Method for producing a polyisobutene

Abstract

The present invention relates to a method and a device for producing polyisobutene at low temperatures.

Description

[0001]BASF SE 150272 Process for the preparation of polyisobutene Description The present invention relates to a process and a device for the preparation of polyisobutene at low temperatures. The preparation of isobutene-containing homo- and copolymers with a number-average molecular weight Mn of 500 to 1,000,000 is carried out on an industrial scale under Lewis acid catalysis at temperatures usually well below room temperature. The reaction temperature can be reduced to as low as -90 °C. The higher the desired number-average molecular weight Mn of the polyisobutene, the lower the reaction temperature is generally. Thus, low-molecular-weight polyisobutene with a number-average molecular weight Mn of up to a few 1,000 can be prepared at temperatures around 0 °C, or only slightly below. For polyisobutene with a number-average molecular weight Mn of several 10,000, temperatures in the range of -10°C to -30°C are usually required.For the production of high-molecular-weight polyisobutene with a molecular weight of more than 100,000, however, reaction temperatures of no more than -50°C are required. EP 322241 A2 discloses a process for the production of polybutenes with a molecular weight Mn of 500 to 5000 by reacting isobutene in the presence of BF3 under specific reaction conditions. EP 628575 A1 and WO 2013 / 62763 describe various reactors and their operation for the production of highly reactive polyisobutenes. None of the cited documents addresses energy management during the reaction. US 6525149 B1 describes the conduct of ionic polymerizations in tube-bundle reactors, wherein the reaction mixture is circulated in a specific ratio to the monomer feed. The examples disclose specific flow rates and dimensions of the reaction zone, but heat management remains unclear.US 2019 / 0118158 A1 describes the conduct of ionic polymerizations, e.g., the polymerization of isobutene to polyisobutene, in a circulating heat exchanger reactor system (HERS), which comprises special parallelepiped-shaped channels. A disadvantage of such reactor systems is that they are complex to manufacture, since the side surfaces must be welded. Weld seams are, on the one hand, costly to produce and, on the other hand, are a source of leaks and corrosion. DE 102009019470 A1 discloses the cationic polymerization of monomers, wherein at least the mixing of the streams and preferably also the polymerization takes place in microstructures.The microstructures enable large heat exchanger surfaces, which means that the polymerization can be carried out almost isothermally. However, the microstructures have the disadvantage that they quickly become clogged by polymer deposits. Therefore, the reactions according to DE 102009019470 A1 are carried out exclusively in a straight pass without a circulation mode, since the circulating reaction mixture contains polymer, which would cause deposits or blockage of the microstructures. Due to the low reaction temperatures, these processes for the production of polyisobutene are very energy-intensive, particularly for the generation and maintenance of low temperatures and for the removal of the reaction heat at a low temperature level. Due to the low reaction temperature, the use of favorable cooling media, especially water, especially river water, for the removal of the reaction heat during the polymerization to polyisobutene is generally not possible.The task was therefore to provide a process and a device for the production of isobutene homo- and copolymers which reduces the energy expenditure when maintaining low temperatures and optimizes the energy efficiency of the device. An isothermal reaction process may be desirable from a chemical point of view, but is uneconomical. Although it is common practice to insulate cooled equipment and pipelines from the ambient temperature, this insulation cannot be complete: In order to ensure the operability and functionality of plant components that are movable or must be kept easily accessible, e.g. pumps or valves, the insulation is penetrated at numerous points, creating thermal bridges. The contents of the equipment, valves and pipelines are heated at these points via such thermal bridges, which on the one hand reduces energy efficiency.If the contents of the apparatus and piping are still reactive at these points due to the presence of non-deactivated Lewis acid, heating above the desired reaction temperature will further accelerate the reaction, which can lead to the formation of by-products or, in general, to a lack of uniformity in the product, since the reaction temperature, as explained above, has a strong influence on the reaction and particularly on the molar mass distribution, which leads to a high polydispersity of the product. This object was achieved by a process for the preparation of polyisobutenes with a number-average molecular weight Mn of 500 to 1,000,000 from isobutene or an isobutene-containing monomer mixture, in which the reaction is carried out in a reaction system whose ratio of cooled surface area Ocooling to reaction volume V satisfies the inequality O. Kühl / V ≥ 50 m 2 / m 3, preferably at least 55, particularly preferably at least 60, very particularly preferably at least 70 and in particular at least 80 m 2 / m 3 , and - for whose ratio of cooled surface OKool and uncooled surface ONcool the inequality O Kühl / O NKühl ≥ 5 [m 2 / m 2 ], preferably at least 7, particularly preferably at least 10 and most particularly preferably at least 12 [m 2 / m 2] applies. By these measures, energy inputs into the reaction mixture via the surface of the reaction system are kept as low as possible. A further advantage is that the reaction temperature need not be kept lower than required for the desired polymerization result. Otherwise, the polymerization energy released would have to be dissipated at a lower temperature level, which would reduce the efficiency of the refrigeration machine for generating the cooling medium. Definitions In the context of this document, the reaction system is understood to be the part of the plant in which isobutene or an isobutene-containing monomer mixture reacts to form polyisobutene with an increase in the number average molecular weight Mn. In the plant, this includes in particular reactors, heat exchangers, lines, pumps, valves, nozzles and flanges, preferably reactors, heat exchangers and lines.This is preferably the part of the plant in which the isobutene or an isobutene-containing monomer mixture is in contact with at least one Lewis acid. The reaction is generally terminated by deactivating the at least one Lewis acid, for example by adding water or alcohols, preferably water. Therefore, the plant components and apparatus with which the reaction mixture comes into contact after deactivation of the Lewis acid are no longer considered part of the reaction system. With particular preference, the reaction system is therefore understood to mean the part of the plant in which the isobutene or an isobutene-containing monomer mixture is in contact with a non-deactivated Lewis acid, so that the apparatus in which the deactivation takes place is no longer considered part of the reaction system.The reaction system therefore most preferably comprises the reactor and the heat exchangers, pumps, and lines leading from the reactor to the inlet of the apparatus in which the deactivation takes place. Another possibility for terminating the reaction is the removal of monomeric isobutene from the Lewis acid-containing reaction mixture, for example by distillation or extraction, preferably single- or multi-stage distillation. Complete conversion of the monomeric isobutene is also conceivable. In general, the reaction no longer takes place when there is no longer any contact between the reactive Lewis acid and the monomeric isobutene.The reaction system comprises at least one part, for example one to three, preferably one or two or particularly preferably one part, in which the reaction mixture is in direct contact via a thermally conductive surface with a colder medium than the reaction mixture, and at least one part, for example one to three, preferably one or two or particularly preferably one part, in which this is not the case. The reaction mixture is circulated via at least one cooled heat exchanger. Fresh isobutene-containing feed is added to this circulating reaction mixture, and the appropriate amount of reaction product is withdrawn in order to keep the total amount of reaction mixture constant.To homogenize both the reaction mixture and the heat of reaction during the reaction, the volume of the circulating reaction mixture should preferably be at least 15 times the fresh isobutene-containing feed, preferably at least 20 times, and particularly preferably at least 25 times. 100 times the volume of the fresh isobutene-containing feed should not be exceeded, since otherwise the energy required for the pumping power would make the overall process uneconomical; preferably, 80 times, particularly preferably 70 times, most preferably 60 times, and in particular 50 times should not be exceeded. BASF SE 150272 5 Furthermore, a high pumping power introduces energy into the system, which in turn must be removed from the reaction mixture by cooling. Pumping a high circulating volume therefore counteracts the optimization of energy efficiency.The fresh isobutene-containing feed is generally mixed with the reaction mixture by introducing energy through shear energy. This can be done, for example, in dynamic mixing apparatus, i.e. by mixing using a stirrer or by pumping (natural circulation or forced circulation) or pumping with static mixing devices such as static mixers or nozzles in the pumping circuit, by static mixing apparatus such as static mixers, nozzles, orifices, Y- or T-pieces, or by dynamic mixing apparatus such as mixing pumps. Mixing is preferably carried out via a mixing section with static mixers or nozzles, preferably static mixers. Heat exchangers are selected from the group consisting of temperature-controlled tubular reactors, shell-and-tube heat exchangers, plate heat exchangers, and temperature-controlled tubular reactors with internals, preferably shell-and-tube heat exchangers and plate heat exchangers, with particular preference being given to shell-and-tube heat exchangers.It is conceivable, although less preferred, to carry out the reaction in at least one stirred-tank reactor or a cascade of at least two stirred-tank reactors, preferably 2 to 5, particularly preferably 2 to 3, each of which is cooled via a jacket and / or internal heat exchanger. An advantage of the present invention is that it can be carried out in these conventional reactors and does not require complicated reactor systems according to US 2019 / 0118158 A1. The cooled surface O. Kühlis the area of ​​all plant components of the reaction system which is in direct contact with a colder medium than the reaction mixture in the plant components in question via a thermally conductive surface. The temperature difference between the colder medium and the reaction mixture is, for example, at least 5 K, preferably at least 8 K, more preferably at least 10 K and most preferably at least 12 K. A temperature difference between the colder medium and the reaction mixture which is greater than 30 K generally offers no advantage; preferably the temperature difference is not more than 25 K, more preferably not more than 20 K, most preferably not more than 18 K and in particular not more than 15 K. Large temperature differences cause supercooling of the reaction mixture close to the wall, so that polyisobutene with a higher molecular weight is formed there.Furthermore, the lower temperature locally increases the viscosity, making mixing more difficult and increasing the residence time on the reactor wall. This results in the average molecular weight of the polyisobutene becoming more uneven, i.e., increasing polydispersity. The colder medium can be, for example, brine, particularly glycol / water mixtures, or ammonia. It is also possible to use organic hydrocarbons such as propane, butane isomer mixtures or ethene, or partially or fully fluorinated hydrocarbons. The thermally conductive surfaces are, for example, the surfaces of the heat exchanger tubes or plates or casings of reactors or pipes that carry the colder medium. Ocool refers to the inside of the aforementioned surfaces, which is in direct thermal contact with the reaction mixture and any gas phase above it.The uncooled surface ONcool, on the other hand, is the area of ​​all system components in the reaction system that are in direct heat-conducting contact with a medium that has a higher temperature than the reaction mixture. Here, too, the inside of the mentioned surfaces is used. The medium that has a higher temperature than the reaction mixture is preferably the ambient atmosphere under the prevailing conditions. The respective surfaces are usually thermally insulated from the outside, which means they still have a higher temperature than the reaction mixture. Typical uncooled surfaces are pumps, fittings, valves and pipes, unless they are provided with a jacket (tracing cooling) that carries a colder medium. To distinguish from O. Kühl and O Nkühland to calculate the relevant surfaces, the points are used where the wall temperature is equal to or higher than the temperature of the reaction mixture. The reaction volume V is the total volume of the reaction system enclosed by plant components in which isobutene or an isobutene-containing monomer mixture reacts to increase the number-average molecular weight Mn. The preferred embodiments described above in the definition of the reaction system are to be applied accordingly here. BASF SE 150272 7 The volume is the volume enclosed by the surface defined above. With regard to the reaction volume, it may be useful to distinguish between lines and containers: In this document, lines are defined as machines and apparatus which have a surface area: volume ratio (S / V) of at least 4 m 2 / m 3preferably at least 8, particularly preferably at least, most preferably at least 16 and in particular at least 40 m 2 / m 3 These are especially true for pipelines with a nominal diameter (NW) of 100 mm to 1000 mm, preferably 300 to 500 mm. It is irrelevant whether the surface is cooled or uncooled. For example, a pipeline with a nominal diameter of 100 mm has an O / V ratio of approximately 40 m. 2 / m 3 on, NW 250 mm an S / V of 16 m 2 / m 3 and NW 1000 mm an S / V of 4 m 2 / m 3The O / V ratio depends solely on the radius r of the pipe: O / V = 2 / r. In this document, "vessels" are machines and devices that have a lower surface area to volume ratio than the pipes defined above. Assuming a vessel as an ideal cylinder with a circular, flat bottom and lid, typical O / V ratios for a radius r of 0.5 m and a height h = 1 m are O / V = 6 m. 2 / m 3 , at r = 1 m and h = 1 m, O / V = 4, at r = 0.5 m and h = 0.5 m, O / V = 8 and at r = 1 m and h = 5 m, O / V = 2.4 m 2 / m 3 . In general, for such an ideal cylinder, O / V = (2 ^h + 2 ^r) / (r ^h) applies. From these O / V ratios, it is clear that to achieve the inventive ratio of Ocool / V ≥ 50 m 2 / m 3The cooling of the pipes in particular has a strong influence, since these have a high O / V ratio according to the above definition, i.e. a large surface area in relation to their volume, over which heat exchange can take place. The ratio of O Kühl / V at least 55, particularly preferably at least 60, most preferably at least 70 and in particular at least 80 m 2 / m 3 . A ratio of O Kühl / V of 250 m 2 / m 3 is preferably not exceeded, particularly preferably the ratio of O Kühl / V not more than 200, most preferably not more than 175, especially not more than 150 m 2 / m 3 . High ratios O Kühl / V are essential for achieving an isothermal reaction, which is not required for optimizing energy efficiency according to the invention. According to the invention, the ratio of non-cooled surface to cooled surface must be at least 5 [m 2 / m 2 ], preferably at least 7, particularly preferably at least 10 and most preferably at least 12 [m 2 / m 2]. This makes it possible to limit the temperature rise in the uncooled reaction system to no more than 2.5 K, preferably no more than 2.2 K, more preferably no more than 2.0 K, and most preferably no more than 1.8 K compared to the inlet temperature into the uncooled reaction system. By limiting the temperature rise in this way, the reaction mixture is exposed to smaller temperature fluctuations that would lead to a non-uniform product mixture. However, it is not necessary to carry out the reaction isothermally, i.e., without a temperature rise. However, an isothermal reaction would require a large cooled area. It has been found that a temperature rise in the uncooled reaction system of 0.3 K or more is acceptable without the product mixture becoming excessively non-uniform, preferably 0.5 K or more, more preferably 0.75 K or more, most preferably 1.0 K or more, and in particular 1.5 K or more. The present invention further provides a production plant for producing polyisobutene, comprising at least the following plant components: - at least one apparatus for providing the isobutene-containing monomer mixture, selected from the group consisting of a line, a tank, and a buffer vessel, - at least two storage vessels for at least one Lewis acid and at least one cocatalyst, - a metering device for the at least one Lewis acid and at least one cocatalyst into the isobutene-containing monomer mixture, - a refrigeration system, - at least one reaction system for converting the isobutene-containing monomer mixture, comprising at least one reactor, heat exchanger, and lines, - at least one vessel for terminating the reaction by metering in water or alcohol, and - at least one distillation apparatus for working up the reaction mixture,in the BASF SE 150272 9 in the reaction system the ratio of cooled surface OKool and reaction volume V the inequality OKool / V ≥ 50 m, 2 / m 3 , preferably at least 55, particularly preferably at least 60, very particularly preferably at least 70 and in particular at least 80 m 2 / m 3 , and - for whose ratio of cooled surface OKool and uncooled surface ONcool the inequality OKool / ONcool ≥ 5 [m 2 / m 2 ], preferably at least 7, particularly preferably at least 10 and most particularly preferably at least 12 [m 2 / m 2] applies. The statements made for the process also apply to the production plant according to the invention. The present invention further relates to the use of such a production plant for producing polyisobutene having a number-average molecular weight Mn of 500 to 1,000,000 from isobutene or an isobutene-containing monomer mixture. The present invention further relates to a process for producing polyisobutene from an isobutene-containing monomer mixture in the presence of at least one Lewis acid in such a production plant. In the context of this document, polyisobutene is understood to mean homopolymers and copolymers containing isobutene in polymerized form.For the production of such homopolymers or copolymers containing isobutene in polymerized form, suitable isobutene sources include, in one preferred embodiment, both pure isobutene and, in another preferred embodiment, isobutene-containing C4 hydrocarbon streams, for example, C4 raffinates, in particular "raffinate 1," C4 cuts from isobutane dehydrogenation, C4 cuts from steam crackers, and from FCC crackers (fluid-catalyzed cracking), provided they are largely freed of the 1,3-butadiene present therein. A C4 hydrocarbon stream from an FCC refinery unit is also known as a "b / b" stream. Other suitable isobutene-containing C4 hydrocarbon streams include, for example, the product stream from a propylene-isobutane cooxidation or the product stream from a metathesis unit, which are generally used after conventional purification and / or concentration.Suitable C4 hydrocarbon streams generally contain less than 500 ppm, preferably less than 200 ppm, of butadiene. The presence of 1-butene and cis- and trans-2-butene is largely uncritical. Typically, the isobutene concentration in the C4 hydrocarbon streams mentioned in BASF SE 150272 10 is in the range of 40 to 60 wt. Thus, raffinate 1 generally consists essentially of 30 to 50 wt.% isobutene, 10 to 50 wt.% 1-butene, 10 to 40 wt.% cis- and trans-2-butene, and 2 to 35 wt.% butanes; In the polymerization process according to the invention, the unbranched butenes in the raffinate 1 generally behave virtually inertly and only the isobutene is polymerized. In a preferred embodiment, a C4 hydrocarbon stream having an isobutene content of 1 to 100% by weight, in particular 1 to 99% by weight, especially 1 to 90% by weight, particularly preferably 30 to 60% by weight, is used as the monomer source for the polymerization.-%, in particular a raffinate 1 stream, a b / b stream from an FCC refinery unit, a product stream from a propylene-isobutane co-oxidation or a product stream from a metathesis unit. The isobutene content in the C4 hydrocarbon stream is preferably at least 40, more preferably at least 45 wt. In particular when using a raffinate 1 stream as the isobutene source, the use of water as the sole or as a further initiator has proven useful, especially when polymerizing at temperatures from -20°C to +30°C, in particular from 0°C to +20°C. At temperatures from -20°C to +30°C, in particular from 0°C to +20°C, however, the use of an initiator can also be dispensed with when using a raffinate 1 stream as the isobutene source. The isobutene-containing monomer mixture mentioned may contain small amounts of contaminants such as water, carboxylic acids or mineral acids without causing critical losses in yield or selectivity.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. The proportion of comonomers other than isobutene in the polyisobutene is generally not more than 10% by weight, preferably not more than 7.5% by weight, more preferably not more than 5% by weight, most preferably not more than 2.5% by weight, and in particular not more than 1% by weight. The process according to the invention can be used to prepare polyisobutenes having a number-average molecular weight Mn of 500 to 1,000,000, preferably 500 to 250,000, more preferably 550 to 200,000, most preferably 650 to 100,000, in particular 750 to 50,000 g / mol.BASF SE 150272 11 The polydispersity, i.e. the quotient of weight-average molecular weight Mw and number-average molecular weight Mn (PDI = Mw / Mn), is from 1.05 to 3.5, preferably from 1.05 to 2.5, in particular from 1.05 to 2.0, and especially from 1.1 to 1.85. Typical values ​​for PDI under optimal process conditions are from 1.2 to 1.7. The molecular weight is determined by gel permeation chromatography using polystyrene as the standard. Among these polyisobutenes, those having a high content of terminally arranged ethylenic double bonds (α-double bonds) are preferred, particularly those having a content of α-double bonds of at least 50 mol%, preferably at least 60 mol%, particularly preferably at least 70 mol%, very particularly preferably at least 80 mol% and in particular at least 90 mol%.These are referred to as highly reactive polyisobutenes, and the formation of the terminal ethylenically unsaturated double bonds is particularly sensitive to variations in reaction temperature. The process according to the invention is operated at various reaction temperatures depending on the desired molecular weight of the polyisobutene: In a preferred embodiment of the present invention, when a polyisobutene with a number-average molecular weight Mn of 500 to 5000 is desired, the reaction temperature in the reaction mixture is preferably from minus 5 (-5) to 25 °C. This homo- or copolymer is particularly preferably a highly reactive polyisobutene with a terminal vinylidene group content of at least 70 mol%, preferably at least 80 mol%, particularly preferably at least 90 mol%.In an alternative preferred embodiment of the present invention, when a number-average molecular weight Mn of 10,000 to 100,000 is desired, the reaction temperature in the reaction mixture is preferably from -10 to -30 °C. In an alternative preferred embodiment of the present invention, when a number-average molecular weight Mn of at least 150,000 to 1,000,000 is desired, the reaction temperature in the reaction mixture is preferably from -50 to -90 °C. The average residence time in the reaction system without mixing, e.g. by circulation, stirring or backmixing, should generally be less than 2 hours, preferably less than 90 minutes and particularly preferably less than 60 minutes. BASF SE 150272 12 The polymerization is generally carried out at a pressure of 700 mbar to 20 bar, in particular at a pressure of 1 bar to 10 bar, especially at a pressure of 1.2 bar to 7 bar.Overpressure normally offers advantages for the C4 hydrocarbon mixture used and for some inert diluents that may be used. In principle, all reactor types suitable for such liquid-phase polymerizations, whether operated batchwise or continuously, can be used as the polymerization reactor for the process according to the invention, for example stirred tanks, stirred tank cascades, tubular reactors, or loop reactors. If the polymerization according to the process according to the invention takes place at or above the boiling point of any inert diluents used or of the monomer to be polymerized, it is preferably carried out in pressure vessels, for example in autoclaves or pressure reactors. Furthermore, it is important that polymerizable compounds are agitated in containers, for example by stirring, natural circulation, or pumping (forced circulation).The polymerization according to the process of the invention can be carried out in the presence of an inert diluent. The inert diluent used should be suitable for reducing the increase in viscosity of the reaction solution that generally occurs during the polymerization reaction to such an extent that the dissipation of the resulting heat of reaction can be ensured. Suitable diluents are solvents or solvent mixtures that are inert to the reagents used.Suitable diluents are, for example, aliphatic hydrocarbons such as n-butane, n-pentane, n-hexane, n-heptane, n-octane and isooctane, cycloaliphatic hydrocarbons such as cyclopentane and cyclohexane, aromatic hydrocarbons such as benzene, toluene and the xylenes, and halogenated hydrocarbons, in particular halogenated aliphatic hydrocarbons such as methyl chloride, dichloromethane, trichloromethane (chloroform), 1,1-dichloroethane, 1,2-dichloroethane, trichloroethane and 1-chlorobutane as well as halogenated aromatic hydrocarbons and alkyl aromatics halogenated in the alkyl side chains such as chlorobenzene, monofluoromethylbenzene, difluoromethylbenzene and trifluoromethylbenzene, as well as mixtures of the aforementioned diluents. It is advisable to free the diluents from impurities such as water or mineral acids before use, for example by adsorption on solid adsorbents such as activated carbon, molecular sieves or ion exchangers.The inert components of the isobutene-containing C4 hydrocarbon mixture used as starting material naturally also act as diluents or as constituents of the solvent mixtures mentioned. BASF SE 150272 13 In a preferred embodiment, no further diluents, which increase the heat capacity of the reaction mixture and necessarily have to be cooled, are added to the reaction mixture. The reaction mixture can be single-phase or multiphase, preferably single-phase. Typically, gaseous, liquid, and solid constituents are present alongside one another; the reaction mixture is preferably kept liquid, although the liquid reaction mixture may also contain solid constituents, for example polymer or Lewis acid. The polymerization according to the process according to the invention is preferably carried out under largely aprotic, in particular under largely anhydrous, reaction conditions.Largely aprotic or largely anhydrous reaction conditions mean that the water content (or the content of protic impurities) in the reaction mixture is less than 50 ppm by weight and in particular less than 5 ppm by weight. Therefore, the starting material is generally dried by physical and / or chemical means before use. It has proven particularly useful to add an organometallic compound, for example an organolithium, organomagnesium, or organoaluminum compound, to aliphatic or cycloaliphatic hydrocarbons used as diluents after conventional pre-purification and pre-drying in an amount sufficient to largely remove traces of water from the solvent. The solvent thus treated is then preferably condensed directly into the reaction vessel.A similar procedure can be used with the isobutene-containing C4 hydrocarbon mixture. Drying with other conventional drying agents such as molecular sieves or pre-dried oxides such as aluminum oxide, silicon dioxide, calcium oxide, or barium oxide is also suitable. Halogenated solvents, for which drying with metals such as sodium or potassium or with metal alkyls is not an option, are freed from water or traces of water using suitable drying agents, for example with calcium chloride, phosphorus pentoxide, or molecular sieves. To terminate the reaction, the reaction mixture is preferably deactivated, for example by adding sufficient amounts of a protic compound, in particular by adding water, alcohols such as methanol, ethanol, n-propanol, and isopropanol, or mixtures thereof with water, or by adding an aqueous base, e.g.an aqueous solution of an alkali metal or alkaline earth metal hydroxide such as sodium hydroxide, potassium hydroxide, magnesium hydroxide or calcium hydroxide, of an alkali metal or alkaline earth metal carbonate such as sodium, potassium, magnesium or calcium carbonate, or of an alkali metal or alkaline earth metal hydrogencarbonate such as sodium, potassium, magnesium or calcium hydrogencarbonate. BASF SE 150272 14 The workup of the reaction mixture comprises at least one phase separation in which the water, alcohol or aqueous solution added to terminate the reaction is separated from the organic phase of the reaction mixture. This can optionally be supported by adding further water or aqueous solutions to wash the reaction mixture. Low boilers present in the reaction mixture are then removed by distillation. The low boilers can be unreacted isobutene, unreacting constituents of the C4 hydrocarbon stream used as starting material and solvents.Furthermore, low oligomers, for example diisobutene, triisobutene, etc., are separated off. Unreacted isobutene and the lower oligomers are preferably at least partially recycled back into the reaction. The distillative removal of these low boilers can be carried out batchwise or, preferably, continuously and in one or more stages, preferably in one or two stages, particularly preferably in two stages. Preference is given to a continuous and at least two-stage distillative separation, with the last stage in each case being carried out under vacuum, for example from 10 mbar to atmospheric pressure, preferably 20 to 900 mbar, and particularly preferably 30 to 800 mbar, and a temperature of 60 to 250 °C, preferably 80 to 200 °C, and particularly preferably 100 to 190 °C. Of course, the distillation can also be carried out in a falling-film, thin-film, or wiped-blade evaporator.For this purpose, the mixture is passed through the apparatus, preferably several times in a cycle, under reduced or normal pressure, for example from 10 mbar to normal pressure, preferably 20 to 900 mbar and particularly preferably 30 to 800 mbar and a temperature of 60 to 250 °C, preferably 80 to 200 and particularly preferably 100 to 190 °C. Advantageously, an inert gas, preferably argon or a nitrogen-containing gas, particularly preferably argon, nitrogen or a mixture of air and nitrogen (lean air), very particularly preferably nitrogen, can be introduced into the distillation apparatus, for example 0.1 - 1, preferably 0.2 - 0.8 and particularly preferably 0.3 - 0.7 m³ / m³h, based on the volume of the liquid mixture. In a particularly preferred variant, the degassing can also take place in a vessel with internals that promote degassing.BASF SE 150272 15 Possible apparatus for degassing also include paddle dryers with or preferably without a cooling zone, optionally with forced discharge devices. Such usable paddle dryers preferably have no separation into heating and cooling zones in order to avoid a sudden drop in temperature. Instead, the temperature of the residue rises as the residue passes through the apparatus, preferably by a temperature gradient which changes by no more than 50 °C as the residue passes through the apparatus, and particularly preferably no significant temperature change as the residue passes through the apparatus, i.e. less than 20 and in particular less than 10 °C. Such paddle dryers are essentially horizontal in design, and the residue is generally conveyed via one or two mixing and kneading shafts inside the apparatus.In the specialist literature, these apparatuses are also referred to as particle bed reactors, kneading dryers, or kneading reactors. Preferably, the paddle dryer has forced conveying in the axial direction. Forced conveying is achieved, for example, by an inclined position of the surfaces of the conveying elements. Axial transport through the apparatus can preferably be achieved by arranging conveying, kneading, and / or mixing elements, for example, disc elements, shafts, screws, blades, wipers, or rotors. To narrow the residence time distribution in the paddle dryer, the product-carrying interior is preferably divided into different segments using baffle-like discs. Particular preference is given to using at least two discs. Heating is via the wall and can be implemented in any desired manner.Heating preferably occurs not only through the outer wall of the apparatus, but also through internal components such as cleaning hooks, segmenting discs, and kneading shafts. The thermal energy introduced into the reactor contents through the walls is typically more than 120 kJ / kg of reactor contents and less than 2400 kJ / kg of reactor contents, preferably more than 220 kJ / kg of reactor contents and less than 1800 kJ / kg of reactor contents, particularly preferably more than 300 kJ / kg of reactor contents and less than 1400 kJ / kg of reactor contents, and most particularly preferably more than 360 kJ / kg of residue and less than 900 kJ / kg of reactor contents. BASF SE 150272 16 The heating section of the reaction mixture fed to the paddle dryer is preferably more than 10% and less than 70% of the total length of the paddle dryer, preferably more than 20% and less than 60%, particularly preferably more than 30% and less than 50% of the total length of the paddle dryer.In the apparatus, a mechanical energy input of 5 W / kg or more is generally sufficient, preferably 10 or more W / kg, more preferably 20 or more, most preferably 40 or more, in particular 80 or more, and especially 100 W / kg or more. As a rule, an energy input of more than 200 W / kg brings no advantages. The specified specific power input here is the input power per amount of reaction mixture in the apparatus. It is also advantageous that the paddle dryer has forced cleaning of the internal, product-contacting surfaces of at least 50%, preferably at least 60%, most preferably at least 70%, and especially at least 80% of these internal, product-contacting surfaces. Forced cleaning is ensured by the proximity of the conveying elements to the outer wall or by the proximity of cleaning hooks to the conveying elements.Such apparatuses are offered, for example, by List AG, Arisdorf, Switzerland, under the trade names Discotherm® B or List-CRP or AP, and by Buss-SMS-Canzler GmbH, Butzbach, Germany, under the names Reasol® or Reactotherm®. Optional discharge devices for the forced discharge of the reaction product can be provided, for example screws, preferably twin screws. In most cases, however, the mechanical conveying devices of the apparatus are sufficient to discharge the product from the apparatus. Boron halides, for example boron trichloride, boron trifluoride, or boron tribromide, are preferred as Lewis acids, preferably boron trifluoride. The reactivity of the boron halides is moderated with the aid of one or more cocatalysts, which form a complex with the boron halide.Preferred cocatalysts are: BASF SE 150272 17 C1 to C5 alkanols, for example methanol, ethanol, n-propanol, isopropanol, n-butanol, sec-butanol, isobutanol, tert-butanol. These alkanols are also preferably used as a mixture of two alkanols, preferably as a mixture of a primary and a secondary alcohol. Aldehydes or ketones having 1 to 20, preferably 2 to 10 carbon atoms, for example formaldehyde, acetaldehyde, propionaldehyde, n-butyraldehyde, isobutyraldehyde, acetone, methyl ethyl ketone or diethyl ketone. Ethers having 2 to 10 carbon atoms, for example dialkyl ethers or cyclic ethers. Examples of these are dimethyl ether, diethyl ether, diisopropyl ether, di-n-butyl ether, tert. Butyl methyl ether, bis(chloroethyl) ether, chloroethyl ethyl ether, tetrahydrofuran or dioxane.Carboxylic acid esters having 2 to 20 carbon atoms, for example C1-C4-alkyl formates, C1-C4-alkyl acetates, or C1-C4-alkyl butyrates. The molar ratio of boron trihalide to the sum of all cocatalysts is usually from 2:1 to 1:10, preferably from 1.5:1 to 1:5, in particular from 1.25:1 to 1:2, especially from 1:1 to 1:1.5. The amount of cocatalyst can influence the setting of the desired molecular weight of the isobutene homopolymer and thus serve to specifically adjust its molecular weight. The amount of polymerization catalyst to be used depends essentially on the type of cocatalyst and the reaction conditions, in particular the reaction temperature and the desired molecular weight of the polymer. It can be determined based on a few random tests for the respective reaction system. Generally, the polymerization catalyst is used in amounts of 0.0001 to 1 wt.-%, in particular 0.0005 to 0.5 wt.%, especially 0.001 to 0.1 wt.%, based in each case on the boron trihalide content in the catalyst complex and on the isobutene used. As an alternative to the preferred boron trihalides, aluminum halides, C1- to C4-alkylaluminum halides, C1- to C4-alkoxyaluminum halides, C1- to C4-alkoxy-C1- to C4-alkylaluminum halides, iron halides, gallium halides, titanium halides, zinc halides, tin dihalides, and tin tetrahalides can also be used as Lewis acids. Among these, aluminum halides, C1- to C4-alkylaluminum halides, iron halides, and titanium halides are preferred; aluminum halides and C1- to C4-alkylaluminum halides are particularly preferred, and aluminum halides are very particularly preferred. BASF SE 150272 18 Among the halides, chlorides and bromides are preferred for these last-mentioned Lewis acids; chlorides are particularly preferred.The latter Lewis acids are also used together with the above-mentioned cocatalysts; the type and amount of cocatalysts are also applicable to these Lewis acids. An advantage of the process according to the invention is that the energy efficiency of the production process is improved, since the external heat input into the reaction mixture is reduced, the released reaction heat can be dissipated at a higher temperature level, which improves the efficiency of the refrigeration machines, and / or allows smaller amounts of solvent or diluent to be used. By increasing energy efficiency, the energy consumption of the process is reduced, which indirectly reduces the emissions associated with the energy supply.Accordingly, the present invention further provides a process for reducing emissions, in particular carbon dioxide emissions, preferably determined as a carbon footprint or life cycle assessment, particularly preferably in accordance with DIN EN ISO 14021, DIN EN ISO 14067, especially the 2019-02 edition, DIN EN ISO 14044, especially the 2006 + A1:2018 edition, and / or DIN EN ISO 14040, especially the 2009-11 edition, by preparing polyisobutene by the process of the invention. Furthermore, the reaction temperature is kept more constant by the measures of the invention than without them.Without these measures, a greater increase in the reaction temperature occurs in the reaction mixture, which, due to the temperature dependence of the molecular weight of the polyisobutene formed, leads to increased formation of higher molecular weight polyisobutene, thus to increased non-uniformity in the molecular weight of the polyisobutene, which is expressed in increased polydispersity. The measures according to the invention can therefore reduce the polydispersity of the polyisobutene in the process.Examples The reaction system for producing polyisobutene with a molecular weight of around 1000 g / mol consists of a loop reactor that can be divided into two sections: The heat exchanger (cooled reaction volume) (reaction temperature -15 °C as the initial temperature of the heat exchanger), which consists of a tube-bundle heat exchanger with the data and dimensions given below, in which the reacting medium (raffinate 1, isobutene content BASF SE 150272 19 46 wt%, feed temperature 4 °C) is pumped through the interior of the tubes via a loop and with the aid of a pumping device. Outside the tubes, but within the tube-bundle heat exchanger, the cooling medium (liquid ammonia at the boiling point) dissipates the reaction heat. The inner surface of these heat exchanger tubes represents the cooled surface.The second part of the reaction system essentially consists of the loop, which is required to keep the reaction medium moving by means of a pumping device and thus ensure the most even distribution possible of isobutene, the resulting polymer and diluent. At the same time, the pumping of the reaction medium promotes the desired heat transfer in the heat exchanger. Although the surface of the loop is usually insulated, it is generally not cooled due to its design. The inner surface of all pipes and equipment in the loop (e.g. the pumping device) therefore represents the uncooled surface. The temperature drop across the heat exchanger, i.e. the difference between the temperature at the outlet and that at the inlet of the cooler, is used as a measure of the efficiency of the tube bundle heat exchanger. This is a measure of the heat to be removed from the uncooled reaction volume.The smaller the temperature drop, the more constant the polymerization reaction temperature is maintained. Example 1 In the energy-efficient example according to the invention, the reactor, a cooled tube bundle heat exchanger, is characterized by the following data: Number of tubes: 1,356 Diameter of the tubes: 0.016 m Length of the tubes: 3.00 m Shell area per tube: 0.15 m² Volume per tube: 0.0006 m³ Total cooled area: 204.5 m² Volume of the tubes (cooled reactor volume): 0.82 m³ BASF SE 150272 20 The non-cooled loop of the reaction system is characterized by the following data: Length of suction side: 7.00 m Length of pressure side: 8.00 m Diameter: 0.40 m Area of ​​loop: 18.85 m² Volume of loop: 1.88 m³ + Volume of 2 dished ends, each 0.4 m³ + Area of ​​2 dished ends, each 0.7 m. 2 Reactor outer surface 8.5 m 2 (Diameter 0.9 m, length 3 m) Total uncooled area 28.75 m 2 Ocool / ONcool 7.11 [m 2 / m 2] Total area of ​​reaction system 223.3 m 2 System: Total volume cooled 0.82 m³ Reaction system: Total volume uncooled 2.68 m 3 Reaction system: total volume 3.50 m 3 Ratio of cooled surface to total volume: 204.5 m² / 3.50 m³ = 58.4 m 2 / m 3 Ratio of cooled surface to non-cooled: 204.5 m² / 28.75 m² = 7.11 [m 2 / m 2] Temperature drop across the cooler: - 1.9 K Example 2 In the energy-efficient example according to the invention, the reactor, a cooled tube bundle heat exchanger, is characterized by the following data: BASF SE 150272 21 Number of tubes: 1,356 Diameter of the tubes: 0.016 m Length of the tubes: 8.00 m Shell area per tube: 0.40212 m² Volume per tube: 0.0016 m³ Total cooled area: 545.28 m² Volume of the tubes (cooled reactor volume): 2.18 m³ The non-cooled loop of the reaction system is characterized by the following data: Length of suction side: 7.00 m Length of pressure side: 8.00 m Diameter: 0.40 m Area of ​​loop: 18.85 m² Volume of loop: 1.88 m³ + volume of 2 dished heads each 0.4 m³ + area 2 dished ends each 0.7 m 2 Reactor outer surface 22.6 m 2 (Diameter 0.9 m, length 8 m) Total uncooled area 42.85 m 2 Ocool / ONcool 12.73 [m 2 / m 2 ] Total area of ​​reaction system - 588.1 m 2System: Total volume cooled 2.18 m³ Reaction system: Total volume uncooled 2.68 m 3 Reaction system: Total volume (cooled + 4.86 m 3 uncooled) Ratio of cooled surface to total volume: 588.1 m² / 4.86 m³ = 121.0 [m 2 / m 3 ] Ratio of cooled surface to non-cooled surface: 588.1 m² / 42.85 m² = 13.7 [m 2 / m 2 ] BASF SE 150272 22 Temperature drop across the cooler: -1.3 K The temperature drop in the cooler is offset by an equally high temperature increase due to the reaction heat released in the uncooled part of the reaction system, which is dissipated in the cooler. It can be seen that with increasing quotients Ocool / V or Ocool / ONcool, this temperature increase in the reaction system is lower, so that the reaction system is subject to smaller temperature fluctuations.

Claims

BASF SE 150272 23 claims 1. A process for the preparation of polyisobutenes having a number-average molecular weight Mn of 500 to 1,000,000 from isobutene or a monomer mixture containing isobutene, characterized in that the reaction is carried out in a reaction system, - for the ratio of cooled surface OKool and reaction volume V the inequality OKool / V ≥ 50 m 2 / m 3 , preferably at least 55, particularly preferably at least 60, very particularly preferably at least 70 and in particular at least 80 m 2 / m 3 , and - for its ratio of cooled surface Ocool and uncooled surface ON-cool the inequality O Kühl / O NKühl ≥ 5 [m 2 / m 2 ], preferably at least 7 and particularly preferably at least 10 [m 2 / m 2], very particularly preferably at least 12, wherein the reaction mixture is circulated via at least one cooled heat exchanger and isobutene-containing feed is introduced into this circulated reaction mixture and the corresponding amount is removed from the circuit, wherein the heat exchanger is selected from the group consisting of temperature-controllable tubular reactors, tube bundle heat exchangers, plate heat exchangers and temperature-controllable tubular reactors with internals.

2. The process according to claim 1, characterized in that the ratio of cooled surface OKool and reaction volume V OKool / V has the value of 250 m 2 / m 33. Process according to claim 1 or 2, characterized in that the isobutene-containing monomer mixture is pure isobutene.

4. Process according to claim 1 or 2, characterized in that the isobutene-containing monomer mixture is a C4 hydrocarbon stream which, in addition to isobutene, also contains 1-butene and cis- and trans-2-butene.

5. Process according to one of the preceding claims, characterized in that the reaction is carried out in the presence of at least one Lewis acid selected from the group consisting of boron halides, aluminum halides, C1- to C4- BASF SE 150272 24 alkylaluminum halides, C1- to C4-alkoxyaluminum halides, C1- to C4-alkoxy-C1- to C4-alkylaluminum halides, iron halides, gallium halides, titanium halides, zinc halides, tin dihalides, and tin tetrahalides.

6. The process according to claim 5, characterized in that, in addition to the Lewis acid, at least one cocatalyst selected from the group consisting of C1- to C5-alkanols, aldehydes or ketones having 1 to 20 carbon atoms, ethers having 2 to 10 carbon atoms, and carboxylic acid esters having 2 to 20 carbon atoms is used.

7. Process according to one of the preceding claims, characterized in that the cooled surface OKool is in contact with a cooling medium whose temperature difference to the reaction mixture is at least 5 K. 8.Process according to one of the preceding claims, characterized in that the cooled surface OKool is in contact with a cooling medium whose temperature difference to the reaction mixture is not more than 30 K.

9. Process according to one of the preceding claims, characterized in that the volume of the circulating reaction mixture is preferably at least 15 times and not more than 100 times the fresh isobutene-containing feed.

10. Process according to one of the preceding claims, characterized in that the temperature rise in the uncooled reaction volume is not more than 2.5 K compared to the inlet temperature into the uncooled reaction volume, preferably not more than 2.2, more preferably not more than 2.0, and most preferably not more than 1.8 K. 11.Process according to one of the preceding claims, characterized in that the temperature rise in the non-cooled reaction volume is 0.3 K or more.

12. Process for preparing polyisobutenes having a number-average molecular weight Mn of 500 to 5000 according to one of the preceding claims, characterized in that the reaction temperature in the reaction mixture is from -5 to 25 °C. BASF SE 150272 25 13. A process for preparing polyisobutenes having a number-average molecular weight Mn of from 10,000 to 100,000 according to any one of claims 1 to 9, characterized in that the reaction temperature in the reaction mixture is from -10 to -30 °C.

14. A process for preparing polyisobutenes having a number-average molecular weight Mn of from 150,000 to 1,000,000 according to any one of claims 1 to 9, characterized in that the reaction temperature in the reaction mixture is from -50 to -90 °C.

15. Process for reducing emissions in the production of polyisobutenes with a number-average molecular weight Mn of 500 to 1,000,000 from isobutene or a monomer mixture containing isobutene, in particular carbon dioxide emissions, preferably determined as a carbon footprint or life cycle assessment, particularly preferably in accordance with DIN EN ISO 14021, DIN EN ISO 14067, here in particular the 2019-02 edition, DIN EN ISO 14044,here in particular the 2006 + A1:2018 edition and / or DIN EN ISO 14040, here in particular the 2009-11 edition, characterized in that the polyisobutene is prepared by a process according to one of claims 1 to 11.

16. Use of a production plant for producing polyisobutenes having a number-average molecular weight Mn of 500 to 1,000,000 from isobutene or an isobutene-containing monomer mixture, comprising at least the following plant components: - at least one apparatus for providing the isobutene-containing monomer mixture, selected from the group consisting of a line, a tank, and a buffer vessel, - at least two storage vessels for at least one Lewis acid and at least one cocatalyst, - a metering device for adding the at least one Lewis acid and at least one cocatalyst to the isobutene-containing monomer mixture - a refrigeration system - at least one reaction system for converting the isobutene-containing monomer mixture,comprising at least one reactor, heat exchanger and lines, wherein the heat exchanger is selected from the group consisting of temperature-controlled tubular reactors, tube bundle heat exchangers, plate heat exchangers and temperature-controlled tubular reactors with internals, and wherein the reaction mixture can be circulated via at least one cooled heat exchanger, - at least one container for terminating the reaction by metering in water or alcohol, and, BASF SE 150272 26 - at least one distillation apparatus for processing the reaction mixture, in which the ratio of cooled surface OKool and reaction volume V in the reaction system satisfies the inequality OKool / V ≥ 50 m 2 / m 3 , preferably at least 55, particularly preferably at least 60, very particularly preferably at least 70 and in particular at least 80 m 2 / m 3, and - for the ratio of cooled surface OKool and uncooled surface ON- Cool the inequality OKool / ONCool ≥ 5 [m 2 / m 2 ], preferably at least 7, particularly preferably at least 10 and most particularly preferably at least 12 [m 2 / m 2 ] applies.