solvents. 30 ° C to -146 ° C in the presence of isoprene to higher molecular weight polyisobutene polymerize. Reaction accelerators are not included. WO 2007/042516 A1 discloses a process for the preparation of relatively high molecular weight polyisobutene, in which a liquid mixture of isobutene and an inert diluent such as an aliphatic saturated or unsaturated C 2 - to C 6 -hydrocarbon in the gas space of a polymerization apparatus with the polymerization catalyst brings in contact. The mixture of isobutene and the inert diluent is introduced as a liquid jet, liquid film or in the form of droplets in the gas space of the reactor by means of a spray nozzle or spray disk.

WO 2007/1 13196 A1 describes a process for the preparation of relatively high molecular weight isobutene homopolymers in which isobutene in the liquid phase in the presence of Lewis acids as polymerization catalysts and an inert solvent such as ethene in an apparatus which simultaneously has mixing and conveying properties , polymerized. As such a polymerization apparatus, in particular a single-shaft or multi-shaft kneading reactor is described. International Patent Application PCT / EP2013 / 053042 (disclosure No. WO

2013/120984) relates to a process for the preparation of isobutene homopolymers by polymerization of isobutene in the liquid phase at -80 ° C to -190 ° C in an inert solvent in the presence of a polymerization catalyst based on Lewis acids and with concomitant use of reaction accelerators and / or chain length regulators.

The object specified for the present invention has been achieved by a process for the preparation of isobutene homopolymers having a weight-average molecular weight of 75,000 to 10,000,000 by polymerization of isobutene in a liquid phase polymerization apparatus

(a) at temperatures from -80 ° C to -190 ° C,

(b) in a suitable inert solvent which is liquid under the polymerization conditions,

(c) in the presence of a Lewis acid-co-catalyst polymerization catalyst complex characterized in that the isobutene in a mixing device comprising a mixing chamber and feeds for the partial streams, with the inert solvent, the Lewis Mixed acid catalyst and optionally other polymerization-controlling additive substances, this mixture in the form of a liquid jet or droplets in the liquid phase space of the polymerization fed and there can run the polymerization of the isobutene, wherein the mixing device designed so that the substances used in at least three partial streams of the mixing chamber are fed coaxially, wherein the isobutene-containing partial streams and the Lewis acid catalyst-containing partial streams are separated from each other by the inert solvent substreams entering the mixing chamber, so that no reactive Ge has mixed contact with the walls of the mixing device.

The inert solvent according to measure (b) is referred to as fluid, i. as the liquid carrier medium, in which either isobutene dissolves homogeneously in the liquid phase or in dispersed form because it is immiscible with the isobutene.

The mixing device used according to the invention can also be referred to as a mixing nozzle. It has a mixing chamber and feeds for the partial streams into the mixing chamber, which are usually mounted inside the mixing device, and - depending on the structural design - more customary for this purpose components. In particular, in the inner wall region of the mixing chamber there is a risk of the occurrence of caking and clogging. The isobutene-containing substreams and the catalyst-containing substreams are still completely separated from one another when they enter the mixing chamber due to the inert solvent. They do not intermix in the mixing chamber until after a certain distance and are completely free from the reactive mixture until they exit the mixing chamber or outside the mixing chamber, ie in the liquid phase space of the polymerization, mixed.

In a preferred embodiment, Lewis acid catalyst and co-catalyst are metered into the mixing chamber in different, separate partial streams, so that the polymerization catalyst complex only forms when these partial streams are mixed. In the present invention, structural and fluidic measures in the mixing device used ensure in particular that no reactive mixture comes into contact with its walls, in particular the walls of the mixing chamber, and can deposit polymer thereon. By suitable combination of the partial streams in the mixing device used and optionally additionally by mixing-active devices and / or internals (eg swirl generators that generate upstream in the individual streams turbulence) to achieve a sufficiently homogeneous mixing of the reactants in the polymerization, very short residence times of the reactive mixture in the mixing chamber and a defined liquid jet or film or a defined and controllable division of the reactive mixture into droplets.

For the supply of the isobutene, the coaxial feed is best suited as the most internal partial flow into the mixing chamber, since in the case of non-excludable malfunctions, however, the reactive mixture reaching the nozzle wall would have only a very low isobutene concentration, so that only a little unwanted polymer could deposit there. In a preferred embodiment, the mixing device is designed so that the substances used are supplied in three partial streams of the mixing chamber, wherein the coaxial internal partial flow of isobutene, co-catalysts and optionally further polymerization-controlling additives or from isobutene, co-catalysts, inert solvent and optionally further polymerization-controlling additives, the coaxial middle Teil- Ström of inert solvent and optionally other polymerization-controlling additives and the coaxial outer partial flow of Lewis acid catalyst, inert solvent and optionally other polymerization-controlling additives.

In a further embodiment, the mixing device is designed such that the substances used are supplied to the mixing chamber in three partial streams, wherein

the coaxial internal partial stream of isobutene, co-catalysts and optionally further polymerization-controlling additives or of isobutene, co-catalysts, inert solvent and optionally further polymerization-controlling additives,

- The coaxial mean partial flow of Lewis acid catalyst, inert solvent and, where appropriate, other polymerization-controlling additives and

- The coaxial outer partial flow of inert solvent and optionally other polymerization-controlling additives.

The attached FIG. 1 (according to the embodiment with three partial flows) is intended to illustrate schematically these embodiments of the mixing device without limiting it. The attached numbering here has the following meanings: 1 = mixing chamber outlet

2 = mixing chamber wall

3 = inlet of the internal partial flow

4 = inlet of the middle partial flow

5 = entry of the external partial flow

6 = partitions of the partial flow feeders

7 = fluid zone in which the pure internal partial flow is present

8 = fluid zone in which the pure mean partial flow is present

9 = fluid zone in which the pure external partial flow is present

10 = fluid zone in which there is a mixture of middle and outer partial flow

1 1 = fluid zone in which a mixture of middle and internal partial flow is present

12 = fluid zone in which there is a mixture of all three partial streams

13 = inlet cross sections of the three partial streams into the mixing chamber The distance between the inlet cross sections of the three partial streams into the mixing chamber (13) and the mixing chamber outlet (1) is expediently chosen such that the assumed region in which all three partial streams are mixed together (12) , not reached the mixing chamber wall (2). Further, since the inner and the outer partial flow are separated from each other in the region of the respective inlet cross section into the mixing chamber (13) by the inert central Teil-, there is also at the mixing chamber ends of the feed partitions (6) no reactive mixture. Thus, there may be no polymerization of the isobutene on the nozzle walls and it is an early blockage avoided. In a further preferred embodiment, the mixing device is designed so that the substances used are fed into four partial streams of the mixing chamber, wherein the coaxial internal partial flow of isobutene, co-catalysts and optionally further polymerization-controlling additives or of isobutene, co-catalysts, inert Solvent and optionally further polymerization-controlling additives, the coaxially counted from the inside first average substream of inert solvent and optionally other polymerization-controlling additives coaxial from the inside counted second average part of Lewis acid catalyst, inert solvent and optionally other polymerization-controlling additives and the external partial flow in turn, consists of inert solvent and optionally further polymerization-controlling additives.

The attached FIG. 2 (according to the embodiment with four partial flows) is intended to illustrate schematically these embodiments of the mixing device without limiting it. The attached numbers have the following meanings here: 1 = mixing chamber outlet

2 = mixing chamber wall

3 = inlet of the internal partial flow 4 = inlet of the first middle partial flow counted from the inside

5 = inlet of the second middle partial flow counted from the inside

6 = entry of the external partial flow

7 = partitions of the partial flow feeders

8 = fluid zone in which the pure internal partial flow is present

9 = fluid zone, in which the pure counted from the inside first middle partial flow exists

10 = fluid zone in which the pure is counted from the inside second average partial flow

1 1 = fluid zone in which the pure external partial flow is present

12 = fluid zone in which a mixture of the first counted from the inside middle and

internal partial flow is present

13 = fluid zone in which there is a mixture of the two middle part streams

14 = fluid zone in which there is a mixture of the second and the outer substream counted from the inside

15 = fluid zone in which there is a mixture of the two middle part streams and the inner part stream

16 = fluid zone in which there is a mixture of all partial streams

17 = inlet cross sections of the four partial streams into the mixing chamber

In this embodiment, with four partial streams is particularly advantageous that in the immediate vicinity of the nozzle wall no polymerization catalyst complex is present and thus the

Likelihood of polymer deposition is further reduced. In addition, with this embodiment, an even more homogeneous mixture of isobutene-containing and catalyst-containing substream can be produced, since the distance between the inlet cross sections of the partial streams into the mixing chamber (17) and the mixing chamber outlet (1) in comparison to the above-described embodiment with four partial streams can be chosen even larger and thus a longer mixing distance is available.

It has been found to be particularly advantageous that the mixing device is designed so that the flow cross-sections of the individual partial flows completely envelop the further inner partial streams. For this purpose, from a structural point of view, it is further preferred that the mixing device is designed such that the flow cross sections of the individual partial flows are aligned axially symmetrically, in particular circular, elliptical, oval, square, rectangular or polygonal, for example triangular, quadrangular, pentagonal or hexagonal.

The resulting gap-shaped channel cross sections in the case of rectangular or oval flow cross-sections have the advantage that the mixing path lengths are smaller than in the case of circular arrangements. This results in lower mixing times and thus the length of the mixing chamber and the residence time of the reactive mixture can be shortened here with a rather mixing effect. This in turn reduces the blockage tendency of the mixing nozzle. In a preferred embodiment, all the individual partial streams of the mixing device flow completely or almost completely isokinetically, ie with the same or almost the same speed. The average residence time of the isobutene fed into the mixing chamber with the corresponding partial flows, ie the time between entry from the feed channels and leaving (partially already polymerized) at the mixing chamber exit, is usually less than 10 seconds, in particular less than 1 second all less than 0.1 seconds. A typical residence time for the isobutene is 0.001 to less than 0.1 second.

As a rule, the spatial concentration fluctuations of the reactive components in the stream at the mixing chamber outlet amount to less than 40%, preferably less than 20%, in particular less than 10%, especially less than 5%, of the calculated theoretical concentration average of the respective component.

It has proved to be advantageous if the partial streams to be fed into the mixing device are cooled to temperatures in the region of the polymerization temperature. Preferably, it is cooled to a temperature which deviates absolutely no more than 20 ° C and in particular not more than 10 ° C from the reaction temperature in the polymerization, with temperatures below the reaction temperature are generally not critical, unless they lead to a separation of the Reactants and / or lead to the solidification of individual components.

Preferably, the mixing device is designed such that the feeds of the partial flows into the mixing chamber, which are usually mounted inside the mixing device, contain devices and / or internals for generating flow turbulences.

In particular, these devices and / or internals for generating turbulence turbulence generate by turbulence. As a rule, these swirl-producing devices or internals serve to displace sub-streams which enter the mixing chamber adjacent to one another. Such an opposite twisting causes a particularly strong shear at the junction of the partial flows, which entails an increased flow turbulence, which in turn accelerates the mixing process. This in turn can reduce the required residence time in the mixing chamber and thus reduce the tendency to clog. In this case, the swirl can be induced in the feed channels of the partial flows, in particular by means of moving or preferably non-moving devices or internals whose construction and design correspond to the usual designs known to the person skilled in the art. From a fluidic point of view, it is preferred that the flow rate of the outermost sub-flow is equal to or higher than that of all other sub-streams. If the other internal partial flows have different flow speeds, the flow velocity of the outermost substream should be equal to or higher than that of the fastest flowing substream. This avoids that there is a backflow near the wall, which can possibly transport polymerizable mixture to the wall. Thus, isobutene is less likely to reach the outer wall of the mixing chamber, and again this reduces the tendency to clog.

The above-described return flow can also be avoided by keeping the mean flow velocity of the total flow in the mixing chamber along the main flow direction constant or increasing it with the progression of the flow. The constant maintenance or the increase in the flow rate can be achieved by suitable technical measures familiar to the person skilled in the art. An increase in the flow rate is achieved for example by a successive reduction of the flow cross-section in the mixing chamber or - at a constant flow cross-section - by increasing the formation of gas in the total flow or by heating and thus density reduction of the gas-containing total flow. By increasing the flow velocity, a pressure gradient decreasing along the main flow direction is generated, which effectively prevents backflow. The actual polymerization of the isobutene is carried out in a suitable polymerization apparatus equipped with conventional means for removing or discharging the product and means for cooling. For example, the polymerization device may be a belt reactor. However, the polymerization apparatus is preferably selected from tubular reactors, loop reactors and reactors which simultaneously have mixing and conveying properties.

The feed of the mixture of isobutene, inert solvent, Lewis acid catalyst, co-catalyst and optionally further polymerization-controlling additives from the mixing device into the polymerization is generally advantageously at the top, at the bottom or on the side of the polymerization. By "upper end" is meant to mean the highest point or, in the case of a horizontal construction, the point furthest from the outlet of the polymerisation device. By \ "lower end \", the lowest point or horizontal shape is also meant to be the most remote point of the polymerizer from the outlet. If the feed is on the side, a position in the middle region of the polymerisation device is usually selected.

In order to better control the polymerization process, it may be useful to feed the reactive mixture not only at a single point, ie by means of a single mixing device in the liquid phase space of the polymerization, but the feed by several, in particular 2, 3, 4, 5 or 6 mixing devices, which on various locations of the polymerization are arranged under geometrically optimized aspects to make.

Depending on the nature of the inert solvent used and the molecular weight of the isobutene homopolymer obtained, the polymerization, which is generally carried out continuously or semi-continuously, takes place in a liquid homogeneous phase, ie. as solution or precipitation polymerization, or in liquid heterogeneous phase, i. as suspension polymerization, instead. Solution polymerization occurs at the beginning of the polymerization when the monomer isobutylene is well soluble in the inert solvent to be selected, and then successively undergoes a precipitation polymerization when insoluble high molecular weight polymer has formed. However, depending on the choice of inert solvent and the molecular weight of the resulting polymer, there is also the possibility that it will remain dissolved more or less completely dissolved in this solvent. In the case of heterogeneous-phase polymerization, monomer droplets may initially form in an inert solvent that is immiscible with the isobutene, which then convert to largely solid polymer particles in the further course of the polymerization reaction.

As a rule, in the case of the abovementioned polymerization apparatuses, the adjustment and constant maintenance of the polymerization temperature are carried out by direct or indirect cooling of the inert solvent used as the carrier liquid. Direct cooling is to be understood as meaning, in particular, cooling through the walls of the polymerization apparatus or cooling by boiling of the evaporating inert solvent. Indirect cooling is understood in particular as meaning the cooling via an external heat exchanger, such as a heat exchanger. Preferably, the setting and keeping constant the polymerization temperature in a cooling circuit via an external heat exchanger with or without bypass heat exchanger, by evaporative cooling, in particular by the evaporating inert solvent, or by wall cooling of the polymerization. It is advisable to ensure that all walls and components in contact with the polymerization medium have a temperature below the glass transition temperature of the polymer produced. In the embodiment of temperature regulation by means of an external heat exchanger, temperature variations in the fluid phase are possible to a limited extent. These temperature variations are generally less than 60 ° K, preferably less than 50 ° K, in particular less than 40 ° K, especially less than 30 ° K. In one embodiment of such a cooling circuit, the temperature regulation of the fluid phase takes place by means of a bypass heat exchanger, which is described below.

A suitable for the present invention tubular reactor usually has a channel-shaped structure, which is generally cylindrical, but also other cross-sectional shapes such as squares, rectangles, polygons, ovals or ellipses are possible. It often has the shape of a column and is then also referred to as a column reactor. A column reactor has a vertical arrangement. In a tube reactor, the height or length of the reaction space is usually greater than its diameter. A technically common version The form of such a polymerization device is, for example, a vertically arranged spray reactor, in particular a spray tower.

The height (length) to diameter ratio of the channel of the tubular reactor is preferably H / D = 0.5 to H / D = 100, in particular H / D = 1 to H / D = 50, particularly preferred

H / D = 2 to H / D = 30, especially H / D = 5 to H / D = 20. The diameter of the channel is to be understood as a hydraulic diameter. The height or length of the reactor is chosen so that the residence time of the reaction mixture in the reactor is generally 0.02 to 20 minutes, preferably 0.05 to 15 minutes, more preferably 0.1 to 10 minutes. The average flow rate of the reaction medium is usually 0.02 to 20 m / s, preferably 0.04 to 10 m / s, particularly preferably 0.06 to 5 m / s. In one embodiment, the tubular reactor H = 10 m long and has a diameter of D = 0.65 m and the residence time is 2 min. The channel of the tubular reactor can be designed with and without internals. As internals come, for example, a flow tube into consideration. Preferably, the channel is executed without internals. The walls of the channel are usually isolated by vacuum or suitable for the selected temperature range insulating material and / or tempered by means of a flowing in a double jacket fluid.

The removal of the heat of reaction from the tubular reactor is usually carried out by direct cooling of the reactor walls and / or by an external heat exchanger. Preferably, the heat of reaction is dissipated predominantly by an external heat exchanger. For this purpose, the inert solvent used as the heat transfer medium is separated from the polymer produced at the reactor outlet and cooled via an external circulation with a heat exchanger and returned to the reactor. The channel of the tube reactor is before the start of polymerization with the carrier medium, i. the inert solvent, and the reactants are either miscible with the carrier medium, whereby the polymer precipitates, or they form a second phase.

The polymerization takes place substantially along the reactor length of the tubular reactor, whereby the polymer produced can be moved either downwards or upwards by a possible difference in density between the inert solvent as the carrier medium and the polymer. It is also possible that the tube reactor is arranged horizontally and not vertically and the tube is not completely filled with carrier medium. The polymer in this case can "float" or settle on the carrier medium.

An advantage of the described cooling is that it reduces the risk of fouling, which could occur with a cooled double jacket. The carrier medium is circulated and thereby cooler carrier medium is available directly at the reactor inlet to absorb the heat of polymerization. A tubular reactor or column reactor with double jacket cooling is also suitable for the present invention. In this case, the carrier medium, ie the inert solvent which leaves the reactor with the polymer produced and is then separated therefrom, is not returned indirectly via a heat exchanger but directly to the mixer at the inlet of the reactor. However, a large part of the carrier medium remains in the reactor and cools again by contact with the reactor wall. The temperature of the carrier medium is adjusted by adjusting the temperature of the external cooling medium in the double jacket.

A loop reactor suitable for the present invention is, in principle, a vertical tubular reactor as described above, which does not operate with an external but with an internal circulation of the continuous phase. In this case, the reactor column usually consists of two concentrically arranged tubes, wherein the polymerization (when operating from bottom to top) in the inner column ascending, the polymer produced is separated at the top and serving as a carrier medium inert solvent in the annular gap between the inner and the outer column is conveyed back down. Here, the temperature setting is also carried out by an external cooling medium via the double jacket on the reactor. The continuous medium is only replaced to the extent that it is lost at the discharge. In addition, the continuous phase remains largely static in the reactor.

A suitable reactor for the present invention, which simultaneously has mixing and conveying properties, is referred to in specific structural embodiments as a kneading reactor or extruder reactor. As such a reactor can also be a self-cleaning apparatus with conveying and mixing action, as described in WO2007 / 1 13196 A1, are used. In a reactor which simultaneously has mixing and conveying properties, it may in particular be a double- or multi-shaft kneader or extruder. Preferably, a twin-screw kneader or a twin-screw extruder is used, particularly preferably a co-rotating twin-screw extruder. In the case of the twin-screw extruder, this can be made up of different types of screw elements. The residence time of the reaction mixture in the mentioned reactors, which simultaneously have mixing and conveying properties, is generally from 0.02 to 20 minutes, preferably from 0.05 to 15 minutes, more preferably from 0.1 to 10 minutes.

The design of such a reactor, which at the same time has mixing and conveying properties, takes place in such a way that it causes only a small specific power input, so that the reaction mixture scarcely heats up as a result of the dissipated mechanical power. The ratio of power input of the reactor to the heat of reaction liberated per time is less than 10, preferably less than 5, more preferably less than 2, in particular less than 1.

Such a reactor, which at the same time has mixing and conveying properties, can be inclined downwards in the conveying direction with respect to the horizontal, in order to ensure backmixing of the reaction mixture. to avoid contact in the inlet area and thus prevent clogging of the inlet. The angle of inclination has a value of 0 to 60 °, preferably 0 to 30 °, particularly preferably 0 to 20 °. The supply of the reactants and serving as the heat transfer medium inert solvent via the described blockage-free three- or multi-flow mixing nozzle or optionally on positively conveying conveying members, which also counteract a blockage of the inlet. The reactor can be isolated and / or tempered again. The mixture of polymer and heat transfer medium is separated at the end of the reactor and the heat transfer medium in turn is returned via an external heat exchanger to the inlet of the reactor.

In principle, any heat exchanger known to the person skilled in the art can be used as the external heat exchanger suitable for the reactor types mentioned, for example plate heat exchangers, tube bundle heat exchangers or static mixers designed as heat exchangers. Preferably, a bypass concept is used, in which the heated heat transfer medium coming from the reactor is partially recooled in a recycle loop well below the desired reactor inlet temperature and the recycle is then mixed with the heated heat transfer medium from the reactor. The advantage here is that the heated heat transfer medium is not directly in contact with a cold wall, but is cooled by mixing with the recirculated, supercooled stream. The partially dissolved polymer in the heat transfer medium does not precipitate as a solid on the wall during cooling, which would lead to a formation of deposits on the heat exchanger surface and thus to a deterioration of the heat transfer. On the contrary, when the supercooled stream is mixed with the heated heat medium, polymer particles are formed, which then flow with the cooled heat transfer medium and leave the heat exchanger again without forming a polymer film on the walls.

In the context of the present invention, isobutene homopolymers are understood to mean those polymers which, based on the polymer, are composed of at least 98 mol%, preferably at least 99 mol%, of isobutene.

For the use of the isobutene or of the isobutene-containing monomer mixture as the monomer to be polymerized, the isobutene source is in particular pure isobutene, which generally contains at most 0.5% by volume of residual impurities such as 1-butene, 2 Butenes, water and / or C 1 to C 4 alkanols. In principle, however, it is also possible to use isobutene-containing technical C4 hydrocarbon streams, for example C4 raffinates, C4 cuts from isobutane dehydrogenation, C4 cuts from steam crackers and FCC crackers (fluid catalysed cracking), insofar as they are largely derived from 1, 3-butadiene contained therein are liberated. Suitable technical C4 hydrocarbon streams generally contain less than 500 ppm, preferably less than 200 ppm, butadiene. The isobutene from such technical C4 hydrocarbon streams polymerized here largely selectively to the desired isobutene homopolymer, without appreciable amounts of other C4 monomers in the Polymer chain to be installed. Typically, the isobutene concentration in said C4 hydrocarbon commercial streams is in the range of 40 to 60 weight percent. However, in principle the process according to the invention can also be operated with isobutene-containing C 4 hydrocarbon streams which contain less isobutene, for example only 10 to 20% by weight. The isobutene-containing monomer mixture may contain small amounts of contaminants such as water, carboxylic acids or mineral acids, without resulting in critical yield or selectivity losses. It is expedient to avoid an accumulation of these impurities by removing such pollutants from the isobutene-containing monomer mixture, for example by adsorption on solid adsorbents such as activated carbon, molecular sieves or ion exchangers.

In accordance with measure (a), the polymerization process according to the invention is carried out in the liquid polymerization medium at temperatures of -80.degree. C. to -190.degree. In a preferred embodiment, it is carried out at temperatures in the vicinity of the lower limit of the above-mentioned temperature range, namely at -130 ° C to -190 ° C, especially at less than -160 ° C to -185 ° C, in particular at -165 ° C to -180 ° C, in a typical procedure at -168 ° C to -173 ° C. In an alternative preferred embodiment, the process is carried out at temperatures of -100 ° C to -150 ° C, preferably at -105 ° C to -147 ° C, especially at -1 10 ° C to -140 ° C, in particular at -1 15 ° C to -135 ° C, in a typical procedure at -120 ° C to -130 ° C, by. The controlled low polymerization temperatures have an advantageous effect on the product properties.

The polymerization process according to the invention is generally carried out at an absolute pressure of 500 mbar to 5 bar, in particular at a pressure of 800 mbar to 2 bar. Most advantageously and most economically, the polymerization reactor is operated at or near the ambient pressure (normal pressure). A slight overpressure may provide benefits for some of the possible inert solvents. Although a procedure of polymerization with overpressure is possible in principle, bring higher pressures, especially those above 5 bar, in general, no additional advantages.

According to measure (b), certain inert solvents or mixtures of such inert solvents are used in the liquid polymerization medium. In this case, inert solvents are understood to mean not only fluids in which isobutene dissolves homogeneously in the liquid phase, but also fluids with which isobutene is immiscible and is present in dispersed form. Suitable inert solvents are suitable for a C 1 to C 8 hydrocarbons, preferably C 1 to C 8 hydrocarbons, in particular C 2 to C 4 hydrocarbons, which are usually saturated or simply ethylenically unsaturated and generally have a linear or slightly branched structure , Of course, if they are ethylenically unsaturated, they must not polymerize themselves under the reaction conditions of the present invention; they usually have only primary and / or secondary olefinic carbon atoms. Typical examples of such C to Cs-hydrocarbons are methane, ethane, ethene, propane, propene, n-butane, isobutane, n-pentane, 2-methylbutane, 2,3-dimethylbutane, 2- Methylpentane, 3-methylpentane, 3-ethylpentane, 2,2-dimethylpentane, 2,3-dimethylpentane, 2,4-dimethylpentane, 2-methylhexane, 3-methylhexane, 3-ethyl-2-methylpentane, 2,2-dimethylhexane, 2,3-dimethylhexane, 3,3-dimethylhexane, 4-methylheptane, 2,2,3-trimethylpentane and 3-methylheptane. As such inert solvents, on the other hand, halogenated C 1 to C 8 hydrocarbons, preferably halogenated C 1 to C 8 hydrocarbons, in particular fluorinated and / or chlorinated C 1 to C 2 or C 1 to C 6 hydrocarbons, such as methyl chloride, methyl fluoride , Difluoromethane, dichloromethane, fluoroethane, 1-fluoropropane, 1, 1, 1, 2,3,3,3-heptafluoropropane, octafluoropropane or 1-fluorobutane, suitable; In particular perfluorinated d- to Cs- or d- to Cs-hydrocarbons or such d- to Cs- or d- to Cs-hydrocarbons come here, in which at least half of the hydrogen atoms are replaced by fluorine atoms into consideration. It is also possible to use mixtures of C.sub.1- to C.sub.s- or C.sub.1- to C.sub.s-hydrocarbons, mixtures of halogenated C.sub.1- to C.sub.s- or Cs-hydrocarbons or mixtures of one or more C.sub.1- or C.sub.1- to C.sub.1- Hydrocarbons and one or more halogenated d- to Cs or C to Cs-hydrocarbons are used. Also suitable as inert solvents in accordance with measure (b) are siloxanes, for example disiloxane or halogenated siloxanes, such as octachlorotrisiloxane, hexachlorodisiloxane, hexafluorodisiloxane or 1,1,1-trichloro-2,2,2-trifluorodisiloxane.

In a preferred embodiment, the inert solvent is selected according to measure (b) from C 1 to C 8 hydrocarbons, halogenated C 1 to C 8 hydrocarbons and mixtures thereof. Typical examples of these are ethane, ethene, propane, propene, n-butane, isobutane, 1,1,1,3,3,3,3-heptafluoropropane, octafluoropropane or mixtures thereof.

The weight ratio of isobutene to the inert solvents according to measure (b) in the polymerization reactor is generally 1: 0.1 to 1:50, preferably 1: 0.1 to 1:40, especially 0.1: 1 to 1: 20, especially 1: 0.5 to 1:10.

The polymerization catalyst complexes based on Lewis acids and cocatalysts to be used as polymerization catalyst according to measure (c) are preferably those based on boron trifluoride, iron halides, aluminum trihalides or aluminum alkyl halides as Lewis acids or based on a Lewis acid in combination with organic Sulfonic acids, each together with suitable co-catalysts. Co-catalysts are to be understood here in particular as activators or moderators in the form of proton sources, donors and initiators.

The said iron halide, aluminum trihalide and aluminum alkyl halide complexes contain, in addition to the Lewis acid, a donor in the form of an organic compound having at least one ether function or one carboxylic ester function. The combination of Lewis acids, in particular of boron trifluoride, iron halides, aluminum trihalides or aluminum alkyl halides, with organic sulfonic acids as initiators contain at least one organic sulfonic acid of the general formula Z-SO3H in which Z is a C 1 -C 20 -alkyl radical, C 1 -C 20 Haloalkyl radical, Cs-Cs-cycloalkyl radical, C 6 -C 20 -aryl radical or a C7- C2o-aralkyl radical; a typical such organic sulfonic acid is methanesulfonic acid.

However, according to measure (c), a complex of boron trifluoride and a proton source is preferably used as the polymerization catalyst. As such proton sources, which play the role of an activator or moderator in the catalyst complex, especially ethers, in particular C 1 to C ₄ dialkyl ethers such as diethyl ether, and alcohols, in particular low molecular weight monohydric aliphatic alcohols are suitable. In a particularly preferred embodiment, the polymerization catalyst used is a complex of boron trifluoride and a Cr to C ₄ -alkanol, for example. Methanol, ethanol, n-propanol, isopropanol, n-butanol, isobutanol, sec-butanol, and tert-butanol, a. It can be used as proton source and mixtures of said ethers and / or alcohols. The proton source can be used in a substoichiometric, stoichiometric or superstoichiometric amount in relation to boron trifluoride. Typical molar ratios of proton source to boron trifluoride are in the range of 0.3: 1 to 3: 1, in particular 0.5: 1 to 2: 1, especially 0.7: 1 to 1, 3: 1 (in each case based on 1 Proton equivalent of the proton source). In a preferred embodiment, mixtures of primary and secondary C 1 - to C ₄ -alkanols can be used as cocatalyst, as described in WO 2013/120859, particularly preferably mixtures of methanol and isopropanol, mixtures of methanol and isobutanol, and mixtures of methanol and sec-butanol, very particularly preferably mixtures of methanol and isobutanol.

The amount of polymerization catalyst complex (c) to be used depends essentially on the type of catalyst and on the reaction conditions, in particular the reaction temperature and the desired molecular weight of the polymer. It can be determined by means of fewer random tests for the respective reaction system. In general, the polymerization catalyst complex in amounts of 0.0001 to 1 wt .-%, in particular 0.0005 to 0.5 wt .-%, especially 0.001 to 0.1 wt .-%, each based on the Lewis Acid content or boron trifluoride content in the catalyst complex and on isobutene used.

In a preferred embodiment, the polymerization process according to the invention is carried out in the presence of at least one of the polymerization-controlling additives

(D) at least one reaction accelerator in the form of an ethylenically saturated hydrocarbon compound which at least one oxygen atom and no abstractable

Contains proton, and / or

(e) at least one chain length regulator containing at least one tertiary olefinic carbon atom. Particularly preferably, the measures (d) and (e) are carried out simultaneously. A reaction accelerator according to measure (d) is a compound which, under the chosen polymerization conditions, influences and thus controls the catalytic activity of the boron trifluoride in the desired manner. Such reaction accelerators are saturated hydrocarbon compounds which contain at least one oxygen atom, preferably as an ether oxygen atom or as part of a carbonyl function. In a preferred embodiment, the polymerization is carried out as measure (d) in the presence of at least one reaction accelerator selected from ketones, aldehydes, ethers, acetals and hemiacetals. Usually, such reaction accelerators are low molecular weight compounds having 1 to 40, in particular 1 to 16, especially 1 to 8 carbon atoms; their structure can be open-chain or cyclic; they may be aliphatic, aromatic or heteroaromatic in nature.

Typical representatives of such reaction accelerators are ketones such as acetone, butanone, cyclohexanone, acetophenone or benzophenone, aldehydes such as formaldehyde, trioxane, paraformaldehyde, acetaldehyde, propionaldehyde, butyraldehyde, benzaldehyde, cyclohexylaldehyde or glyoxal, dialkyl ethers such as dimethyl ether, diethyl ether or di-n butyl ether, cyclic ethers such as tetrahydrofuran or dioxane, as well as acetals and hemiacetals, by reacting the above-mentioned ketones and aldehydes with alcohols such as methanol, ethanol, n-propanol, isopropanol, n-butanol, isobutanol, sec-butanol or tert-butanol are available. Very particular preference is given to formaldehyde as such a reaction accelerator.

The reaction accelerators mentioned can usually advantageously be combined with one or more medium molecular weight alcohols, in particular monohydric aliphatic, cycloaliphatic or araliphatic alcohols, especially C 1 to C 10 alcohols, eg. n-pentanol, n-hexanol, n-octanol, 2-ethylhexanol, n-decanol, 2-propylheptanol, cyclohexanol or benzyl alcohol. On the one hand, such medium molecular weight alcohols act as activators or moderators in the catalyst complex, similar to the low molecular weight alcohols used as proton source in accordance with measure (c), but usually with a weaker activating effect; on the other hand, they act as solvents for the reaction accelerators. If aldehydes or ketones are used as reaction accelerators, the abovementioned medium molecular weight alcohols as well as some of the low molecular weight alcohols mentioned can form with these acetals or hemiacetals or ketals (ketone acetals), which also act as reaction accelerators. If one uses formaldehyde as a reaction accelerator, a corresponding alcoholic solution, for. As formaldehyde in isobutanol, are used. If such medium molecular weight alcohols are used, their weight ratio to the reaction accelerator is generally 0.05: 1 to 15: 1, but preferably 0.1: 1 to 5: 1, in particular 0.5: 1 to 2.5: 1, especially 0.75: 1 to 1, 5: 1.

The reaction accelerator itself is normally used in amounts of from 0.0001 to 1% by weight, preferably from 0.0003 to 0.75% by weight, in particular from 0.0005 to 0.5% by weight, in particular from 0.001 to 0, 1 wt .-%, each based on isobutene used. A chain length regulator according to measure (e) normally represents an ethylenically unsaturated system and contains one or more tertiary olefinic carbon atoms, optionally in addition to one or more primary and / or secondary olefinic carbon atoms. Most such chain length regulators are mono- or polyethylenically unsaturated hydrocarbons having 5 to 30, in particular 5 to 20, especially 5 to 16 carbon atoms; their structure can be open-chain or cyclic. Typical representatives of such chain length regulators are isoprene (2-methyl-1,3-butadiene), 2-methyl-2-butene, diisobutene, triisobutene, tetraisobutene and 1-methylcyclohexene. In a particularly preferred embodiment, the polymerization is carried out as measure (e) in the presence of isoprene and / or diisobutene as chain length regulators. Diisobutene (isooctene) is usually understood to mean the mixture of isomers of 2,4,4-trimethyl-1-pentene and 2,4,4-trimethyl-2-pentene; Of course, the individually used isomers 2,4,4-trimethyl-1-pentene and 2,4,4-trimethyl-2-pentene also act as chain length regulators according to measure (e). The amount of chain length regulators used according to the invention allows the molecular weight of the isobutene homopolymers produced to be adjusted in a simple manner: the higher the amount of chain length regulators, the lower the molecular weight generally becomes. The chain length regulator usually controls the molecular weight by incorporating it sooner or later into the polymer chain thus leading to chain termination at this point.

The chain length regulator is normally used in amounts of 0.0001 to 2 wt .-%, in particular 0.0005 to 1 wt .-%, especially 0.001 to 0.5 wt .-%, each based on isobutene used.

The isobutene homopolymers prepared by the process according to the invention preferably have a weight-average molecular weight (M _w ) of from 150,000 to 8,000,000, in particular from 250,000 to 6,000,000, especially from 400,000 to 5,000,000. Alternatively, they preferably have a number average molecular weight (M _n ) (determined by gel permeation chromatography) of from 25,000 to 2,000,000, particularly preferably from 45,000 to 1,500,000, in particular from 55,000 to 1,000,000, especially from 65,000 to 750,000.

In general, the isobutene homopolymers produced by the process according to the invention have a polydispersity (PDI = M _w / M _n ) of from 2 to 20, in particular from 3 to 15, especially from 5 to 10.

It is advantageous to carry out the process according to the invention at high conversions, if possible at full conversion or near full conversion, for example at a conversion of the isobutene used to the desired product of 85% to 100%, in particular from 90% to 100%. However, it is also possible to carry out the process according to the invention with partial conversion, for example at a conversion of the isobutene used to the desired product of from 10% to 85%, in particular from 30% to 60%. In a preferred embodiment, the polymerization conditions for the process according to the invention are selected such that the Isobutene used in the polymerization reactor with a conversion of at least 90%, in particular of at least 95%, especially of at least 99%, is converted to isobutene homopolymers having a weight-average molecular weight of 75,000 to 10,000,000. The present invention also provides a process for the preparation of relatively high molecular weight isobutene homopolymers, as described above, in which the isobutene homopolymers produced are discharged from the polymerization apparatus at temperatures of less than -80 ° C, thereby completely or substantially separating them from the inert solvent and then optionally subjecting them to a thermal cleaning process at temperatures of more than + 80 ° C. Preferably, the complete or extensive separation of the produced isobutene homopolymers from the inert solvent is preferably carried out with the aid of sieves, filters, centrifuges, extruders, blades, settling tanks or floating tanks.

For example, due to its particle size, typically from about 0.1 mm to about 5 cm, usually from about 0.5 mm to about 10 mm, the polymer produced can be retained on a polytetrafluoroethylene-based sieve while the solvent passes through the mesh of the sieve penetrate and can be returned to the process. The polymer is subsequently recovered as a filter cake. Both the suspension feed and the filter cake removal is usually continuous or semi-continuous. It can also be a self-cleaning or regularly backwashed filter with automated changing device, eg. As a rotating screen changer, or a vacuum drum filter can be used.

When using centrifuges, the polymer should be separated from the continuous medium by centrifugal force sedimentation due to differences in density (depending on the solvent). The generation of the centrifugal force field and the separation are carried out in a rotating apparatus, wherein the phase with the higher density is more strongly detected by the centrifugal force and thus deflected more radially. Depending on the type of centrifuge, the separated phases are kept isolated from each other by various devices or conveyors and then discharged. It is important that the separation takes place in such a way that there is always some solvent left over for washing away or transporting away the polymer in order to prevent buildup of caking.

For the separation of polymer and carrier medium, it is also advantageously possible to use extruders (also called screw conveyors and screw pumps in special designs). Both single-shaft and multi-shaft extruders are suitable. In twin- and multi-shaft extruders, the waves can work in the same direction or in opposite directions. Two- or multi-screw extruders are preferably used, since they have an effective self-cleaning. Two-screw extruders, in particular co-rotating twin-screw extruders, are particularly preferably used. The waves in single and multi-shaft extruders are normally occupied by kneading and / or conveying elements. These devices usually promote self-cleaning. The speeds of rotation of the waves are generally in the range of 10 to 500, in particular special 15 to 350 revolutions per minute. In a special design, the waves may be formed as worm shafts, whose gears are engaged with each other and whose inner shaft diameter is preferably constant over the entire length. Preferred building materials for the described extruders are steels or stainless steels. In the case of an ascending polymerization, the mixture of polymer produced and carrier medium is discharged from the top of the polymerization apparatus and introduced directly into the extruder. In a descending polymerization from top to bottom, the mixture of polymer and carrier medium via a screw or similar device is fed into the extruder and the separation is then carried out analogously to the ascending polymerization.

Such an extruder serving for the separation of the polymer produced can be set up with an angle of attack of 0 to 20 °, in particular 5 ° to 15 °, especially 0 ° to 10 °. The separation takes place via the rotation of the screws. The polymer is moved in the extrusion direction and discharged at the top of the extruder. Due to the angle of attack and the resulting gravitational force as well as the lower viscosity compared to the polymer, the carrier medium is not detected as strongly as the polymer and discharged at the bottom against the conveying direction of the extruder. The carrier medium is collected and can be supplied to the external heat exchanger circuit. The angle of attack avoids that too much carrier medium is discharged with the polymer, which then has to be degassed in another extruder.

For the separation of polymer and carrier medium, it is also possible to use blades for the reactor types which are operated without the constant external exchange of solvent. In this case, it is possible to remove the polymer similar to a waterwheel when the polymer is rising to the surface of the continuous phase ("frame"). For this purpose, a construction based on a wheel or a bath can be used. The emptying of the individual blades can be done mechanically and self-cleaning.

For separation of polymer and carrier medium, settling tanks or floating pools can also be used. In this case, due to differences in density (depending on the solvent), the polymer produced is separated from the continuous medium by gravity sedimentation. Depending on the density ratio between polymer and continuous medium, the polymer is withdrawn as sediment in the settling tank or discharged as an overflow in the swimming pool.

A desired thermal cleaning after the discharge of the product from the polymerization reactor is advantageously carried out by using one or more extruders. For this purpose, the extruder can be used, which provide for the discharge of the product, if they are heated accordingly. In this case, the isobutene homopolymers are heated to temperatures of more than 80 ° C., in particular more than 100 ° C. By mechanical action of the extruder shafts and possibly of internals in the extruder, the inner surface for better degassing of the volatile components in the product such Residual monomers and solvents constantly renewed. The degassing and the purification of the product can be facilitated by applying a vacuum, in particular one works for this purpose at an absolute pressure of less than 700 mbar, in particular less than 200 mbar, especially less than 100 mbar. Advantageously, an inert gas, for example nitrogen, is introduced into one or more segments of the extruder in order to promote the degassing process.

The inventive method has the advantage that the isobutene homopolymers produced in the inert solvent used (usually hydrocarbons and / or halogenated hydrocarbons) have only a low solubility - this applies to a greater extent at low temperatures - and thus largely precipitate as a solid. This precipitated solid has no tendency to stick at the low temperatures used and under the specific process conditions of the present invention, so that the crude product can be easily discharged and processed, especially in the region of collection of the product from the reactor in the workup at any point Temperatures above the glass transition temperature of the polymer occur.

Claims

21

claims

1 . A process for producing isobutene homopolymers having a weight-average molecular weight of 75,000 to 10,000,000 by polymerization of isobutene in a liquid phase polymerization apparatus

(a) at temperatures from -80 ° C to -190 ° C,

(c) in the presence of a polymerization catalyst complex based on

Lewis acids and co-catalysts, characterized in that mixing the isobutene in a mixing device, comprising a mixing chamber and feeds for the partial streams, with the inert solvent, the Lewis acid catalyst and optionally further polymerization-controlling additives, this mixture in the form a liquid jet or droplets in the liquid phase space of the polymerization fed and there run the polymerization of isobutene, wherein the mixing device designed so that the substances used in at least three partial streams of the mixing chamber are fed coaxially, wherein the isobutene-containing partial streams and the Lewis Acid catalyst containing part streams are separated by the inert solvent existing part streams on entry into the mixing chamber, so that no reactive mixture has contact with the walls of the mixing device.

A method according to claim 1, characterized in that the mixing device is designed so that the substances used are supplied in three partial streams of the mixing chamber, wherein the coaxial internal partial flow of isobutene, co-catalysts and optionally other polymerization-controlling additives or of isobutene, co-catalysts , inert solvent and optionally further polymerization-controlling additives, the coaxial mean partial flow of inert solvent and optionally further polymerization-controlling additives and the coaxial outer partial flow of Lewis acid catalyst, inert solvent and optionally other polymerization-controlling additives.

A method according to claim 1, characterized in that the mixing device is designed so that the substances used are supplied in four partial streams of the mixing chamber, wherein the coaxial internal partial flow of isobutene, co-catalysts and optionally further polymerization-controlling additives or of isobutene, co-catalysts , inert solvent and, if appropriate, further polymerization-controlling additives, which are coaxially counted from the inside first average partial stream of inert solvent and optionally further polymerization-controlling additives, the co- 22

axially second from the inside counted second partial stream of Lewis acid catalyst, inert solvent and optionally further polymerization-controlling additives and the external partial stream again consists of inert solvent and optionally other polymerization-controlling additives.

4. The method according to claims 1 to 3, characterized in that the mixing device is designed so that the flow cross-sections of the individual partial flows completely envelop the respective further inner sub-streams. 5. Process according to claims 1 to 4, characterized in that all individual partial flows of the mixing device flow completely or almost completely isokinetically.

6. The method according to claims 1 to 5, characterized in that the polymerization device is selected from tubular reactors, loop reactors and reactors, which simultaneously have mixing and conveying properties.

7. The method according to claims 1 to 6, characterized in that the feed of the mixture of isobutene, inert solvent, Lewis acid catalyst, co-catalyst and optionally further polymerization-controlling additives from the mixing device in the polymerization at the top, at the bottom or at the

Side of the polymerization takes place.

8. The method according to claims 1 to 7, characterized in that the adjustment and keeping constant the polymerization temperature by cooling serving as a carrier liquid inert solvent in a cooling circuit via an external heat exchanger with or without bypass heat exchanger, by boiling cooling or by wall cooling of the polymerization ,

9. A process according to claims 1 to 8, wherein the polymerization is carried out at temperatures of -1 10 ° C to -140 ° C or less than -160 ° C to -185 ° C.

A process according to claims 1 to 9, wherein the inert solvent (b) is selected from C 1 to C 8 hydrocarbons, halogenated C 1 to C 8 hydrocarbons, and mixtures thereof.

1 1. Process according to Claims 1 to 10, wherein the polymerization catalyst complex (c) used is a Lewis acid complex based on boron trifluoride, iron halides, aluminum trihalides or aluminum alkyl halides or a Lewis acid in combination with organic sulfonic acids, in each case together with suitable cosolvents. Catalysts used. 23

12. The method according to claims 1 to 1 1, wherein the polymerization in the presence of at least one of the polymerization-controlling additives at least one reaction accelerator in the form of an ethylenically saturated hydrocarbon compound containing at least one oxygen atom and no abstractable proton, and / or

(e) at least one chain length regulator which contains at least one tertiary

contains olefinic carbon atom carried out.

13. The method according to claims 1 to 12, characterized in that one discharges the isobutene homopolymers produced from the polymerization at temperatures of less than -80 ° C and thereby completely or largely separated from the inert solvent and then optionally a thermal cleaning process at temperatures of more than + 80 ° C.

14. The method according to claim 13, characterized in that one carries out the complete or extensive separation of the isobutene homopolymers produced by the inert solvent with the aid of screens, filters, centrifuges, extruders, kneaders, blades, sedimentation o- and Aufschwimmbecken.

15. The method according to claim 13 or 14, characterized in that one carries out the complete or extensive separation of the isobutene homopolymers produced by the inert solvent in a relative to the horizontal inclined extruder or kneader, wherein the tougher polymer in the conveying direction of the extruder or kneader moved up while the less viscous inert solvent moves against the conveying direction of the extruder or kneader down.