ES2365799T3 - PROCESS TO PREPARE POLYOLEFINIC PRODUCTS. - Google Patents

Abstract

A liquid phase polymerization process for preparing very reactive, low molecular weight polyisobutylene, said process comprising: providing a feed load comprising isobutylene; providing a catalytic composition comprising a complex of BF3 and a primary alcohol of C1 to C8; introducing said feed charge and said catalytic composition into a residual reaction mixture in a reaction zone; intimately intermingling said residual reaction mixture, said feed charge and said catalytic composition to present an intimately mixed reaction mixture in said reaction zone; keep the reaction mixture intimately mixed in its intimate intermingling condition, and keep it at a temperature of at least about 0 ° C while it is in said reaction zone, thereby causing the isobutylene in it to undergo polymerization to form said polyisobutylene; extracting a product stream comprising very reactive low molecular weight polyisobutylene from said reaction zone; and controlling the introduction of said feed charge into said reaction zone and the extraction of said product stream from the reaction zone so that the residence time of the isobutylene undergoing polymerization in the reaction zone is less than 4 minutes.

Description

CROSS REFERENCE TO RELATED APPLICATIONS

This request is a continuation in part of the provisional application Series No. 60 / 154,263, filed on September 16, 1999, pending along with this, of the provisional application Series No. 60 / 160,357, filed on October 19, 1999, in process together with this, and the application for utility Series No. 09 / 515,790, filed on February 29, 2000, pending along with this.

BACKGROUND OF THE INVENTION

Field of the Invention

The present invention relates to polymerization of olefins and the preparation of polyolefin products. In particular, the present invention relates to the preparation of a variety of polyolefin products using a liquid phase polymerization process. In this latter respect, the invention relates to a new liquid phase process for the polymerization of olefins using a modified BF3 catalyst, which is stabilized with a complexing agent.

Background of the prior art

The polymerization of olefins using Friedel-Crafts type catalysts, including BF3, is a generally known process. The degree of polymerization of the products obtained varies according to which of the various known polymerization techniques are used. In this last respect, it will be understood that the molecular weight of the polymeric product is directly related to the degree of polymerization, and that the degree of polymerization can be manipulated by manipulating process parameters to produce a variety of products having respective desired average molecular weights.

Generally speaking, due to the nature and mechanics of the olefin polymerization process, a polyolefin product has a single double bond that remains in each molecule at the end of the polymerization process. The position of this remaining double bond is often an important feature of the product. For example, it is known that polyisobutylene (GDP) molecules, in which the remaining double bond is in a terminal (vinylidene) position, are more reactive than the GDP molecules in which the remaining double bond is internal, that is, it is not in a terminal position. A product of GDP in which at least 50% of the double bonds are in a terminal position can often be referred to as very vinylidene or very reactive GDP. The degree to which a polyolefin product has double terminal links can also be manipulated by manipulating process parameters.

Current processes for oligomerization of olefins often employ BF3 / cocatalyst systems in which BF3 is complexed with a cocatalyst. This is done for a variety of reasons that are well known to those skilled in the field of olefin polymerization. For example, and as explained in US Patent No. 5,408,018, a complexed BF3 catalyst may be useful for manipulating and attempting to balance molecular weight, vinylidene content and GDP polydispersity. Often, the cocatalyst is propanol or a higher alcohol, and such cocatalytic systems are used regardless of whether the desired product is a poly-alpha-olefin or an internal polyolefin. However, the use of alcohols having beta hydrogen atoms in such cocatalytic complexes is problematic, because, over time, BF3 tends to attack beta hydrogen atoms. This leads to the breakdown of alcohol, which makes the catalyst ineffective. In this way, the cocatalytic complex is unstable and often has a very short expiration period.

To solve this problem, many current processes employ a procedure in which the cocatalytic complex is prepared in situ by mixing the alcohol and gaseous BF3 immediately before the introduction of the cocatalytic complex into a reactor. In addition, it is not unusual when carrying out the processes that employ such cocatalytic systems to use an excess of alcohol and to spray gaseous BF3 in the reaction mass at several points downstream, to replenish the catalytic activity. Such a methodology involves a three-phase reaction and the need to use a stirred tank reactor, to provide means for dispersing gaseous BF3 in the reaction mass. These processes use discontinuous reactors, or a series of continuously stirred tank reactors, in series, to provide both a gas handling capacity and to meet the need for a piston flow reactor configuration.

It is also known that alpha-olefins, particularly GDP, can be manufactured in at least two different grades - with normal vinylidene content, and with high vinylidene content. Conventionally, these two grades of product have been obtained by different processes, but both often and normally use a diluted feed load, in which the concentration of GDP can range from 40-60% by weight. More recently it has been observed that at least the GDP with high vinylidene content can be produced using a concentrated feed load having an isobutylene content of 90% by weight or more. Non-reactive hydrocarbons, such as isobutane, n-butane and / or other lower alkanes, normally present in petroleum fractions, can also be included as diluents in the feed load. The feed charge can also often contain small amounts of other unsaturated hydrocarbons, such as 1-butene and 2butene.

Normal grade GDP can range in molecular weight from 500 to 1,000,000 or more, and is generally prepared in a discontinuous process at low temperature, sometimes as low as -50 to -70 ° C. As catalysts, AlCl3, RAlCl2 or R2AlCl are used. The catalyst is not completely removed from the final GDP product. The molecular weight can be controlled by temperature, since the molecular weight of the product varies inversely with temperature. That is, higher temperatures give lower molecular weights. Often the reaction times are of the order of hours. The desired polymer product has a single double bond per molecule, and the double bonds are mostly internal. Generally speaking, at least about 90% of the double bonds are internal, and less than 10% of the double bonds are in a terminal position. Even though it is believed that the formation of double terminal bonds is kinetically favored, prolonged reaction times and the fact that the catalyst is not completely eliminated both favor the transposition of the molecule, so that isomers with thermodynamically internal double bonds are formed. most favored. Normal GDP can be used as a viscosity modifier, particularly in lubricating oils, as a thickener, and as a thickener for plastic films and adhesives. PIB can also be functionalized to produce intermediates for the manufacture of detergents and dispersants for fuels and lubricating oils.

The GDP with high vinylidene content, a relatively new product on the market, is characterized by a large percentage of double terminal bonds, typically greater than 70%, and preferably greater than 80%. This provides a more reactive product, compared to GDP with normal vinylidene content, and therefore this product is also referred to as very reactive GDP. The terms "very reactive" (HR-PIB) and "with high vinylidene content" (HV-PIB) are synonymous. The basic processes for producing HVB-PIB all include a reactor system, which employs BF3 and / or modified BF3 catalysts, so that the reaction time can be closely controlled, and the catalyst can be neutralized immediately once it has been formed the desired product. Since the formation of the double terminal bond is kinetically favored, short reaction times favor high levels of vinylidene. The reaction is paralyzed, usually with a basic aqueous solution, such as, for example, NH4OH, before significant isomerization can take place in internal double bonds. Molecular weights are relatively low. HV-PIB, which has an average molecular weight of around 950-1050, is the most common product. Conversions, based on isobutylene, are maintained at 75-85%, since attempts to drive the reaction to higher conversions reduce the vinylidene content through isomerization. U.S. patents prior to US 4,152,499, dated May 1, 1979, 4,605,808, dated August 12, 1986, 5,068,490, dated November 26, 1991, 5,191,044, dated March 2, 1993, 5,286,823, dated June 22, 1992, 5,408,018, dated April 18, 1995, and 5,962,604, dated October 5, 1999, all refer to a related subject matter.

The U.S. document 4,152,499 describes a process for the preparation of GDP from isobutylene under a blanket of gaseous BF3 that acts as a polymerization catalyst. The process results in the production of a GDP in which 60 to 90% of the double bonds are in a terminal (vinylidene) position.

The U.S. document 4,605,808 describes a process for preparing GDP in which a catalyst consisting of a complex of BF3 and an alcohol is used. It is suggested that the use of such a catalytic complex allows a more effective control of the reaction parameters. Reaction contact times of at least 8 minutes are required to obtain a GDP product in which at least about 70% of the double bonds are in a terminal position.

The U.S. document 5,191,044 describes a GDP production process that requires careful pretreatment of a BF3 / alcohol complex to ensure that all free BF3 is absent from the reactor. The complex must contain an excess of the alcohol complexing agent, in order to obtain a product in which at least about 70% of the double bonds are in a terminal position. The only reaction time exemplified is 10 minutes, and the reaction is carried out at temperatures below 0 ° C.

In addition to the tight reaction time control, the key to obtaining high levels of vinylidene seems to be the control of the reactivity of the catalyst. This has been done in the past by complexing BF3 with various oxygenates, including sec-butanol and MTBE. One theory is that these complexes are actually less reactive than BF3 itself, disproportionately slowing the isomerization reaction and thus allowing greater differentiation between the vinylidene formation (polymerization) reaction and isomerization reaction rates. Mechanisms have also been proposed that suggest that BF3 complexes are not protonated, and thus are not able to isomerize the double terminal bond. This further suggests that water (which can preferably protonate BF3) from these reaction systems should generally be excluded. In fact, previous publications describing the preparation of GDP using BF3 complexes teach that a low water feed (less than 20 ppm) is critical for the formation of the product with high vinylidene content.

HV-PIB is increasingly replacing the normal grade GDP for the manufacture of intermediates, not only due to its greater reactivity, but also due to the requirements for the development of "chloride-free" materials in final product applications. Major GDP derivatives are GDP amines, GDP alkylates, and GDP adducts with maleic anhydride.

GDP amines can be produced using a variety of procedures that involve different GDP intermediates that provide a reactive site for subsequent amination. These intermediates may include, for example, epoxides, halides, maleic anhydride adducts, and carbonyl derivatives.

The reference to HV-GDP as "very reactive" refers to normal grade GDP. The HV-GDP is not yet, in absolute terms, very reactive with respect to the formation of some of these intermediates. Other classes of compounds, for example polyethers, can be much more reactive in the formation of amines and amine intermediates. Polyether-derived amines are known as polyethermines (PEA’s), and are competitive products with GDP amines.

The use of HV-PIB as an alkylating agent for phenolic compounds is fueled by the higher activity and higher yields attainable with HV-PIB. These very long chain alkylphenols are good hydrophobes for surfactants and similar products.

The largest volume of GDP derivatives are the products of the GDP reaction with maleic anhydride. HVPIB is reacted with maleic anhydride through the double bond giving a product with anhydride functionality. This functionality provides reactivity for the formation of amides and other carboxylate derivatives. These products are the basis for most detergents and dispersants currently manufactured for lubricating oils. As mentioned above, GDP-maleic anhydride products can also be used as intermediates in the manufacture of additives for GDP amines fuels.

Other polyolefins that are commercially useful for a variety of purposes include conventional GDP in which the vinylidene content is less than 50%, low molecular weight oligomers (<350, and even <250) of branched monomers such as isobutylene , oligomers and polymers of higher molecular weight of linear C3-C15 alpha-olefins, and higher molecular weight oligomers and polymers of non-alpha olefins (internal double bond) of linear C4-C15. Although these materials are all well known to the person skilled in the field of olefin polymerization, there is always a need for new developments that improve process efficiency and / or product qualities, and reduce operating costs and / or capital expenditures

SUMMARY OF THE INVENTION

The present invention provides a new process for the efficient and economical production of polyolefin products. Generally speaking, the invention provides a liquid phase polymerization process for preparing a polyolefin product having preselected properties. According to the principles and concepts of the invention, the process includes the steps of providing a liquid feed charge comprising at least one olefinic component, and a catalytic composition comprising a stable complex of BF3 and a complexing agent therefor. The feed charge and the catalytic composition are introduced into a residual reaction mixture in a reaction zone of a loop reactor, in which the residual reaction mixture is recirculated at a speed of recirculation sufficient to cause intimate intermingling of the residual reaction mixture, the feed feed added and the catalytic composition added, to thereby present an intimately mixed, recirculating reaction mixture, in said reaction zone. The intimately recirculated intermixed reaction mixture is maintained in its intimately intermingled condition while the heat of reaction is removed therefrom at a rate calculated to provide a substantially constant reaction temperature of the reaction mixture, at the same time as it is it recirculates in said reaction zone. The constant temperature of the reaction is at an appropriate level to cause the olefinic components introduced into said feed charge to undergo polymerization to form the desired polyolefin product in the presence of the catalytic composition. A product stream is extracted from the reaction zone comprising the desired polyolefin product. According to the invention, the introduction of the feed charge into the reaction zone, and the extraction of the product stream from the reaction zone, are controlled so that the residence time of the olefinic components that undergo polymerization in the zone The reaction is appropriate for the production of the desired polyolefin product.

According to a preferred form of the invention, the reaction zone may comprise the tube side of a shell and tube heat exchanger. The heat of the exothermic reaction of the olefinic polymerization can be eliminated simultaneously with its generation by circulating a refrigerant on the side of the exchanger housing. The residence time of the olefinic components that undergo polymerization can be no more than 4 minutes, preferably about 3 minutes. Even more preferably, such residence time may be no longer than about 2 minutes. Even more preferably, such residence time may be no longer than about 1 minute. Ideally, the residence time can be less than 1 minute.

The complexing agent should be such that a stable catalytic complex is formed with BF3. This is particularly advantageous at relatively high reaction temperatures, necessary for oligomerization processes. In this regard, the complexing agent comprises a primary C1-C8 alcohol. In a very preferred form of the invention, the alcohol should not have a hydrogen atom in a carbon P. In this very preferred form of the invention, the alcohol can be, for example, methanol or neopentanol.

According to yet another preferred form of the invention, the complexing agent may comprise a glycol, preferably a glycol in which each hydroxyl group of the glycol is in a primary position, and even more preferably a C1-C8 glycol in which each hydroxyl group of the glycol is in a primary position. In this very preferred form of the invention, the glycol can be, for example, ethylene glycol.

In accordance with the concepts and principles of another aspect of the invention, the molar ratio of BF3 to complexing agent in the catalytic complex can range from about 0.5: 1 to about 5: 1. Preferably, the molar ratio of BF3 to complexing agent in said complex may range from about 0.5: 1 to about 2: 1. Even more preferably, the molar ratio of BF3 to complexing agent in the complex can range from about 0.5: 1 to about 1: 1. Ideally, the molar ratio of BF3 to complexing agent in the complex can be approximately 1: 1. Alternatively, the molar ratio of BF3 to complexing agent in said complex can be approximately 0.75: 1.

According to another aspect of the invention, the process can be carried out desirably so that from about 0.1 to about 10 millimoles of BF3 are introduced into the reaction mixture with said catalytic composition for each mole of olefinic component introduced into said mixture in said feed load. Preferably, from about 0.5 to about 2 millimoles of BF3 can be introduced into the reaction mixture or with the catalytic composition for each mole of olefinic component introduced into the mixture in said feed charge.

Another important preferred feature of the invention involves the continuous recirculation of the reaction mixture at a first volumetric flow rate, and the continuous introduction of the feed and catalyst composition at a second combined volumetric flow rate. Desirably, the ratio of the first volumetric flow rate to the second volumetric flow rate may range from about 20: 1 to about 50: 1. Preferably, the ratio of the first volumetric flow rate to the second volumetric flow rate can range from about 25: 1 to about

40: 1 Ideally, the ratio of the first volumetric flow rate to the second volumetric flow rate can range from about 28: 1 to about 35: 1. With respect to this last aspect of the invention, the ratio of the first volumetric flow rate to the second volumetric flow rate may be such that the concentrations of ingredients in the reaction mixture remain essentially constant, and so that they are established and maintained in said reaction mixture essentially isothermal conditions.

In accordance with the principles and concepts of the invention, the feed load and the catalytic composition can be premixed and introduced into the reaction zone together as a single stream in said second volumetric flow rate. Alternatively, the feed charge and the catalytic composition can be introduced separately into the reaction zone as two streams, whose flow rates together add up to said second volumetric flow rate.

In accordance with the principles and concepts of the invention, the configuration of the reactor, the properties of the reaction mixture, and the first volumetric flow rate can preferably be such that a turbulent flow is maintained in said reaction zone. In this regard, in an ideal form of the invention, a Reynolds number of at least about 2000 is maintained in said reaction zone. Still further, in accordance with the principles and concepts of the invention, the reactor can take the form on the tube side of a shell and tube heat exchanger. In this regard, in an ideal form of the invention, a U of at least about 50 Btu / min is maintained in the reaction zone. ft2 ºF.

Preferably, according to the invention, the feed charge may comprise at least about 30% by weight of said olefinic component. Additionally, the feed charge may include non-reactive hydrocarbon diluents. In this last respect, the feed charge may comprise at least about 30% by weight of said olefinic component, the rest being non-reactive hydrocarbon diluents.

The polymerization process of the invention can be a cationic process. Alternatively, the polymerization process of the invention can be a covalent process. An important feature of the invention is that the polyolefin product of the process of the invention can have a molecular weight of at least about 350, but not more than about 5000. Alternatively, the polyolefin product of the process of the invention can have a molecular weight not greater than about 350, and perhaps not greater than about 250.

According to an important aspect of the invention, the olefinic component that is subject to polymerization may comprise isobutylene, and the polyolefinic product may comprise GDP. Additionally, according to this aspect of the invention, the GDP may have a vinylidene content of at least about 50%. Alternatively, the GDP may have a vinylidene content of no more than about 50%.

According to yet another important aspect of the invention, the olefinic component may be a branched compound, and the product may comprise a one, two, three or four member oligomer. The olefinic component used in the process of the invention may comprise isobutylene, and the polyolefinic product may comprise a GDP oligomer of C12, C16, C20, or C24. Alternatively, the olefinic component may comprise a linear alpha-olefin of C3 to C15, or a non-reactive olefin of C4 to C15, such as 2-butene.

The present invention also provides a new process for the efficient and economical production of HV-GDP. Generally speaking, the invention provides a process for the production of HV-GDP, in which the reaction takes place at higher temperatures and at lower reaction times that were thought of as possible in the past. In particular, the present invention provides a liquid phase polymerization process for preparing very reactive, low molecular weight polyisobutylene. Generally speaking, the process may involve cationic polymerization. However, under some conditions, the polymerization reaction may be covalent. The process includes the provision of a feed load comprising isobutylene and a catalytic composition comprising a complex of BF3 and a complexing agent. The feed charge and the catalyst composition are introduced separately or as a single mixed stream in a residual reaction mixture in a reaction zone. The residual reaction mixture, the feed load and the catalytic composition are intimately intermingled to present an intimately mixed reaction mixture in said reaction zone. The reaction mixture is kept in its intimately intermingled condition, and is maintained at a temperature of at least about 0 ° C while it is in said reaction zone, whereby the isobutylene in the reaction mixture is caused to suffer the polymerization to form a polyisobutylene product. Then a product stream comprising a very reactive, low molecular weight polyisobutylene is extracted from the reaction zone. The introduction of the material into said reaction zone, and the extraction of the product stream from the reaction zone, are controlled so that the residence time of the isobutylene that the polymerization undergoes in the reaction zone is not greater than around 4 minutes According to the invention, it is possible to carry out the reaction so that the residence time is not more than about 3 minutes, not more than about 2 minutes, not more than about 1 minute, and ideally even less than 1 minute .

According to the concepts and principles of the invention, the process can be carried out in such a way that the polyisobutylene thus produced has a molecular weight in the range of about 250 to about 5000, in the range of about 600 to about 4000, in the range of about 700 to about 3000, in the range of about 800 to about 2000, and ideally in the range of about 950 to about 1050. According to the invention, it is possible to control the process so that a particular molecular weight can be achieved, such as, for example, a molecular weight of about 1000.

A main purpose of the invention is to provide a process that can be sufficiently controlled to ensure the production of a polyisobutylene product having a vinylidene content of at least about 70%. More preferably, the GDP product may have a vinylidene content of at least about 80%. Through the use of the invention, vinylidene contents of at least about 90% can also be achieved.

The complexing agent comprises a primary C1-C8 alcohol, and ideally it can be methanol.

To achieve the desired results of the invention, the molar ratio of BF3 to complexing agent in the complex can range from about 0.5: 1 to about 5: 1. Preferably, the molar ratio of BF3 to complexing agent in the complex may range from about 0.5: 1 to about 2: 1. Even more preferably, molar of BF3 to complexing agent in the complex can range from about 0.5: 1 to about 1: 1, and ideally molar of BF3 to complexing agent in the complex can be about 1: 1.

According to the principles and concepts of the invention, it is preferred that from about 0.1 to about 10 millimoles of BF3 be introduced into the reaction mixture with the catalytic composition for each mole of isobutylene introduced into the mixture in the feed charge . Even more preferably, from about 0.5 to about 2 millimoles of BF3 can be introduced into the reaction mixture with said catalytic composition for each mole of isobutylene introduced into the mixture in the feed charge.

The invention provides a process whereby the polydispersity of said polyisobutylene may be no greater than about 2.0, and desirably may be no greater than about 1.65. Ideally, the polydispersity can be in the range from about 1.3 to about 1.5.

According to a preferred aspect of the invention, the reaction zone may comprise a loop reactor in which the reaction mixture is continuously recirculated at a first volumetric flow rate, and said feed charge and said catalytic composition are introduced continuously. at a second combined volumetric flow rate. The ratio of said first volumetric flow rate to said second volumetric flow rate may desirably range from about 20: 1 to about 50: 1, may preferably range from about 25: 1 to about 40: 1, and can ideally range from from about 28: 1 to about 35: 1. In order to achieve the benefits of the invention, the ratio of said first volumetric flow rate to said second volumetric flow rate may preferably be such that the concentrations of ingredients in the reaction mixture remain essentially constant, and / or so that they are established. and essentially isothermal conditions are maintained in said reaction mixture.

The feed charge and the catalytic composition can be premixed and introduced into the reaction zone together as a single stream at said second volumetric flow rate. Alternatively, the feed charge and the catalytic composition can be introduced separately into the reaction zone as two respective streams, whose flow rates together add up to said second volumetric flow rate.

To achieve the desired results of the invention, the configuration of the reactor, the properties of the reaction mixture, and the first volumetric flow rate may be such that a turbulent flow is maintained in said reaction zone. In particular, the system can be such that a Reynolds number of at least about 2000 is achieved and maintained in said reaction zone. The system can also be such that a coefficient of at that reaction zone is achieved and maintained in said reaction zone. thermal transfer (U) of at least about 50 Btu / min. ft2 ºF. For this purpose, the reactor may preferably be the tube side of a shell and tube heat exchanger.

Additionally, according to the concepts and principles of the invention, the feed charge may generally comprise at least about 30% by weight of isobutylene, the rest being non-reactive hydrocarbon diluents.

In a more specific sense, the invention provides a liquid phase polymerization process for preparing polyisobutylene having an average molecular weight in the range of about 500 to about 5000, and a vinylidene content of at least 70%. The process may comprise providing both a feed load comprising isobutylene, as well as a separate catalytic composition formed by a complex of BF3 and a primary alcohol of C1 to C8. The molar ratio of BF3 to alcohol in said complex may be desirably in the range of about 0.5: 1 to about 2: 1. The feed charge and the catalyst composition can be introduced separately, or together as a single stream, into a residual reaction mixture in a reaction zone, and the residual reaction mixture, feed charge and catalyst composition can be mixed. intimately to present an intimately mixed reaction mixture in said reaction zone. The introduction of the catalytic complex into the reaction mixture can preferably be controlled so that about 0.1 to about 10 millimoles of BF3 are introduced for each mole of isobutylene introduced with the feed charge. The intimately intermingling condition of the reaction mixture should preferably be maintained, and its temperature should be maintained at about 0 ° C or higher, while the mixture is in the reaction zone, whereby the isobutylene in the The mixture undergoes polymerization to form said polyisobutylene. Then, a product stream comprising the polyisobutylene product can be extracted from the reaction zone. The introduction of said feed charge into the reaction zone, and the extraction of the product stream from the reaction zone, may preferably be such that the residence time of the isobutylene that the polymerization undergoes in the reaction zone is not greater than about 4 minutes.

Even more desirably, the invention can provide a liquid phase polymerization process for preparing polyisobutylene having an average molecular weight in the range of about 950 to about 1050, a polydispersity in the range of about 1.3 to about 1.5, and a vinylidene content of at least about 80%. According to this preferred aspect of the invention, the process comprises providing both a feed load formed by at least about 40% by weight of isobutylene, as well as a separate catalytic composition formed by a complex of BF3 and methanol, in which the ratio Molar of BF3 to methanol in the complex ranges from about 0.5: 1 to about 1: 1. The feed charge and catalyst composition are introduced separately or together in a residual reaction mixture in a reaction zone. The residual reaction mixture, the feed charge and the catalytic composition are intimately mixed by turbulent flow in said reaction zone, whereby an intimately mixed intermingled reaction mixture is present in the reaction zone. Preferably, the catalytic complex is introduced into the reaction mixture at a rate such that about 0.5 to about 2 molols of BF3 are introduced for each mole of isobutylene introduced into the feed charge. The intimate intermingling condition of the reaction mixture is maintained, and its temperature is maintained at about 0 ° C or higher, while it is in the reaction zone, whereby the isobutylene in it is caused to undergo polymerization. to form said polyisobutylene. A product stream comprising said polyisobutylene is extracted from said reaction zone. According to the invention, the introduction of the feed charge into the reaction zone, and the extraction of the product stream from it, are controlled so that the residence time of the isobutylene that undergoes polymerization in the reaction zone is in the range from about 45 to about 90 seconds.

BRIEF DESCRIPTION OF THE DRAWING

Figure 1 is a schematic illustration of a reactor in the form of a casing heat exchanger and multi-pass tubes, which is useful for carrying out the improved process of the invention; Y

Figure 2 is a schematic illustration of an alternative reactor in the form of a single pass casing and tube exchanger, which is also useful for carrying out the improved process of the invention.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS

According to a very important embodiment of the present invention, an improved liquid phase process is provided for efficient and economical production of GDP. According to this embodiment of the invention, a feed charge current containing isobutylene is contacted in a reaction zone with a catalyst that facilitates the polymerization reaction. Appropriate reaction conditions are provided in the reaction zone. After an appropriate residence time, a product stream containing GDP is extracted from the reaction zone. As mentioned above, many techniques are known to carry out the reaction. However, from a commercial point of view, it is always desirable to improve the efficiency and economy of the process. With the foregoing in mind, the present invention provides an improved process for producing GDP, which can be easily controlled and manipulated to efficiently and economically provide a very reactive, relatively low molecular weight GDP product.

The improved process of the present invention highlights the use of a BF3 catalyst that can be desirably complexed with a complexing agent that appropriately alters the behavior of the catalyst. Normal experts in the field of related art know many other potentially useful catalysts. In particular, many useful catalysts are described in the previous patents cited above. However, for the purposes of the present invention, the complexing agent is a C1-C8 primary alcohol, and ideally methanol.

The molar ratio of BF3 to complexing agent in the catalyst composition may generally be in the range of about 0.5: 1 to about 5: 1, desirably in the range of about 0.5: 1 to about 2: 1, and preferably in the range of about 0.5: 1 to about 1: 1. Ideally, the catalyst composition may simply be a 1: 1 complex of BF3 and methanol. In some preferred embodiments of the invention, the molar ratio of BF3 to complexing agent in said complex may be approximately 0.75: 1.

The temperature in the reaction is greater than 0 ° C, the residence time in the reactor is less than 4 minutes, and the desired vinylidene content (terminal unsaturation) in the GDP product may be preferable and generally greater than about 70%. With these parameters, it is possible to carry out the process to achieve efficiencies and economies not previously thought of as available. According to the present invention, the catalytic concentration and the ratio of BF3 / complexing agent can be manipulated as required to achieve the desired vinylidene content of 70% with a reaction temperature greater than 0 ° C and a residence time in the reactor less than 4 minutes Generally speaking, for the production of GDP, the amount of the BF3 catalyst introduced into the reaction zone should be in the range of about 0.1 to about 10 millimoles per mole of isobutylene introduced into the zone of reaction. Preferably, the BF3 catalyst can be introduced at a rate of about 0.5 to about 2 millimoles per mole of isobutylene introduced into the feed charge.

The process itself includes steps that result in the intimate mixing of the isobutylene-containing reactant stream and the catalytic complex, and / or the removal of heat during the reaction. Intimate mixing can be achieved desirably by turbulent flow. Turbulent flow also enhances heat removal. These conditions, separately or together, allow higher operating temperatures (for example> 0 ° C) and shorter residence times in the reactor (for example <4 minutes) provided by the invention. These important parameters can be achieved by having the catalyzed reaction take place in the tubes of a shell and tube heat exchanger, at a flow rate that results in a turbulent flow.

Many potentially valuable reactors are well known to those of ordinary skill in the art to which the invention pertains. However, for the purposes of a preferred embodiment of the invention, the reactor may be a shell and four-pass heat exchanger as shown in Figure 1, in which it is identified by the number 10. For example, The reactor may have 80 3/8 inch tubes, with a wall thickness of 0.022 inches, each thereby providing an internal tube diameter of 0.331 inches. The reactor can be three feet long, and can have baffles and internal partitions to provide 4 passes with 20 tubes per pass. Such a construction is well known in the techniques of reactors and heat exchangers, and no further explanation is believed necessary.

In operation, according to the preferred method for producing highly reactive GDP, the feed load containing isobutylene enters the reactor system through the pipe 15, which is preferably located close to the lower head 11 of the reactor 10. The pipe 15 directs the feed load on the suction line 20 of a recirculation pump 25. The catalytic complex can be injected into the reactor circulation system through a pipe 30 located next to the lower head 11 of the reactor 10. It should be noted here that, according to the principles and concepts of the invention, the catalytic complex could be also injected separately into the reactor, in which case a separate pump for the catalyst may be necessary.

A catalytic modifier can be added to the feed load via the pipe 16 before the feed charge enters the reactor system. The purpose of the modifier is to help control the vinylidene content of the GDP product. The catalytic modifier can be any compound that contains a solitary pair of electrons, such as an alcohol, an ester or an amine. However, it is noted in this regard that if the amount of modifier is too large, it can actually deactivate the catalyst. The feed load containing the modifier enters the reactor system in the suction line 20 of the circulation pump 25. The composition of the catalytic complex enters the reactor system via line 30 at a location downstream of a pump 25 and adjacent to the first pass as shown in Figure 1. The catalytic complex is preferably a methanol / BF3 complex, with a 1: 1 molar ratio of BF3 to methanol. The amount of modifier added via line 16 may vary from 0 to about 1 mole for each mole of BF3 added as a complex via line 30.

The circulation pump 25 pushes the reaction mixture through the line 35, the control valve 40 and the line 45 into the lower head 11 of the reactor 10. A flowmeter 46 can be placed on the line 45, as shown . The reaction mixture travels in an upward direction through pass 50, down through pass 51, up through pass 52 and down through pass 53. As previously explained, each pass 50, 51, 52 and 53 may preferably include 20 separate tubes. For clarity, only a single respective tube is schematically illustrated in each pass in Figure 1. These tubes are identified by reference numbers 50a, 51a, 52a and 53a. However, as explained above, each pass may preferably consist of a plurality, for example 20, of these individual tubes.

It should be noted here that the reaction mixture should preferably be circulated through the tubes 50a, 51a, 52a, 53a of the reactor at a sufficient flow rate to obtain a turbulent flow, to thereby achieve intimate intermingling between the catalyst complex and the reactants, and an appropriate thermal transfer coefficient to provide proper cooling. In this regard, the flow rate, the properties of the reaction mixture, the reaction conditions and the configuration of the reactor should be appropriate to produce a Reynolds number (Re) in the range of about 2000 to about 3000, and a Thermal transfer coefficient (U) in the range of about 50 to about 150 Btu / min. ft2 ºF in the reactor tubes. Such parameters can generally be obtained when the linear flow rate of a typical reaction mixture through a tube having an internal diameter of 0.331 inches is in the range of about 6 to 9 feet per second.

The circulating reaction mixture leaves the reactor 10 via the suction line 20. The recirculation rate in the system is preferably sufficiently high so that the reactor, in essence, is a continuous stirred tank reactor (CSTR). In this regard, and generally speaking, the speed of recirculation of the reaction mixture should preferably be such that essentially steady state conditions are maintained in the reactor. It is pointed out in the latter respect that, regardless of whether the system is in a non-stationary or stationary state, the design equations can be reduced to those of a CSTR when the recirculation rate is sufficiently high. The reactor can also be of the type sometimes referred to as a loop reactor. With this system, which is only a preferred system since there are many other arrangements that would be apparent to those of normal skill in the art, the flow rate of the reactant mixture in the reactor can be adjusted and optimized regardless of the introduction of feed load and product removal rates in order to achieve a thorough intermingling of the catalyst complex and the reactants, and an appropriate temperature control.

Preferably, a product outlet line 55 may be provided at the upper head 12 at a point that is approximately adjacent to the transition zone between the third and fourth pass. Such placement may be desirable to avoid any potential for loss of unreacted isobutylene. In addition, the placement of the outlet line 55 should be appropriate to facilitate gas purge from the reactor during commissioning. Desirably, a refrigerant can be circulated on the side of the reactor housing, at a rate to remove the heat of reaction and maintain the desired temperature in the reactor.

The product leaving the system via line 55 should be deactivated immediately with a material capable of deactivating the catalyst, such as, for example, ammonium hydroxide. In this way, any potential transposition of the polymer molecule that would displace the double bond away from the terminal position is minimized. The high vinylidene isobutylene product can then be directed to a treatment system (not shown), in which the catalytic salts can be removed and the isobutylene product can be separated from unreacted isobutylene and other undesirable contaminants, such as diluents, etc. The latter materials can then be recycled or diverted for other uses, using known methodology.

With the recirculation system described, the feed rate introduction into the reaction mixture and the product removal rate are each independent of the circulation speed. As those of normal skill in the art will appreciate, the number of passes through the reactor, and the size and configuration of the latter, are simply matters of choice. The flow rates of feed and product extraction can preferably be chosen so that the residence time of the mixture of

The reaction within the reactor is 4 minutes or less, desirably 3 minutes or less, preferably 2 minutes or less, even more preferably 1 minute or less, and ideally less than 1 minute. From a commercial operational point of view, the flow rate should be such that the residence time of the reaction mixture in the reactor is in the range of about 45 to about 90 seconds. In relation to the above, the residence time is defined as the volume of the total reactor system divided by the volumetric flow rate.

10 The recirculation flow rate, that is, the flow rate of the reaction mixture in the system induced by the recirculation pump 25, is controlled, as described above, to achieve appropriate turbulence and / or heat transfer characteristics. This recirculation flow rate is often a function of the system itself and other desired process conditions. For the system described above, the ratio of the recirculation flow rate to the incoming feed flow rate (recirculation ratio) should generally be maintained in the

15 ranges from about 20: 1 to about 50: 1, desirably in the range of about 25: 1 to about 40: 1, and ideally in the range of about 28: 1 to about 35: 1. In particular, in addition to causing turbulence and providing an appropriate heat transfer coefficient, the recirculation flow rate of the reaction mixture should be sufficient to keep the concentrations of the ingredients there essentially constant, and / or to minimize temperature gradients in the circulating reaction mixture, so that

20 essentially isothermal conditions establish and are maintained in the reactor.

As mentioned above, the recirculation ratios can generally be in the range of about 20: 1 to about 50: 1 when the desired product is very reactive GDP. Higher recirculation ratios increase the degree of mixing, and the reactor approaches an isothermal operation, leading to narrower polymer distributions. Lower recirculation ratios decrease the amount of mixing in the reactor, and, as a result, there is a greater discrepancy in the temperature profiles. As the recirculation ratio approaches zero, the design equations for the reactor are reduced to those for a piston flow reactor model. On the other hand, as the recirculation ratio approaches infinity, the model equations are reduced to those for a CSTR. When CSTR conditions are achieved, both the temperature and the composition remain constant, and the composition

30 of the product stream leaving the reactor is identical to the composition of the reaction mixture that is recirculated in the reactor.

Needless to say, after a steady state or near steady state operation has been established in the reactor, as the feed load enters the system, an equal volume of product is pushed out of the reactor loop . Under CSTR conditions, the point at which the product stream is drawn is independent of the geometry of the reactor. However, the upper part of the third pass was chosen for this particular embodiment of the invention so that any non-condensable species or air in the reactor can be conveniently purged at startup. Also, it is preferred that the extraction point be as far as possible from the point at which recent feed load is introduced into the system, only to be sure that the conditions in the reactor have achieved a steady state operation, and therefore they are

40 as stable as possible.

When very reactive GDP is the desired product, the feed load entering the system through line 15 can be any stream containing isobutylene, such as, but not limited to, isobutylene concentrate, dehydrous effluent, or a stream of raff 1 typical. These materials are described respectively below in Tables 1, 2 and 3.

TABLE 1 - Isobutylene Concentrate

Ingredient% by Weight

component of C3 0.00

l-butane 6.41

n-butane 1.68

1-butene 1.30

I-butene 89.19

trans-2-butene 0.83 cis-2-butene 0.38

1,3-butadiene 0.21

TABLE 2 - Dehydrated Effluent

Ingredient% by weight C3 components 0.38 I-butane 43.07 n-butane 1.29 1-butene 0.81 I-butene 52.58 trans-2-butene 0.98 cis-2-butene 0.69 1,3-butadiene 0.20

TABLE 3 -Raff-1

Ingredient% by Weight

C3 components 0.57

I-butane 4.42

n-butane 16.15

1-butene 37,22

I-butene 30.01

trans-2-butene 8.38

cis-2-butene 2.27

1,3-butadiene 0,37

MTBE 0.61

For commercial and process savings, the isobutylene content of the feed load generally should be at least about 30% by weight, the remainder comprising one or more non-reactive hydrocarbon diluents, preferably alkane.

The desired product is a very reactive polyisobutylene of relatively low molecular weight. Thus, the polyisobutylene leaving the reactor via line 55 should have an average molecular weight that is less than about 10,000. Generally speaking, the isobutylene produced should have a weight

10 molecular average in the range of about 500 to about 5000, desirably from about 600 to about 4000, preferably from about 700 to about 3000, even more preferably from about 800 to about 2000, and ideally from around 950 to around 1050. By carefully controlling the various process parameters, it may even be possible to produce a product in which the average molecular weight is consistent and with a certain desired number, for example 1000.

15 Polydispersity of GDP can also be important. The term "polydispersity" refers to the distribution of molecular weights in a given polymer product, and is generally defined as the ratio of the molecular weight of the highest molecular weight molecule to the molecular weight of the lowest molecular weight molecule. Polydispersity can be controlled by carefully maintaining constant concentrations of monomers and isothermal conditions in the reaction mixture. Generally speaking, it is desirable that the polydispersity be as low as possible in order to decrease the content of polyisobutylenes of relatively low or unwanted molecular weights in the product, and thereby improve the quality of the latter. Following the concepts and principles of the present invention, it has been found that the polydispersity of the product can be controlled in

5 more than about 2.0. Preferably, by using the invention, a polydispersity of no more than about 1.65 can be achieved. Even more desirably, the polydispersity can be controlled to be within the range of about 1.3 to about 1.5.

The polyisobutylene product obtained by using the present invention should generally have a terminal (vinylidene) unsaturation content of at least about 70%. That is, at least about 70% of the double bonds remaining in the polymerized product should preferably be in a terminal position. Ideally, the vinylidene content should be not less than about 80%, or even higher. However, the vinylidene content is indirectly related to the conversion rates. That is, the higher the conversion rate, the lower the vinylidene content. In addition, the vinylidene content is directly related in the same way to molecular weight. Consequently, in each process

15 requires a balance between molecular weight, conversion rate and vinylidene content.

EXAMPLE 1

Using the principles and concepts of the invention, a reactor, such as the reactor illustrated in Figure 1, was used to produce a very reactive, low molecular weight polyisobutylene. The feed charge was essentially the same as shown above in Table 1, and the refrigerant circulated on the side of the housing of the

The reactor was a mixture of 35% by weight of methanol and 65% by weight of water. The coolant inlet temperature was 32ºF. A 1: 1 complex catalyst of BF3 / methanol was used. All relevant reactor data and dimensions are set forth below in Table 4.

TABLE 4

Feed load flow rate 1.7 gpm Recirculation flow rate 50 gpm Feed load density 5 lb / gal Conversion 63% by weight Isobutylene concentration in the feed load 92% by weight H of the reaction 398 Btu / lb  of the reaction mixture 4.5 cP = 0.0030 lb / ft seconds Cp of the reaction mixture 0.46 Btu / lb ° F. Effective reaction density 44.9 lb / ft3 Thermal conductivity 0.075 Btu / h ft ° F. Total volume of reactor recirculation system 390.2 in3 Residence time 59.6 seconds Linear velocity inside the tubes 9.3 ft / s Reynolds number 3180 Surface area of the tubes 23.6 ft2 Heat generated 1961 Btu / min . Tlm 37.3 ºF. Heat flow 83.2 Btu / min. ft2 U 133.7 Btu / min. ft2 ºF. Coolant Cp 0.86 Btu / lb ºF.

Refrigerant density 7.70 lb / gal

Coolant flow 39.3 gpm

T of refrigerant 8.0ºF.

Heat removed 2074 Btu / min.

The composition of the product thus obtained is as set forth below in Table 5.

TABLE 5 - Gross Polyisobutylene Product

Ingredient% by weight

C3 components 0.00

I-butane 6.41

n-butane 1.68

1-butene 1.30

I-butene 33.00

trans-2-butene 0.83

cis-2-butene 0.38

1,3-butadiene 0.21

polyisobutylene 56.19

Again it should be noted that one of the main objectives according to the invention is to provide a flow rate to

5 through the reactor and other parameters so that the reaction mixture is in a generally constant state of turbulent flow during the reaction. The turbulent flow results in a two-fold increase in the overall process. First, turbulent flow results in intimate intermingling of reactor contents to enhance reaction kinetics. Secondly, turbulent flow results in an enhancement of the heat transfer coefficient on the side of the tubes, to thereby improve the

10 heat removal from the reaction. These results can be achieved by carrying out the reaction on the tube side of a shell and tube heat exchanger reactor, and by circulating a refrigerant on the side of the shell.

The above description refers to the methodology that allows the polymerization reaction of GDP to be carried out at higher temperatures and shorter residence times than current processes. According to this embodiment of

In the present invention, a stable BF3 catalytic system (BF3 / methanol) can be used. In addition, an improved turbulent loop reactor configuration is advantageously employed that includes a heat exchanger to effect simultaneous heat removal. The turbulent flow also allows intimate mixing of the two-phase reaction system.

In addition to very reactive GDP, the process of the invention provides an improved process for preparing oligomers.

20 and higher molecular weight polymers from olefinic precursors. In general, the process of the invention can be used to produce conventional GDP, low molecular weight oligomers of branched olefins, preferably isobutylene, higher molecular weight oligomers and polymers of linear C3-C15 alpha-olefins, and oligomers and polymers of higher molecular weight of non-alpha reactive olefins of C4-C15. According to this aspect of the invention, and particularly when the desired product is a relatively low molecular weight oligomer.

25 (<350, and perhaps even <250), the catalytic complex is desirably stable, even at the relatively higher reaction temperatures necessary for the production of oligomeric olefinic products.

Examples of processes for the production of relatively low molecular weight oligomers of olefinic monomer components are set forth below. In these examples, a loop reactor is advantageously used as illustrated in FIG. 2. As illustrated in FIG. 2, the reactor 100 may consist of a single tube 102 30 surrounded by a housing 104 of the heat exchanger. In all other essential aspects, the recirculation system may preferably be the same as described in relation to reactor 10 of FIG. 1, except

that a recirculation line 106 is provided to return the residual recirculating mixture from the top of the reactor tube 102 to the pump suction line 20. The output line 55 is directly connected to the recirculation line 106 as shown.

EXAMPLE 2

5 A stream containing 2.19% by weight of isobutane, 61.5% by weight of n-butane, 0.64% by weight of 1-butene, 28.18% by weight of trans-2-butene is introduced and 7.49% by weight of cis-2-butene (35.66% by weight of 2-butene) in a loop reactor system of FIG. 2 via feed line 15 at a flow rate of 156 ml / min. (93.6 g / min.). A catalyst complex containing BF3 / methanol complex (one mole of BF3 to one mole of methanol) is fed to the reactor at a flow rate of 8 ml / min. (10.4 g / min.). The reaction temperature is kept constant at 90 ° F. Everybody

10 the relevant data and dimensions of the reactor are set forth below in Table 6. The effluent from the reactor exits the top of the reaction loop via line 55, and is fed to a decanter (not shown) in which preferably the catalyst is separated from the organic layer. Then, a portion of the catalyst can be returned back to the reactor, reducing the amount of recent catalyst required. The product that leaves the upper part of the decanter is mixed with NH4OH to paralyze any catalyst that remains in the phase

15 organic, and sent to a second decanter. The products are washed twice more with water, and decanted to remove the last traces of catalyst. The oligomeric product composition is given in Table 7.

TABLE 6

HC flow rate 0.0412 gpm Flow around the pump 1.5 gpm HC density 5 lb / gal% Conversion 51% by weight% 2-butene in feed load 36.55% by weight xHrxn 318 Btu / lb

 0.6 cP = 0.0004 lb / ft-s Cp 0.46 Btu / lb-ºF OD of the reactor 0.375 In Thickness of the reactor walls 0.035 In ID of the Reactor 0.305 In Length of the reactor 10.5 Ft Volume of the reactor 9.2 In3 No. of tubes 1 No. of passes 1 Residence time 58.02 seconds Linear speed 6.59 ft / s Surface area 1.03 ft2 Heat generated 12.2 Btu / min.

Tlm 3.0 ° F

Heat flow 11.8 Btu / min.-ft2 U 237.0 Btu / hr-ft2-ºF Re 15531

TABLE 7

C8 7.9% by weight C12 29.8% by weight C16 35.9% by weight C20 16.1% by weight C20 + 10.3% by weight

EXAMPLE 3

It was fed into the loop reactor of FIG. 2 a stream containing 94.0% by weight of 1-decene and 6.0% by weight of C10 isomers, at a flow rate of 10 ml / min. (7.4 g / min.). A catalyst complex containing BF3 / methanol complex (one mole of BF3 to one mole of methanol) was fed into the reactor at a flow rate of 1 ml / min. (1.3 g / min.). The reaction was maintained at a constant temperature of 70 ° F. All relevant reactor data and dimensions are given in Table 8. Both the reactor assembly and the downstream catalyst removal steps are identical to Example 2. The product stream contained about 59.8% by weight of oligomers of C20, and about 40.2% by weight of C30 oligomers.

TABLE 8

HC flow rate 0.00264 gpm Flow around the pump 1.5 gpm HC density 6.2 lb / gal% Conversion 90% by weight% 1-decene in feed load 94% by weight xHrxn 318 Btu / lb

 1.2 cP = 0.0008 lb / ft s Cp 0.50 Btu / lb ° F OD of the reactor 0.375 in Thickness of the reactor walls 0.035 in ID of the reactor 0.305 in Length of the reactor 10.5 ft Volume of the reactor 9, 2 In3 No. of tubes 1 No. of passes 1 Residence time 905.13 seconds Linear speed 6.59 Ft / s Surface area 1.03 ft2 Heat generated 4.4 Btu / min. Tlm 1.2 ºF

Heat flow 4.3 Btu / min-ft2 U 213.2 Btu / hr-ft2-ºF

Re 9604.4

As can be seen from the previous examples, the invention provides a process for preparing a polyolefin product having preselected properties. According to the invention, the process advantageously employs a stable complex of BF3 and a complexing agent therefor. The residual reaction mixture 5 in the reaction zone is recycled at a sufficient recirculation rate to cause intimate intermingling of the reaction mixture. The rate is also such that the heat of reaction is removed from the reaction mixture at a rate calculated to provide a substantially constant temperature of the reaction in the reaction mixture while it is recirculated in the reaction zone. The introduction of the feeding load and the extraction of the product stream are controlled so that the residence time of the

10 olefinic components that undergo polymerization in the reaction zone are suitable for the production of the desired polyolefinic product.

Although the text and examples above have focused on processes in which a single monomer is included in the feed load, it will be apparent to those common in the olefin polymerization technique that, according to the principles and concepts of the present invention, the Feeding load may desirably include, from time to time

15 when, two or more monomers to produce useful copolymer products.

Claims (83)

one.

 A liquid phase polymerization process for preparing very reactive, low molecular weight polyisobutylene, said process comprising: providing a feed load comprising isobutylene; providing a catalytic composition comprising a complex of BF3 and a primary alcohol of C1 to C8; introducing said feed charge and said catalytic composition into a residual reaction mixture in a reaction zone; intimately intermingling said residual reaction mixture, said feed charge and said catalytic composition to present an intimately mixed reaction mixture in said reaction zone; keep the reaction mixture intimately mixed in its intimate intermingling condition, and keep it at a temperature of at least about 0 ° C while it is in said reaction zone, thereby causing the isobutylene in it to undergo polymerization to form said polyisobutylene; extracting a product stream comprising very reactive low molecular weight polyisobutylene from said reaction zone; and controlling the introduction of said feed charge into said reaction zone and the extraction of said product stream from the reaction zone so that the residence time of the isobutylene undergoing polymerization in the reaction zone is less than 4 minutes.

2.

A liquid phase polymerization process as set forth in claim 1, wherein said polyisobutylene has an average molecular weight in the range of about 500 to about 5000 and a vinylidene content of at least 70%, said catalytic composition it comprises a complex of BF3 and a primary alcohol of C1 to C8, in which the ratio of said BF3 to said alcohol in said complex is in the range of about 0.5: 1 to about 1: 1, and the amount The catalytic complex introduced into said reaction zone is controlled so that about 0.1 to about 10 millimoles of BF3 are introduced for each mole of isobutylene introduced into said feed charge.

3.

A liquid phase polymerization process as set forth in claim 1, wherein said polyisobutylene has an average molecular weight in the range of about 950 to about 1050, a polydispersity in the range of about 1.3 to 1 , 5, and a vinylidene content of at least about 80%, said feed charge comprises at least about 40% by weight of isobutylene, said catalytic composition comprises a complex of BF3 and methanol in which the ratio of said BF3 to said methanol is in the range of about 0.5: 1 to about 1: 1, the amount of catalytic complex introduced into said reaction zone is controlled so that about 0.5 to about 2 millimoles are introduced of BF3 for each mole of isobutylene introduced into said feed charge, and the introduction of said feed charge into said reaction zone and the extraction of said product stream from the reaction zone is controlled olan so that the residence time of the isobutylene that undergoes polymerization in the reaction zone is in the inclusive range of about 45 to about 90 seconds.

Four.

A process as set forth in claim 1, 2 or 3, wherein said intimate intermingling of said residual reaction mixture, said feed charge and said catalytic composition is achieved by recirculating them in said reaction zone.

5.

A process as set forth in claim 1, 2, 3 or 4, wherein said reaction zone comprises a loop reactor in which the reaction mixture is continuously recirculated at a first volumetric flow rate, and said charge of feed and said catalytic composition are introduced continuously at a second combined volumetric flow rate.

6.

A process as set forth in any one of claims 1 to 5, further comprising adding a deactivating material capable of deactivating the catalyst to the extracted product stream.

7.

A process as set forth in any one of claims 1 to 6, wherein the speed of recirculation in said reaction zone is such that a Reynolds number of at least about 2000 is maintained therein.

8.

 A process as set forth in claim 1 or 2, wherein said alcohol comprises methanol.

9.

 A process as set forth in claim 1, wherein said complexing agent comprises a glycol.

10.

 A process as set forth in claim 9, wherein said complexing agent comprises ethylene glycol.

eleven.

 A process as set forth in claim 1, wherein the molar ratio of BF3 to complexing agent in said complex ranges from about 0.5: 1 to about 1: 1.

12.

A process as set forth in claim 1, wherein said feed charge and said catalyst composition are premixed and introduced into the reaction zone together as a single stream.

13.

A process as set forth in claim 1, wherein said feed charge and said catalyst composition are introduced separately into the reaction zone as two streams.

14.

A process as set forth in claim 1, wherein said feed load comprises at least 30% by weight of said isobutylene.

fifteen.

A process as set forth in claim 1, wherein said polyolefin product has a molecular weight ranging from about 250 to about 5000.

16.

A process as set forth in claim 1, wherein said polyisobutylene has a vinylidene content of at least 50%.

17.

A process as set forth in claim 1, wherein said liquid feed charge comprises a raff-1 stream.

18.

A process as set forth in claim 1, wherein said feed load comprises an isobutylene concentrate.

19.

 A process as set forth in claim 1, wherein said feed charge comprises a dehydrated effluent.

twenty.

A process as set forth in claim 1, wherein said residence time is not more than 3 minutes.

twenty-one.

A process as set forth in claim 20, wherein said residence time is not more than 2 minutes.

22

A process as set forth in claim 21, wherein said residence time is not longer than 1 minute.

2. 3.

A process as set forth in claim 1, wherein the molar ratio of BF3 to complexing agent in said complex is approximately 1: 1.

24.

 A process as set forth in claim 1, wherein the molar ratio of BF3 to complexing agent in said complex is approximately 0.75: 1.

25.

A process as set forth in claim 1, wherein the polydispersity of said polyisobutylene is not greater than about 2.0.

26.

A process as set forth in claim 25, wherein the polydispersity of said polyisobutylene is not greater than about 1.65.

27.

A process as set forth in claim 7, wherein a turbulent flow is maintained in said reaction zone.

28.

A process as set forth in claim 5, wherein said reaction zone comprises a tube side and a shell and tube heat exchanger.

29.

 A process as set forth in claim 1 or 28, wherein a U of at least about 50 Btu / min is maintained in said reaction zone. ft2 ºF. (5.12 W / M2 ° C).

30

A process as set forth in claim 1 or 28, wherein said heat of reaction is eliminated simultaneously with its generation by the circulation of a refrigerant on the side of the exchanger housing.

31.

A process as set forth in claim 1, wherein the polydispersity is in the range of about 1.3 to about 1.5.

32

A process as set forth in claim 1, wherein a catalytic modifier is introduced into said residual reaction mixture in said reaction zone separately from said catalytic composition.

33.

A process as set forth in claim 32, wherein said catalytic modifier comprises essentially the same composition as said complexing agent.

3. 4.

A process as set forth in claim 33, wherein said catalytic complexing agent and said catalytic modifier each comprise methanol.

35

A reaction system for olefinic polymerization, comprising:

a reactor (10; 100) defining a reaction zone and including an inlet for the olefinic polymerization reaction mixture and an outlet for the olefinic polymerization reaction mixture, the inlet and outlet being in fluid communication with said zone said reactor being adapted and arranged to facilitate the conduit in said zone of an exothermic olefinic polymerization reaction in said olefinic polymerization reaction mixture in the presence of a catalytic composition comprising a catalyst and a catalytic modifier; the reactor being included in a reactor recirculation system; a pipe (15) for the introduction of a feed load containing olefin into the reaction circulation system; and an outlet line (55) for extracting a product stream comprising a polyolefin from the reactor circulation system; the reactor circulation system including a circulation pump (25) arranged to circulate said reaction mixture in said zone at a flow rate that is independent of the rate of introduction of the olefin-containing feed charge in said zone; a line (30) of catalytic composition in communication with the reactor circulation system, for the introduction of a catalytic composition in the reaction mixture at least one line (16) of catalytic modifier in fluid communication with the reactor circulation system, for the introduction of a catalytic modifier in the reaction mixture independently of the introduction of the catalytic composition in the reaction mixture.

36.

A reactor system according to claim 35, wherein said line (30) of catalytic composition is in a different location from line (16) of the catalytic modifier.

37.

A reactor system according to claim 35 or 36, wherein said catalytic composition line (30) is connected to a pipe that is part of the reactor circulation system.

38.

 A method for carrying out an olefin polymerization reaction, comprising:

recirculating an olefin polymerization reaction mixture in a reaction zone of a reactor (10; 100); introducing an olefin-containing feed charge into said recirculated olefinic polymerization reaction mixture, said olefinic polymerization reaction mixture being recirculated at a flow rate that is independent of the rate of introduction of said feed charge into said zone; introducing a catalyst complex composition comprising a catalyst and a catalyst complexing agent in said reaction mixture; subjecting said polymerization reaction mixture to isothermal olefinic polymerization reaction conditions in said zone in the presence of said catalytic composition; introducing a catalytic modifier into said recirculating olefinic polymerization reaction mixture, regardless of the introduction of the composition of the catalyst complex in said zone; Y extracting from said reaction zone a product stream comprising a polyolefin.

39.

A method as set forth in claim 38, wherein said olefin comprises isobutylene, and said polyolefin comprises polyisobutylene.

40

 A method as set forth in claim 38 or 39, wherein said catalyst comprises BF3.

41.

A method as set forth in claim 38 or 39, wherein said reaction zone comprises a loop reactor in which the reaction mixture is continuously recirculated at a first volumetric flow rate, and said feed charge and said catalytic composition are continuously introduce a second combined volumetric flow rate.

42

A method as set forth in claim 38 or 39, wherein said complexing agent comprises methanol.

43

 A method as set forth in claim 42, wherein said catalytic modifier comprises methanol.

44.

 A method as set forth in claim 38 or 39, wherein said catalytic composition of the intimately mixed reaction mixture is maintained at a temperature of at least about 0 ° C while it is in said reaction zone.

Four. Five.

A method as set forth in claim 38, 39 or 44, which includes controlling the introduction of said feed charge into said reaction zone and the removal of said product stream from the reaction zone, such that the residence time of the olefins that undergo polymerization in the reaction zone is not greater

That about 4 minutes.

46.

 A method as set forth in claim 45, wherein the introduction of said feed charge into said reaction zone and the extraction of said product stream from the reaction zone are controlled such that the residence time of the olefin that undergoes polymerization in the reaction zone is not greater than about 3 minutes.

47

A liquid phase polymerization process for preparing a polyolefin product having preselected properties, said process comprising providing a liquid feed charge comprising at least one olefinic component; providing a catalytic composition comprising a complex of BF3 and a complexing agent thereof, said complexing agent comprising a glycol or an aliphatic alcohol having a primary hydroxyl group and no hydrogen in the beta carbon atom, said complex being stable at temperatures necessary to produce said polyolefin product; introducing said feed charge and said catalyst composition into a residual reaction mixture in a reaction zone of a loop reactor; intimately mixing the residual reaction mixture, the feedstock and the catalytic composition by recirculating them in said area to thereby present a recirculating reaction mixture, intimately intermingled, of the residual reaction mixture, the feedstock and the composition catalytic;

keeping the recirculating reaction mixture intimately mixed in its intimate intermingling condition, and removing heat from the reaction from the reaction mixture as it is generated to thereby provide a constant reaction temperature in the reaction mixture while that it is recirculated in said reaction zone, said constant reaction temperature being at an appropriate level to cause the olefinic components introduced into said feed charge to undergo polymerization to form said polyolefin product in the presence of said catalytic composition; extracting a product stream comprising polyolefin product and catalytic composition of said reaction zone; and controlling the introduction of said feed charge into said reaction zone and the extraction of said product stream from the reaction zone so that the residence time of the olefinic components that undergo polymerization in the reaction zone is appropriate for the production of said polyolefin product.

48.

A process as set forth in claim 47, further comprising adding a deactivating material capable of deactivating the catalyst in the extracted product stream.

49.

 A process as set forth in claim 47 or 48, wherein the speed of recirculation in said reaction zone is such that a Reynolds number of at least about 2000 is maintained therein.

fifty.

A process as set forth in claim 47, wherein said residence time is not longer than about 4 minutes.

51.

 A process as set forth in claim 47, wherein said complexing agent comprises an alcohol.

52

 A process as set forth in claim 51, wherein said complexing agent comprises a C1-C8 alcohol.

53.

 A process as set forth in claim 52, wherein said alcohol comprises methanol.

54

 A process as set forth in claim 47, wherein said complexing agent comprises a glycol.

55.

 A process as set forth in claim 54, wherein said complexing agent comprises ethylene glycol.

56.

A process as set forth in claim 47, wherein the molar ratio of BF3 to complexing agent in said complex ranges from about 0.5: 1 to about 1: 1.

57.

 A process as set forth in claim 47, wherein said feed charge and said catalyst composition are premixed and introduced together in the reaction zone as a single stream.

58.

 A process as set forth in claim 47, wherein said feed charge and said catalytic composition are introduced separately into the reaction zone as two streams.

59.

A process as set forth in claim 47, wherein said feed load comprises at least 30% by weight of said olefinic component.

60

A process as set forth in claim 47, wherein said polyolefin product has a molecular weight ranging from about 250 to about 5000.

61.

A process as set forth in claim 47, wherein the olefin component comprises isobutylene, and the polyolefin product comprises polyisobutylene.

62

A process as set forth in claim 61, wherein said polyisobutylene has a vinylidene content of at least 50%.

63.

A process as set forth in claim 47, wherein the olefinic component comprises a linear alpha olefin from C3 to C15.

64.

 A process as set forth in claim 47, wherein the olefinic component comprises a non-alpha olefin from C4 to C15.

65

A process as set forth in claim 61, wherein said liquid feed charge comprises a raff-1 stream.

66.

A process as set forth in claim 61, wherein said feed load comprises an isobutylene concentrate.

67.

A process as set forth in claim 61, wherein said feed charge comprises a dehydrated effluent.

68.

A process as set forth in claim 50, wherein said residence time is not longer than 3 minutes.

69.

A process as set forth in claim 68, wherein said residence time is not more than 2 minutes.

70.

A process as set forth in claim 69, wherein said residence time is not longer than 1 minute.

71.

A process as set forth in claim 47, wherein the molar ratio of BF3 to complexing agent in said complex is approximately 1: 1.

72.

 A process as set forth in claim 47, wherein the molar ratio of BF3 to complexing agent in said complex is approximately 0.75: 1.

73

 A process as set forth in claim 47, wherein the reaction mixture is continuously recirculated at a first volumetric flow rate, and said feed charge and said catalytic composition are continuously introduced at a second combined volumetric flow rate.

74.

A process as set forth in claim 61, wherein the polydispersity of said polyisobutylene is not greater than about 2.0.

75.

A process as set forth in claim 74, wherein the polydispersity of said polyisobutylene is not greater than about 1.65.

76

A process as set forth in claim 49, wherein a turbulent flow is maintained in said reaction zone.

77.

A process as set forth in claim 47, wherein said reaction zone comprises a tube side of a shell and tube heat exchanger.

78.

A process as set forth in claim 47 or 77, wherein a U of at least about 50 Btu / min is maintained in said reaction zone. ft2 ºF. (5.12 W / M2 ° C).

79.

A process as set forth in claim 47 or 77, wherein said heat of reaction is eliminated simultaneously with its generation by the circulation of a refrigerant on the side of the exchanger housing.

80.

 A process as set forth in claim 61, wherein the polydispersity is in the range of about 1.3 to about 1.5.

81.

A process as set forth in claim 47, wherein a catalytic modifier is introduced into said residual reaction mixture in said reaction zone separately from said catalytic composition.

82.

A process as set forth in claim 81, wherein said catalytic modifier comprises essentially the same composition as said complexing agent.

83.

A process as set forth in claim 82, wherein said catalytic complexing agent and said catalytic modifier each comprise methanol.