CN108602043B

CN108602043B - Apparatus and process for high pressure polymerization of ethylene

Info

Publication number: CN108602043B
Application number: CN201780010186.0A
Authority: CN
Inventors: P·J·克莱曼斯; H·A·拉曼斯; P·H·克纳利森
Original assignee: ExxonMobil Chemical Patents Inc
Current assignee: ExxonMobil Chemical Patents Inc
Priority date: 2016-02-26
Filing date: 2017-01-06
Publication date: 2021-05-14
Anticipated expiration: 2037-01-06
Also published as: WO2017146823A1; KR20180108719A; MY195216A; EP3419751A1; KR102137105B1; CN108602043A

Abstract

An apparatus and process for polymerizing ethylene at high pressure is disclosed that provides more than one injection point for one initiator injection pump.

Description

Apparatus and process for high pressure polymerization of ethylene

Cross Reference to Related Applications

The present application claims the benefit of U.S. provisional application No.62/300367 filed on 26/2016 and european application No.16167667.1 filed on 29/4/2016, the disclosures of which are incorporated herein by reference in their entireties.

Technical Field

The present invention relates to an apparatus and a process for polymerizing ethylene under high pressure. In particular, the present invention relates to apparatus and methods that provide more than one injection point for one initiator injection pump.

Background

High pressure reactors, such as tubular reactors and autoclaves, are used for the polymerization of ethylene at high pressures, for example pressures in excess of 1000bar (100MPa) and up to 3000bar (300MPa) or more. In such processes, fresh ethylene from an ethylene supply is compressed to reactor pressure and then combined with initiator and any comonomer (if available) in the reactor and polymerized to produce a mixture comprising the predominant polymer and unreacted monomer. This mixture leaves the reactor through a valve (commonly referred to as a high pressure let down valve) and then enters a separation system where unreacted monomer is separated from the polymer and recycled back to the suction of the secondary compressor where it is combined with fresh ethylene from the primary compressor.

The initiator, typically oxygen or an organic free radical initiator, is injected into the reactor system to initiate the polymerization reaction. Generally, different organic peroxides having different half-life time temperatures are used depending on the desired reaction process. The organic peroxide (mixture) is diluted in an inert organic solvent for safety and easier handling and metering purposes. The initiator (mixture) is injected into the reactor at one or more points to start the polymerization reaction. The injection is typically by a high pressure initiator injection pump. A pump (usually with a back-up pump) is used to inject initiator into the reaction zone through an injection point. The pump output (initiator flow) is controlled by the measured temperature in the reactor (either the individual temperature or the temperature from a peak detector in the reactor). See, e.g., U.S. patent application U.S. Pat. No.3628918 or U.S. publication No.2005/0192414, WO2004/108271, WO2007/018871, WO2011/008197 or WO 2013/154690. Other background references include EP2481477A and WO 2014/046835.

In the present invention, an injection pump is designed to inject initiator at more than one injection point. While maintaining the initiator supply for each reaction zone, the inventive concept can increase the overall output of each injection pump by eliminating the limitation of using at least one injection pump to supply only one reaction zone through one injection point, thus reducing the number of injection pumps required, which can promote the desired cost efficiency of a high pressure reactor polymerization plant.

Summary of The Invention

An apparatus and process for polymerizing ethylene at high pressure is provided.

In one embodiment, the invention comprises an apparatus for high pressure polymerizing ethylene comprising a high pressure reactor in which monomers are polymerized to form a product mixture comprising polymer and unreacted monomers, wherein the reactor has more than one reaction zone and at least one injection pump fluidly connected to the reactor, wherein each injection pump is fluidly connected to the reactor at more than one injection point. Preferably, the apparatus further comprises a primary compressor for compressing the monomer to an intermediate pressure, and a secondary compressor for compressing the monomer to the pressure of the reactor upstream of the reactor; a high pressure let down valve downstream of the reactor through which the product mixture from the high pressure reactor flows; a separation system downstream of the high pressure let down valve having at least two stages for separating the product mixture into polymer and unreacted monomer; and a recycling system for returning unreacted monomers to the secondary compressor.

In another embodiment, the invention relates to a process for the high pressure polymerization of ethylene comprising the steps of: compressing the monomer to an intermediate pressure in a primary compressor, then mixing the monomer with recycled monomer, and further compressing the monomer to the pressure of the reactor upstream of the reactor in a secondary compressor; introducing the monomer into a high pressure reactor having more than one reaction zone; injecting initiator into the reactor at more than one injection point by at least one injection pump, wherein each injection pump injects at more than one injection point; contacting the monomer in the reactor with an initiator under high pressure polymerization conditions to polymerize to form a product mixture; discharging the product mixture through a high pressure let down valve, cooling the product mixture, and separating the product mixture into polymer and unreacted monomer in a separation system having at least two stages; and recycling the unreacted monomer to the secondary compressor.

Preferably, the initiator is injected into the reactor at least 2, at least 3 or at least 4 injection points. Preferably, the initiators are injected into different reaction zones having the same initiator composition. Preferably, the reactor is a tubular reactor, an autoclave, or a combination thereof.

Brief description of the drawings

FIG. 1 shows a schematic representation of a high pressure polymerization system with a conventional initiator injection scheme, wherein an injection pump injects initiator into a reaction zone through an injection point.

FIG. 2 shows a schematic representation of a high pressure polymerization system with the initiator injection scheme of the present invention, wherein one injection pump is designed to inject initiator into more than one reaction zone at more than one injection point.

Detailed description of the invention

Various specific embodiments, versions of the invention will now be described, including preferred embodiments and definitions used herein. While the following detailed description sets forth specific preferred embodiments, those skilled in the art will appreciate that these embodiments are merely exemplary, and that the present invention can be practiced in other ways. Any reference to "the invention" may refer to one or more, but not necessarily all, of the inventions defined by the claims. Headings are used for convenience only and do not limit the scope of the invention.

As used herein, "polymer" may be used to refer to homopolymers, copolymers, interpolymers, terpolymers, etc. A "polymer" has two or more identical or different monomer units. A "homopolymer" is a polymer having the same monomer units. A "copolymer" is a polymer having two or more monomer units that are different from each other. The term "different" when used in reference to a monomeric unit means that the monomeric units differ from each other by at least 1 atom or are isomerically different. Also, as used herein, the definition of polymer includes copolymers and the like.

As used herein, the term "monomer" refers to ethylene and any mixture of ethylene with one or more comonomers. Thus, as used herein, the terms "polyethylene", "ethylene polymer" and "ethylene copolymer" denote a polymer or copolymer comprising at least 50 mol% ethylene units (preferably at least 70 mol% ethylene units, more preferably at least 80 mol% ethylene units, even more preferably at least 90 mol% ethylene units, even more preferably at least 95 mol% ethylene units or 100 mol% ethylene units (in the case of homopolymers)).

As used herein, "high pressure polymerization" refers to highly exothermic polymerization reactions conducted in reactors such as tubular reactors and autoclaves at high reactor operating pressures, for example, in excess of 1000bar (100MPa) and up to 3000bar (300MPa) or more.

Polymerization initiator and initiator composition

The initiator is used to initiate free radical polymerization of ethylene and optional comonomer(s). Suitable initiators are organic peroxides. Usually, mixtures of different peroxides are used, so-called "peroxide cocktail" (cocktail). Such mixtures of several peroxide initiators typically include peroxides having different half-lives: typically one being active at the minimum temperature required for a given reaction start temperature (about 120 ℃ to about 160 ℃) and one being active at the maximum temperature (up to about 335 ℃) for the maximum temperature desired. The choice of a suitable combination of different peroxides depends on the reactor setup and the desired reaction temperature profile along the length of the reactor and is within the general knowledge of the person skilled in the art.

Organic peroxides useful as polymerization initiators are well known in the art. Classes of peroxide initiators particularly useful in the present invention are, for example, the following: diacyl peroxides, dialkyl peroxydicarbonates, tertiary alkyl peroxyesters, OO-tertiary alkyl O-alkyl monoperoxycarbonates, di-tertiary alkyl peroxides, di (tertiary alkyl peroxy) ketals, tertiary alkyl hydroperoxides and ketone peroxides.

Non-limiting examples of useful peroxides are for instance the following: dibenzoyl peroxide, dilauroyl peroxide, succinic acid peroxide, diisononanoyl peroxide, dioctanoyl peroxide, tert-butyl peroxybenzoate, tert-butyl peroxyacetate, tert-butyl peroxymaleate, tert-butyl 2-ethylperoxyhexanoate, tert-butyl peroxy-3, 5, 5-trimethylhexanoate, tert-amyl 2-ethylperoxyhexanoate, 2, 5-bis (2-ethylhexanoyl-peroxy) 2, 5-dimethylhexane, tert-butyl peroxypivalate, alpha-cumyl peroxyneoheptanoate, 3-hydroxy-1, 1-dimethylbutyl peroxyneodecanoate, OO-tert-butyl-O- (isopropyl) monoperoxycarbonate, OO-tert-amyl-O- (2-ethylhexyl) monoperoxycarbonate, ethyl-3, 3-di (tert-amylperoxy) butyrate, n-butyl-4, 4-di (tert-butylperoxy)Esterified) pentanoate, 1, 1-di (tert-butylperoxy) cyclohexane, 2, 2-di (tert-butylperoxy) butane, 1, 1-di (tert-amylperoxy) cyclohexane, 2, 5-di- (tert-butylperoxy) -2, 5-dimethyl-3-hexyne, 2, 5-di- (tert-butylperoxy) -2, 5-dimethyl-hexane, 2, 5-di (tert-butylperoxy) -2, 5-dimethylhexane, 1,3(4) -bis (2- (tert-butylperoxy) -1-methylethyl) -benzene, di (tert-butyl) peroxide (DTBP), di (tert-amyl) peroxide, dicumyl peroxide, tert-butylcumyl peroxide, tert-butyl peroxyisopropylcarbonate, tert-butyl peroxyisobutyrate, di (n-propyl) peroxydicarbonate, di (sec-butyl) peroxydicarbonate, di (2-ethylhexyl) peroxydicarbonate, di (n-hexadecyl) peroxydicarbonate, di (4-tert-butylcyclohexyl) peroxydicarbonate, tert-butyl hydroperoxide, tert-amyl hydroperoxide, alpha-cumyl hydroperoxide, 2, 5-dihydroperoxy-2, 5-dimethylhexane, p-terpene-alkane hydroperoxide, m/p-isopropyl-alpha-cumyl hydroperoxide. Such peroxides are for example known under the trade name Trigonox^TMAnd Perkadox^TMCommercially available from Akzo Nobel, or in Luperox^TMCommercially available from Arkema.

Preferred initiator mixtures, especially for tubular reactors, contain a minimum of one and up to five different types of initiator. Suitable mixtures of different organic peroxides (often referred to as peroxide cocktail-type mixtures) are known to the person skilled in the art.

In a preferred embodiment according to the invention, the initiators (or initiator mixtures) are injected into the different reaction zones in the same initiator composition. The polymerization initiator composition used herein comprises at least one, preferably several, of the above-described polymerization initiators (which are dissolved in an organic solvent, as described further below), and optionally one or more additional regulators (also described further below).

Suitable organic solvents may include one or more non-coordinating inert liquids including, but not limited to, straight and branched chain hydrocarbons such as propane, isobutane, butane, pentane, isopentane, hexane, isohexane, heptane, octaneAlkanes, n-octane, dodecane, isododecane, and mixtures thereof; cyclic and alicyclic hydrocarbons such as cyclohexane, cycloheptane, methylcyclohexane, methylcycloheptane and mixtures thereof, for example, are commercially available (Isopars from ExxonMobil)^TM) (ii) a Perhalogenated hydrocarbons, e.g. perfluorinated C₄-C₁₀Alkanes, chlorobenzene, and aromatic and alkyl-substituted aromatic compounds such as benzene, toluene, mesitylene, and xylene. Suitable solvents also include liquid olefins, which may serve as monomers or comonomers, including ethylene, propylene, 1-butene, isobutylene, 1-hexene, 1-pentene, 3-methyl-1-pentene, 4-methyl-1-pentene, 1-octene, and 1-decene. In certain embodiments, the initiator may comprise butane, n-octane, or one or more C₉-C₁₂A mixture of paraffins.

The peroxide initiator (or initiator mixture) may comprise from about 5 to about 50 weight percent of the polymerization initiator composition, preferably from about 5 to about 40 weight percent, and more preferably from about 10 to about 40 weight percent.

Optionally, where additional modifier (other than the initiator solvent) is used, such additional modifier as described above may be added to the reactor with the monomer feed or via one or more separate injection points. The amount of the transfer agent can be as high as is necessary to control the melt index of the product to a concentration in the reaction mixture as high as is required to meet the desired specifications.

Preferably, the organic solvent comprises from about 50 to about 95 wt%, more preferably from about 65 to about 85 wt% and most preferably from about 70 to about 85 wt% of the total solution of peroxide(s) contained in the organic solvent.

The one or more polymerization initiators comprise from about 30 to about 1500 ppm by weight, preferably from about 50 to about 1000 ppm by weight, based on the monomer feed (ethylene and optional comonomer (s)).

The amount of organic solvent in which the initiator is dissolved may correspond to about 100 to about 5000 ppm by weight, preferably about 250 to about 3000 ppm by weight, relative to the monomer feed.

The polymerization initiator composition according to the invention may further comprise conventional additives, such as radical quenchers, to stabilize the initiator composition during storage.

Conditioning agents

The process of the invention preferably comprises the use of a modulator. The term "modifier" is also used herein interchangeably with "chain transfer agent" and refers to a component that can be added to a polymerization process to control the molecular weight of a polymer by facilitating chain transfer.

Examples of regulators may include, but are not limited to, tetramethylsilane, cyclopropane, sulfur hexafluoride, methane, t-butanol, perfluoropropane, deuterated benzene, ethane, ethylene oxide, 2-dimethylpropane, benzene, dimethyl sulfoxide, vinyl methyl ether, methanol, propane, 2-methyl-3-buten-2-ol, methyl acetate, t-butyl acetate, methyl formate, ethyl acetate, butane, triphenylphosphine, methylamine, methyl benzoate, ethyl benzoate, N-diisopropylacetamide, 2, 4-trimethylpentane, N-hexane, isobutane, dimethoxymethane, ethanol, N-heptane, N-butyl acetate, cyclohexane, methylcyclohexane, 1, 2-dichloroethane, acetonitrile, N-ethylacetamide, propylene, 1-butene, N-decane, N-butyl alcohol, methyl acetate, N-butyl acetate, cyclohexane, methylcyclohexane, 1, 2-dichloroethane, acetonitrile, N-ethylacetamide, propylene, 1-butene, n, N-diethylacetamide, cyclopentane, acetic anhydride, N-tridecane, N-butyl benzoate, isopropanol, toluene, hydrogen, acetone, 4-dimethylpentene-1, trimethylamine, N-dimethylacetamide, isobutylene, N-butyl isocyanate, methyl butyrate, N-butylamine, N-dimethylformamide, diethyl sulfide, diisobutylene, tetrahydrofuran, 4-methylpentene-1, p-xylene, p-dioxane, trimethylamine, butene-2, 1-bromo-2-chloroethane, octene-1, 2-methylbutene-2, cumene, butene-1, methylvinyl sulfide, N-butyronitrile, 2-methylbutene-1, ethylbenzene, N-hexadecene, 2-butanone, N-butyl isothiocyanate, Methyl 3-cyanopropionate, tri-n-butylamine, 3-methyl-2-butanone, isobutyronitrile, di-n-butylamine, methyl chloroacetate, 3-methylbutene-1, 2-dibromoethane, dimethylamine, benzaldehyde, chloroform, 2-ethylhexene-1, propionaldehyde, 1, 4-dichlorobutene-2, tri-n-butylphosphine, dimethylphosphine, methyl cyanoacetate, carbon tetrachloride, bromotrichloromethane, di-n-butylphosphine, acetaldehyde, propionaldehyde and phosphine. Additional details and other suitable transfer agents are described in Advances in Polymer Science, Vol.7, pp.386-448 (1970).

Preferably, the polyethylene produced by the apparatus or according to the process described herein comprises one or more C₂-C₁₂An unsaturated modifier. The C is₂-C₁₂The unsaturated modifier contains at least one unsaturation, but may also contain multiple conjugated or non-conjugated unsaturations. In the case of multiple unsaturations, it is preferred that they are non-conjugated. In certain embodiments, the C₂-C₁₂The unsaturation of the unsaturated modifier may be disubstituted in the beta position with one or more alkyl groups. Preferred is C₂-C₁₂The unsaturated modifier comprises propylene, isobutylene, or a combination thereof. The amount of modifier(s) can be from as low as about 0.1 wt%, 0.3 wt%, or 0.8 wt% to as high as about 3.0 wt%, 6.0 wt%, or 10.0 wt%, based on the total weight of the polyethylene.

The modifier may be added to the reaction mixture in any suitable manner. It may be contained in the polymerization initiator composition. Alternatively, the modifier may be injected into the monomer feed, for example into the inlet pipe feeding the secondary compressor. Since the moderator is usually not completely consumed in one pass through the reactor, it is usually also present in the recycled ethylene returned to the secondary compressor in a certain amount.

Comonomer

The process of the present invention can be used not only to produce ethylene homopolymers but also to produce ethylene copolymers. Such comonomer(s) will be pressurized and injected into the primary and/or secondary compressor and then fed to the polymerization reactor together with ethylene.

Typical comonomers include, but are not limited to: vinyl ethers such as vinyl methyl ether, vinyl n-butyl ether, vinyl phenyl ether, vinyl β -hydroxyethyl ether and vinyl dimethylaminoethyl ether; olefins such as ethylene, propylene, butene-1, cis-butene-2, trans-butene-2, isobutylene, 3-dimethylbutene-1, 4-methylpentene-1, hexane-1, octene-1 and styrene; vinyl type esters such as vinyl acetate, vinyl butyrate, vinyl valerate, and vinylene carbonate; halogenated olefins such as vinyl fluoride, vinylidene fluoride, tetrafluoroethylene, vinyl chloride, vinylidene chloride, tetrachloroethylene, and chlorotrifluoroethylene; acrylic type esters such as methyl acrylate, ethyl acrylate, N-butyl acrylate, t-butyl acrylate, 2-ethylhexyl acrylate, α -cyanoisopropyl acrylate, β -cyanoethyl acrylate, o- (3-phenylprop-1, 3-diketo) phenyl acrylate, methyl methacrylate, N-butyl methacrylate, t-butyl methacrylate, cyclohexyl methacrylate, 2-ethylhexyl methacrylate, methyl methacrylate, glycidyl methacrylate, β -hydroxyethyl methacrylate, β -hydroxypropyl methacrylate, 3-hydroxy-4-carbamethoxyphenyl methacrylate, N-dimethylaminoethyl methacrylate, t-butylaminoethyl methacrylate, N-hexylamino methacrylate, N, 2- (1-aziridinyl) ethyl methacrylate, diethyl fumarate, diethyl maleate and methyl crotonate; other acrylic acid type derivatives such as acrylic acid, methacrylic acid, crotonic acid, maleic acid, methyl hydroxymaleate, itaconic acid, acrylonitrile, fumaronitrile, N-dimethylacrylamide, N-isopropylacrylamide, N-t-butylacrylamide, N-phenylacrylamide, diacetoneacrylamide, methacrylamide, N-phenylmethylacrylamide, N-ethylmaleimide, and maleic anhydride; and other compounds such as allyl alcohol, vinyltrimethylsilane, vinyltriethoxysilane, N-vinylcarbazole, N-vinyl-N-methylacetamide, vinyldibutylphosphine oxide, vinyldiphenylphosphine oxide, bis- (2-chloroethyl) vinylphosphonate, and vinylmethylsulfide.

Examples of preferred comonomers are vinyl acetate, methyl acrylate, methacrylic acid, ethyl acrylate, butyl acrylate or acrylic acid or mixtures thereof. In some polyethylenes containing comonomer, the amount of comonomer is less than 10 wt%, but may also be 5 wt% or less, 3 wt% or less, or even 1.5 wt% or less. However, in other co-monomer containing polyethylenes, the amount of co-monomer may be 10 wt% or more, for example 15, 20 or 30 wt% or more, depending on the desired end use of the polymer.

Generally, the purity of the ethylene feed suitable for use in the process according to the invention is provided by a steam cracker of the prior art. In order not to interfere with the free radical initiated reaction, the oxygen content of the feed should be below 5 ppm.

Polymerization process

In one embodiment of the present invention, a process for high pressure polymerization of ethylene may comprise the steps of: compressing the monomer to an intermediate pressure in a primary compressor, then mixing the monomer with recycled monomer, and further compressing the monomer to the pressure of the reactor upstream of the reactor in a secondary compressor; introducing the monomer into a high pressure reactor having more than one reaction zone; injecting initiator into the reactor at more than one injection point by at least one injection pump, wherein each injection pump injects at more than one injection point; contacting the monomer in the reactor with an initiator under high pressure polymerization conditions to polymerize to form a product mixture; discharging the product mixture through a high pressure let down valve, cooling the product mixture, and separating the product mixture into polymer and unreacted monomer in a separation system having at least two stages; and recycling the unreacted monomer to the secondary compressor.

The process for high pressure polymerization of ethylene (optionally with one or more comonomers) according to the present invention is discussed in detail below. The reactors used for the high pressure polymerization may be tubular reactors, autoclaves or combinations thereof, each having multiple reaction zones. The multiple reaction zones of the reactor allow for manipulation of the temperature profile throughout the polymerization process, which facilitates adjustment of product properties.

Ethylene monomer is fed to a primary compressor which compresses the monomer to an intermediate pressure, preferably a pressure of at least about 200bar (20MPa), and the monomer is mixed with recycled monomer to be fed to a secondary compressor. In some existing ethylene tubular reactor plants, the ethylene discharged from the primary compressor is split into two streams, one stream is combined with recycled monomer and fed to the suction of the secondary compressor, and the other stream is injected into the ethylene/polymer mixture downstream of the high pressure let down valve, thereby providing rapid cooling of the ethylene/polymer mixture before it enters the product separation system. In the process of the present invention, it is preferred that substantially the entire output of the primary compressor is fed to the secondary compressor. The discharge pressure of the primary compressor is matched to the pressure of the high pressure ethylene recycle system and may for example be 270bar (27MPa) to 350bar (35MPa) and preferably 280bar (28MPa) to 320bar (32 MPa). In addition, the ethylene monomer is preferably cooled after leaving the primary compressor and before entering the secondary compressor. In an advantageous embodiment, the primary compressor is a reciprocating compressor, having at least 8 cylinders, preferably 8-12 cylinders.

A secondary compressor downstream of and in fluid communication with the primary compressor further pressurizes the feed (including the feed discharged from the primary compressor) to a desired reactor pressure, which is greater than or equal to about 1500bar (150MPa), or greater than or equal to about 2000bar (200MPa), or greater than or equal to about 2500bar (250MPa), or greater than or equal to about 3000bar (300MPa), for feeding to the high pressure reactor. Interstage pressure (i.e., the pressure between the first and second stages of the two-stage compressor) is typically 1100bar (110MPa) to 1600bar (160 MPa). The two-stage compressor used with the tubular polymerization reactor is typically a two-stage reciprocating compressor having, for example, 6 or 8 cylinders, arranged in a compressor frame, and having a common crankshaft driven by an electric motor located at one end of the compressor frame. The temperature of the ethylene monomer should be controlled to allow load balancing between the two compressor stages, thereby optimizing/maximizing compressor throughput. Typically the ethylene is cooled between the first and second compression stages of the two-stage compressor. This can be done by passing the ethylene through a tube with a cooling jacket (usually a water jacket).

Other reaction components may be injected with the ethylene monomer (including one or more other comonomers) into the suction inlet of the secondary compressor. One or more modifiers may also be injected with the ethylene monomer and the comonomer(s), if available, into the suction inlet of the secondary compressor.

The compressed reactor feed stream exiting the secondary compressor may be split into two or more streams. At least one of the separate streams may be heated or cooled in one or more heat exchangers prior to entering the reactor. Other separate streams may be cooled in one or more coolers and introduced into the reactor at different points. The reactor has more than one reaction zone and may also include two or more initiator injection points along its length (if a tubular reactor) or in different zones (if an autoclave). The initiator may be supplied to the reactor in the polymerization initiator composition prepared in a mixing tank prior to being supplied to the polymerization reactor as described herein, or in situ if both the solvent and the initiator are supplied from separate storage tanks and mixed with each other in the line supplying them to the high pressure reactor. The monomer introduced into the front end of the reactor is heated to at least 95 c, preferably at least 135 c or in some cases at least 160 c to promote initiator decomposition and start the polymerization reaction.

In one class of embodiments, the initiator is injected into the reactor at more than one injection point (e.g., at least two, at least three, or at least four injection points). Preferably, the initiators are injected into different reaction zones in the same initiator composition. At least one injection pump is used to inject the initiator, and each injection pump is designed to inject at more than one injection point. As a result, rather than at least the same number of injection pumps as the number of injection points, fewer injection pumps are required to meet the required initiator supply for the polymerization apparatus and process described herein. Any suitable pump may be used for injecting the initiator, such as a hydraulically driven piston pump. The total output of the infusion pump is controlled by a pump discharge pressure controller. At least two flow control valves are used per injection point. The flow of initiator to the injection point is controlled by a temperature controller in the reaction zone acting on a flow control valve in the injection line leading to the reaction zone.

The polymerization reaction may be carried out in a high pressure reactor, wherein the monomers are polymerized, optionally with one or more comonomers, in the presence of a polymerization catalyst to form a product mixture comprising polymer and unreacted monomers. Suitable catalysts and catalyst systems are well known in the art. The polymerization is started immediately downstream of the first reaction zone, thereby causing the temperature of the reaction mixture to increase due to the exothermic nature of the polymerization. As the temperature increases, the initiator decomposes and the polymerization rate increases, which accelerates the generation of heat and causes further temperature increases. As the initiator is consumed, initiation and polymerization slows, and at a point where heat generation equals the heat conducted away from the reaction mixture, the temperature peaks and then begins to drop. Thus, as the reaction mixture travels along the length of the reactor, the reaction mixture temperature increases to a peak and then decreases until the next initiator injection point is reached, whereupon the process begins anew. The region downstream of the initiator injection point, where the polymerization reaction takes place, is referred to by the person skilled in the art as the reaction zone.

In embodiments where the ethylene discharged from the secondary compressor is divided into two or more streams, one stream enters the reactor front end and the other stream(s) enter as sidestream(s), which typically enters the reactor upstream of the initiator injection point, preferably after cooling to, for example, 10-20 ℃ prior to entering the reactor, to reduce the temperature of the reaction mixture. The overall conversion of monomer to polymer is in practice limited primarily by the ability to cool the reaction mixture, so cooling the side stream can allow for increased conversion for a given reactor.

In one embodiment, the high pressure polymerization conditions comprise a temperature of from about 120 ℃ to about 335 ℃. The peak temperature of each reaction zone will advantageously be in the range 200 ℃ to 350 ℃. Preferably, the peak temperature will be in the range of from 280 ℃ to 340 ℃, preferably from 290 ℃ to 315 ℃ in at least one reaction zone. The temperature increase in the reaction zone is proportional to the amount of polymer produced in the reactor zone and therefore operation at high peak temperatures favors high conversions. However, the ethylene polymerization kinetics are such that, as the temperature increases, the chain transfer to the polymer increases relative to the growth of the linear chain and the polydispersity index increases, which results in an increase in the haze value of the polymer produced. Therefore, when it is desired to make a low haze grade polymer, it will be necessary to run at a lower peak temperature. Preferably, in each reaction zone upstream of the initiator injection point (i.e. in all but not the last reaction zone), the reaction mixture is cooled to at least 20 ℃, more preferably at least 40 ℃ and most preferably at least 50 ℃ below the peak reaction zone temperature before the reaction mixture reaches the next initiator injection point.

In another embodiment, the high pressure polymerization conditions comprise a pressure of from about 1200bar (120MPa) to about 3500bar (350 MPa). The proportion of total ethylene entering the reactor, whether in the front-end stream or the sidestream, which is converted to polymer before exiting the reactor, is referred to as conversion. In the process of the invention, the conversion is at least 28%. The conversion achieved is in part related to the reactor operating pressure, and higher front end pressures increase the polymerization rate and make possible a greater pressure drop over the reactor length. However, operating at higher pressures places more strain on the secondary compressor and increases energy consumption, which results in cost disadvantages. For this reason, it may be desirable in some cases to operate at relatively low conversions (which may be, for example, about 28% -32%) at pressures of 2300bar (230MPa) to 2800bar (280 MPa). Alternatively, it may be desirable to operate at high conversion (e.g., in the range of 32% -37%) at pressures of 2800bar (280MPa) to 3100bar (310 MPa). However, the pressure is only one of the factors affecting the conversion, and 30% to 40% conversion is preferable as a whole, and the more preferable range is 30% to 37%.

Tubular reactor

The tubular reactor is a continuous, plug flow loop reactor. In preferred embodiments where the reactor is a tubular reactor, the tubular reactor typically has an initial section to which the monomer(s) are fed from a secondary compressor and in which they are heated to the desired reaction start temperature, typically to at least about 120 ℃, preferably at least about 135 ℃ or in some cases even to at least about 160 ℃. Once the desired temperature is reached, the polymerization initiator composition is injected to start the reaction. The pressure in the tubular reactor is generally in the range from about 2100bar (210MPa) to about 3500bar (350 MPa).

In the process of the present invention, the polymerization is carried out in a tubular reactor having at least three reaction zones, each of which is started at the point of initiator injection. The injected initiator decomposes into free radicals, which start the polymerization. Another point for injecting the initiator composition is located downstream along the length of the reactor. Preferably, the reactor has at least two, preferably at least three, more preferably at least four different injection points in total, thereby producing at least two, at least three or at least four reaction zones, respectively. In each reaction zone, the polymerization is carried out as described previously. In a preferred embodiment, the tubular reactor will generally be equipped with at least one temperature-regulated cooling jacket in each reaction zone. The reaction mixture in any reaction zone may be cooled by the cooling jacket (through which water or another cooling fluid is circulated) or a combination of cooling jackets, and a side stream of introduced cooled ethylene monomer.

The maximum internal diameter of the tubular reactor of the invention is preferably at least 65mm to maintain the pressure drop over the length of the reactor at an acceptable level. In embodiments where a portion of the ethylene discharged from the secondary compressor enters the tubular reactor as a side stream, it would be desirable for the reactor to have regions of different internal diameters, which increases the number of stages down the length of the reactor when the side stream enters. For example, for a process with a two-stage compressor throughput of about 160 metric tons/hour at 3000bar (300MPa), 20% of which enters the front end of the tubular reactor, and the remainder enters as a sidestream, the initial diameter of the tubular reactor may be 35mm to 40mm, and the internal diameter at the point of entry of the first sidestream will increase, depending on the size of the sidestream, etc., until the final internal diameter is 75mm to 80mm after the last sidestream. For any process according to the invention, the choice of the particular maximum internal diameter of the tubular reactor will depend on the throughput of the secondary compressor, the output pressure of the secondary compressor and the length of the tubular reactor used, all of which are related to the pressure drop experienced over the length of the reactor. The tubular reactor preferably has a length of from 1500m to 5000m, more preferably from 3000m to 4500 m.

In a process for polymerizing ethylene in a tubular reactor, once the desired throughput of ethylene through the secondary compressor and into the reactor is established, the reactor pressure is controlled by a high pressure let down valve through which the product mixture exits the reactor. Opening the valve reduces the pressure in the tubular reactor; closing the valve increases the pressure. Furthermore, there is a pressure drop along the length of the tubular reactor that drives the reaction mixture at a desired velocity along the reactor (the term "reactor pressure" refers herein to the maximum pressure in the reactor, i.e. the pressure immediately downstream of the secondary compressor, unless the context clearly indicates another meaning). The pressure drop over the length of the reactor depends on the condition that the pressure should not drop below the point at which phase separation of the reaction mixture takes place. For a given throughput, the pressure drop can be reduced by increasing the internal diameter of the tubular reactor. However, the increased tube diameter also makes it more difficult to effectively cool the reactor mixture.

High-pressure autoclave reactor

The high pressure polymerization process described herein may also be carried out in an autoclave reactor. The autoclave reactor is typically a cylindrical Continuous Stirred Tank Reactor (CSTR) with an agitator to promote good mixing and a residence time of about 20-60 seconds. Ethylene and optionally comonomer(s) are fed at one or more points into the reactor and then into the polymerization initiator composition. If desired, additional modifier may be added with the monomer feed, separately or as part of the polymerization initiator composition.

Furthermore, the autoclave reactor may have several reaction zones with different, increasing polymerization temperatures. The reaction zones may be separated from each other, for example by means of baffles within the reactor. In each such stage, an internally or externally operated stirrer provides back-mixing of the reaction mixture, but back-mixing between the stages is generally to be avoided. The pressure in the autoclave reactor is generally from about 1200bar (120MPa) to about 2100bar (210 MPa).

Product separation and recycle

The heated reactor (both tubular and autoclave) effluent contains polymer, unreacted monomer(s), residual transfer agent (if any) and residual organic solvent (initially used to dissolve the initiator). After leaving the reactor, the above-described reactor effluent (hereinafter "product mixture") may be passed through a downstream high-pressure let-down valve, which reduces the pressure of the product mixture so that the product mixture is no longer in a single phase and begins to form two phases, a monomer-rich phase with unreacted monomer(s) and a polymer-rich phase. The high pressure let down valve may be controlled to maintain a desired pressure in the reactor.

In many existing tubular reactor facilities, a portion of the ethylene discharged from the primary compressor is cooled and transferred in separate streams to a location immediately downstream of the high pressure blowdown valve (ethylene quench) to serve as a rapid quench cooling of the product mixture. However, it is preferred that all of the ethylene discharged from the primary compressor be directed to the secondary compressor and subsequently to the tubular reactor to maximize the amount of polymer produced and to provide an alternative means to cool the product mixture. The combination of high throughput and the above factors relating to the pressure drop over the length of the reactor make it undesirable to provide a greater cooling capacity in the final reaction zone than is required to control the exotherm of polymerization in the final reaction zone in the process of the present invention. Preferably, the temperature of the product mixture at the high pressure let down valve is 260-290 ℃. Thus, additional cooling means are provided downstream of the high pressure blowdown valve and upstream of the product separator.

From the high pressure blow down valve, the product mixture may flow through a jet pump and then into a separation system having at least two stages, which may include one or more high pressure separation ("HPS") vessels and one or more low pressure separation ("LPS") vessels. The first separation of polymer from unreacted ethylene is carried out in the first stage, where the product mixture is fed to an HPS vessel operating at a pressure lower than the exit pressure of the reaction zone, for example at least about 200bar (20MPa), preferably at least about 250bar (25 MPa). The separated gas is fed to a high pressure recirculation system for return to the secondary compressor. The polymer-rich liquid effluent of the HPS vessel is optionally fed to a medium-pressure separation (MPS, see WO2007/134670 and co-pending application PCT/US2008/087501) vessel, the operating pressure of which is between the HPS and the LPS vessel described below, and is therefore, for example, from about 10bar (1MPa) to about 250bar (25 MPa). After the MPS (or directly after the HPS) the polymer rich liquid is fed to the LPS vessel, at a working pressure lower than the working pressure of the MPS and HPS vessels, thus lower than about 20bar (2MPa), preferably lower than 10bar (1MPa) and most preferably lower than about 1bar (0.1 MPa). The separation between ethylene and optional monomers (e.g. vinyl acetate) takes place in the LPS vessel. The overhead from the LPS vessel is recycled back to the primary compressor. Part of the overhead gas of the LPS vessel is sent to the external battery limits for purification, thereby limiting the accumulation of impurities in the system. The molten polymer leaving the LPS vessel is then fed to a conventional extruder where it is combined with conventional additives to modify the properties of the extruded polymer, and then to a pelletizer for finishing.

Figure 1 shows an embodiment of a high pressure polymerization system 1 comprising a tubular reactor 5 with a conventional initiator injection scheme wherein an injection pump injects initiator into a reaction zone through an injection point. Ethylene feed line 2 supplies ethylene monomer to primary compressor 3, which pressurizes ethylene to a pressure of from about 200bar (20MPa) to about 350bar (35 MPa). The outlet of the primary compressor 3 is connected to the inlet of the secondary compressor 4 by a pipe with a valve. The secondary compressor 4 pressurises the reaction feed, which includes the ethylene feed and other reaction components discharged from the primary compressor 3, to a pressure of about 3000bar (300 MPa). The compressed reaction feed exiting the secondary compressor 4 may be split into two streams, one of which enters the front end of the pipe reactor 5 and the other of which is split into one or more sidestreams that enter the pipe reactor 5 at points along the length of the pipe reactor 5. The tubular reactor 5 has two reaction zones (5a and 5b) along its length and two initiator injection points, which are fed from an initiator injection system 6.

In the initiator injection system 6, one injection pump is used to inject the initiator into the reaction zone through one injection point, i.e. the injection pump 6a injects the initiator 6i into the reaction zone 5a and the injection pump 6b into the reaction zone 5 b. The output of each pump is controlled by a temperature controller (15a and 15b) of the reaction zone.

The mixture of polymer and unreacted monomer exiting the tubular reactor 5 is passed through a high pressure let down valve 7 which can be controlled to maintain the desired pressure in the tubular reactor 5. From the high pressure let down valve 7, the product mixture may flow through the jet pump and then into a separation system, which may include an HPS vessel 8 and an LPS vessel 9. The HPS vessel 8 can separate the product mixture into a stream of unreacted monomer gas 10 and a stream of polymer-rich liquid 11. The separated monomer gas may be directed to a recycle gas system 12. The recycle gas system 12 may include one or more waste heat boilers, one or more coolers for cooling the recycle gas, and one or more separation vessels for dewaxing. The cooled and dewaxed gas exiting the recycle gas system 12 may be returned to the reactor feed of the secondary compressor 4. The polymer rich liquid 11 may be further separated in the LPS container 9. The LPS reservoir 9 may be operated at a pressure of from about 0.5bar (0.05MPa) to about 2.0bar (0.2 MPa). The molten polymer exits the LPS container 9 via the bottom outlet of the LPS container 9 and is sent through a pipe to the inlet of a hot melt extruder 13, which hot melt extruder 13 extrudes the polymer into strands, chops it, cools and transfers it to a product storage bin (not shown). In the LPS vessel 9, at least a portion (if not all) of the remaining monomer is recovered as off-gas and compressed into the purge gas compressor 14. Any portion of the compressed purge gas may be recycled to the inlet of the primary compressor 3.

In contrast to fig. 1, fig. 2 shows an alternative embodiment of a high pressure polymerization system 1' which replaces the conventional initiator injection scheme of fig. 1 with the initiator injection scheme of the present invention. In the initiator injection system 6', an injection pump is designed to inject at more than one injection point. It can be seen that the initiator injection scheme of the present invention for the same tubular reactor line as in figure 1 can use only one injection pump 6 'a to regulate the initiator supply to both

reaction zones

5a and 5b by operating the injection pump 6' a to simultaneously inject initiator at two separate injection points of the

reaction zones

5a and 5 b. The total output of the charge pump 6' a is controlled by a pump discharge pressure controller 16. The flow of initiator to the injection point is controlled by temperature controllers (15 'a and 15' b) of the reaction zone, which act on flow control valves (17a and 17b) in the injection lines leading to the reaction zone. This design can increase initiator injection efficiency and result in advantageous cost savings for high pressure polymerization equipment.

Polymer product

The final polymer product (homopolyethylene or copolymer of ethylene and one or more of the above comonomers, preferably vinyl acetate) produced by the apparatus described herein or according to the processes described herein includes a wide range of Low Density Polyethylenes (LDPE) which can be made using free radical initiated high pressure processes, using tubular or autoclave reactors. Typically, the density ranges from about 0.910 to 0.935g/cm³A polydispersity index of from about 5 to about 50, a melt index of from about 0.1 to about 500g/min, and a haze value of from about 1 to 20. If desired, high comonomer contents, for example up to about 40% by weight of vinyl acetate, can be achieved. The molecular weight of the polymer can be varied by using different types and concentrations of modifiers. The polymer density can be affected by the type and amount of comonomer, as well as the polymerization temperature. In addition, haze can be affected by reactor temperature, as well as reactor pressure and choice of polymerization initiator. The main applications of these LDPE grades are film and extrusion coating.

An apparatus for high pressure polymerization of ethylene is also provided. An apparatus for high pressure polymerization of ethylene may comprise a high pressure reactor in which monomers are polymerized to form a product mixture comprising polymer and unreacted monomers, wherein the reactor has more than one reaction zone and at least one injection pump for injecting initiator into the reactor, and each injection pump injects at more than one injection point. Preferably, the apparatus further comprises: a primary compressor for compressing the monomer to an intermediate pressure, and a secondary compressor for compressing the monomer to a pressure of the reactor upstream of the reactor; a high pressure let down valve downstream of the reactor through which the product mixture from the high pressure reactor flows; a separation system downstream of the high pressure let down valve having at least two stages for separating the product mixture into polymer and unreacted monomer; and a recycling system for returning unreacted monomers to the secondary compressor. Preferably, the reactor is a tubular reactor having at least three reaction zones and at least one cooling jacket in each reaction zone.

All documents described herein are incorporated by reference in their entirety unless otherwise specified. When numerical lower limits and numerical upper limits are listed herein, ranges from any lower limit to any upper limit are contemplated. While forms of the invention have been illustrated and described, it will be apparent from the foregoing general description and specific embodiments that various changes can be made without departing from the spirit and scope of the invention. Accordingly, it is not intended that the invention be limited thereby.

Claims

1. An apparatus for high pressure polymerization of ethylene comprising a high pressure reactor in which monomers are polymerized to form a product mixture comprising polymer and unreacted monomers, wherein the reactor has more than one reaction zone and at least one injection pump fluidly connected to the reactor, wherein each injection pump is fluidly connected to the reactor at more than one injection point; wherein at least two flow control valves are used for each injection point; and wherein each injection pump is controlled by a temperature controller of the reaction zone.

2. The apparatus of claim 1, wherein the total output of each injection pump is controlled by a pump discharge pressure controller.

3. The apparatus of claim 1, wherein the reactor is a tubular reactor, an autoclave, or a combination thereof.

4. The apparatus of claim 1, wherein the reactor is a tubular reactor having at least three reaction zones and at least one cooling jacket in each reaction zone.

5. The apparatus of claim 4, wherein the maximum internal diameter of the tubular reactor is at least 65 mm.

6. The apparatus of any one of claims 1-5, further comprising:

a primary compressor for compressing the monomer to an intermediate pressure, and a secondary compressor for compressing the monomer to a pressure of the reactor upstream of the reactor;

a high pressure let down valve downstream of the reactor through which the product mixture from the high pressure reactor flows;

a separation system downstream of the high pressure let down valve having at least two stages for separating the product mixture into polymer and unreacted monomer; and

a recycle system for returning unreacted monomer to the secondary compressor.

7. A process for the high pressure polymerization of ethylene in the apparatus of any one of claims 1 to 6, comprising the steps of:

compressing the monomer to an intermediate pressure in a primary compressor, then mixing the monomer with recycled monomer, and further compressing the monomer to the pressure of the reactor upstream of the reactor in a secondary compressor;

introducing the monomer into a high pressure reactor having more than one reaction zone;

injecting initiator into the reactor at more than one injection point by at least one injection pump, wherein each injection pump injects at more than one injection point and at least two flow control valves are used for each injection point;

controlling the flow of initiator to the injection point with a temperature controller of the reaction zone acting on a flow control valve in an injection line leading to the reaction zone;

contacting the monomer in the reactor with the initiator under high pressure polymerization conditions to polymerize to form a product mixture;

discharging the product mixture through a high pressure let down valve, cooling the product mixture, and separating the product mixture into polymer and unreacted monomer in a separation system having at least two stages; and

the unreacted monomer is recycled to the secondary compressor.

8. The method of claim 7, further comprising controlling the total output of the injection pump using a pump discharge pressure controller.

9. The method of claim 7, further comprising injecting an initiator into the reactor at least 2, at least 3, or at least 4 injection points.

10. The method of claim 7, further comprising injecting initiators into different reaction zones having the same or different initiators.

11. The process of claim 7, wherein the reactor is a tubular reactor, an autoclave, or a combination thereof.

12. The process of claim 7 wherein the reactor is a tubular reactor having at least three reaction zones and at least one cooling jacket in each reaction zone.

13. The method of claim 7, wherein the initiator is oxygen or an organic free radical initiator.

14. The process of claim 7, wherein the high pressure polymerization conditions comprise a temperature of from 120 ℃ to 335 ℃ and a pressure of from 1200bar (120MPa) to 3500bar (350 MPa).

15. The method of claim 7, further comprising injecting at least one modifier with the monomer into a secondary compressor.

16. The process of claim 7 wherein ethylene is the only monomer present and the polymer is an ethylene homopolymer.

17. The method of claim 7, wherein one or more comonomers are present and the polymer is an ethylene copolymer.

18. The process of claim 17, wherein the comonomer is vinyl acetate, methyl acrylate, methacrylic acid, ethyl acrylate, butyl acrylate or acrylic acid or mixtures thereof.