CS214804B2 - A method for producing an extrudate from a rapidly melting ethylene polymer - Google Patents

Abstract

The invention relates to a method for producing an extrudate, preferably in the form of a film, from an ethylene polymer which melts rapidly in an extruder and exhibits slow weakening under shear stress, in which said polymer is extruded in an extruder with a screw having a length to diameter ratio in the range of 15:1 to 21:1.

Description

The invention relates to a method of producing an extrudate, preferably in the form of a film, from an ethylene polymer which melts rapidly in an extruder and exhibits slow weakening under shear stress, wherein said polymer is extruded in an extruder with a screw having a length to diameter ratio in the range of about 15:1 to 21:1.

The majority of industrially produced low-density polyethylene is produced by polymerization in thick-walled autoclaves or tubular reactors at pressures up to 344 MPa and temperatures up to 300 °C. The molecular structure of high-pressure low-density polyethylene is very complex. The permutations in the arrangement of its simple building blocks are essentially infinite. High-pressure resins are characterized by a complex branched architecture with long side chains. These long branches have a significant effect on the rheological properties of the resin melt. High-pressure low-density polyethylene resins also contain a spectrum of short side chains, generally containing 1 to 6 carbon atoms, which affect the crystallinity (density). The frequency distribution of these short branches is such that, on average, most chains contain the same number of branches. The distribution of short side chains, which is a characteristic feature of high-pressure low-density polyethylene, can be considered narrow.

Low-density polyethylene can exhibit many good properties. It is flexible and has a good balance of mechanical properties such as tensile strength, impact strength, burst strength and tear strength. In addition, it retains its strength down to relatively low temperatures. Some resins do not become brittle at temperatures down to -70 °C. Low-density polyethylene has good chemical resistance. It is relatively inert to acids and alkalis and inorganic solutions. However, it is sensitive to hydrocarbons, halogenated hydrocarbons and oils and fats. Low-density polyethylene has excellent dielectric strength.

More than 50% of all low-density polyethylene is processed into films. Films are used primarily in packaging technology, for example for packaging meat, agricultural crops, frozen foods, for the production of ice bags, bags that can be boiled, for the packaging of textile and paper products and goods that are hung on hangers, for the production of lining materials (for example for crates), transport bags and shrink films for packaging pallets with goods. Large quantities of thick films are consumed in the construction and agriculture industries.

Most low-density polyethylene films are produced by extrusion blow molding in the form of so-called tubes or sleeves. The diameters of the tubes produced vary. The narrowest tubes have a diameter of 5 cm or less and are used for the production of bags, and the largest tubes can have a diameter of up to 6 m (when such a tube is cut lengthwise, the width of the strip is 12 m).

Polyethylene can also be produced at low to medium pressure by homopolymerization of ethylene or by copolymerization of ethylene with various α-olefins using heterogeneous catalysts based on transition metal compounds of different valencies. These resins usually contain few, if any, long side chains and the only side chains are therefore short side chains. The length of the side chains depends on the type of co-conomer. The frequency of the side chains depends on the concentration of the comonomer or comonomers used during the copolymerization. The distribution of the frequency of the side chains depends on the type of transition metal catalyst used during the copolymerization. The distribution of short side chains, characteristic of low density polyethylene produced using transition metal catalysts, can be very broad.

U.S. Patent Application No. 892,325 filed March 21, 1978 and refiled July 27, 1979 under No. 014,414 (FJ Karol et al.: Preparation of Ethylene Copolymers in a Fluidized Bed Reactor) discloses that ethylene copolymers with a density of 910 to 960 kg/ ^m3 , a melt flow rate ratio (MFR) of 32 and MFR g 22, which have a relatively low content of catalyst residues, can be produced in the form of granules at relatively high productivity by polymerizing the monomer or monomers in a low-pressure gas phase process using a specific high-performance complex catalyst containing magnesium and titanium, mixed with an inert carrier.

U.S. Patent Application No. 892,322 filed March 31, 1978 and refiled February 16, 1979 under No. 012,720 (GL Goeke et al.: Impregnated Polymerization Catalyst, Process for Its Preparation and Its Use in the Copolymerization of Ethylene) discloses that ethylene copolymers having a density of 910 to 960 kg/ ^m3 , a melt flow rate ratio (MFR) of 32 to 22, and having a relatively low content of catalyst residues can be produced in the form of granules at relatively high productivity by polymerizing the monomer or monomers in a low-pressure gas phase process using a specific high-performance complex catalyst containing magnesium and titanium impregnated in a porous inert support.

U.S. Patent Application No. 892,037, filed March 31, 1978, and refiled February 27, 1979, under No. 014,412 (B. E. Wagner et al.: Polymerization Catalyst, Process for Its Preparation and Use in the Homopolymerization of Ethylene) discloses that ethylene homopolymers having a density of about 958 to 972 kg/ ^m3 (including the upper limits) and a melt flow rate ratio (MFR) of 32 to 22, having a relatively low catalyst residue content, can be produced in granulate form at relatively high productivity by homopolymerizing ethylene in a low-pressure gas phase process using a highly efficient magnesium-titanium complex catalyst mixed with an inert carrier. The granulated pointers thus produced are suitable for numerous end applications.

Polymers made according to these U.S. patent applications (hereinafter referred to as the above-referenced earlier U.S. applications) have the property of melting rapidly and exhibiting slow weakening when subjected to shear stress when used in an extruder.

Polymers such as those produced according to the cited US applications using complex catalysts containing magnesium and titanium have a narrow molecular weight distribution expressed as a ratio Mw/Mn of 2.7 g Mw/Mn g 3.6 and preferably 2.8 g Mw/Mn g 3.4.

The rheological properties of polymeric substances depend to a large extent on their molecular weight and molecular weight distribution. In film extrusion, two aspects of rheological behavior are important: shear stress and stretching. In the extruder and extrusion die, the polymer melt undergoes severe deformation due to shear stress. As the melt is pumped by the extruder screw to the die and forming die, the melt is found in a wide range of shear strain rates. It is assumed that in most film extrusion processes, the melt is subjected to shear at a strain rate of 100 to *5000 s ^{_ l} . Polymer melts, as is known, exhibit weakening under shear stress, i.e. non-Newtonian flow properties. As the shear rate increases, the viscosity (the ratio of shear stress τ to shear rate y'j) decreases. The degree of viscosity decrease depends on the molecular weight, molecular weight distribution, and molecular conformation, i.e., the long side chains of the polymer material. Short side chains have little effect on the injection viscosity. High-pressure low-density polyethylenes typically have a broad molecular weight distribution and exhibit increased shear weakening over the range of shear rates common in film extrusion. The narrow molecular weight distribution resins of the invention exhibit reduced shear weakening over the shear rates common in extrusion. As a result of these differences, narrow molecular weight distribution resins require more energy and higher extrusion pressures than low-density, broad molecular weight distribution high-pressure polyethylene resins. which have the same average molecular weight.

The rheological properties of polymeric materials are usually studied using shear deformation. Under pure shear stress, the velocity gradient of the deforming resin is perpendicular to the flow direction. This deformation method is experimentally convenient, but does not provide important information needed to understand the material response in film production.

Just as shear viscosity can be defined using shear stress and shear strain rate

7jstrih — τ12/χ (lj where strih represents the shear viscosity in Pa.s, τΐ2 represents the shear stress in Pa and y is the shear strain rate vs ^_1 , The elongation viscosity can be defined using the normal stress and the strain rate under tensile stress by the formula 2 rjext — τΐΐ/ε (2) where

Tjext represents the extensional viscosity in Pa.s, τΐΐ represents the normal stress in Pa, and ε represents the strain rate under tensile stress vs ^_1 .

In pure longitudinal flow, unlike shear flow, the velocity gradient is parallel to the direction of flow. Industrial extrusion processes involve both shear and tensile deformations. In film extrusion (slot extrusion and tube blowing), the rheological properties of the resin during stretching are of utmost importance. They can actually be decisive in the process.

The extensional viscosity can be measured by numerous experimental techniques (see, for example, J. L. White, Report No. 104 of the Polymer Science and Engineering., Univ. of Tenn., Knoxvillej. The measurement described in these paragraphs uses the constant stress method. The measurement is made on an Instron tensile testing machine with servo control. The ends of a molten ring of polymer immersed in a silicone oil bath are withdrawn at increasing rates according to the following relationship 3

L(t) = L _o exp (ε tj where

Lt is the distance between the jaws at time t

L _o is the initial distance between the jaws, ε is the tensile strain rate (s~'j — constant, t is time.

The strain is measured by a force transducer. The elongational viscosity is calculated by dividing the stress by the rate of tensile strain and is determined as a function of displacement or time during deformation (temperature ~150 °C).

When high-pressure low-density polyethylene is subjected to deformation according to equation 3, it is found that the elongational viscosity increases exponentially with the logarithm of time. This behavior is illustrated in Fig. 1 for the case of high-pressure low-density polyethylene having a melt flow index of 0.65 g/10 min and a density of 922 kg/m ³ . This property is referred to as melt strain hardening. Strain hardening intensifies with increasing deformation rate. In some cases, the melt may exhibit an unlimited increase in stress.

Ethylene polymers produced by the processes of the above-mentioned U.S. applications do not generally exhibit unlimited stress growth. Some resins with a broad molecular weight distribution do stiffen under strain, but their viscosity during stretching increases linearly with the logarithm of time (see Fig. 2j). Certain resins with a narrow molecular weight distribution described in the above-mentioned U.S. applications exhibit little stiffening under strain when the strain rate is low. Fig. 3 shows that stiffening under strain increases at higher strain rates, but not to the extent observed with high-pressure low-density polyethylene or with copolymers of ethylene and hydrocarbons with a broad molecular weight distribution.

High-pressure low-density polyethylene can be considered soft in shear and stiff in extension compared to ethylene-hydrocarbon copolymers with narrow molecular weight distribution. Ethylene-hydrocarbon copolymers with narrow molecular weight distribution exhibit opposite rheological properties. They are stiff in shear and soft in extension. The terms "soft" and "stiff" as used herein refer to the relative magnitude of the shear viscosity and extension viscosity when comparing the rheology of high-pressure low-density polyethylene and the ethylene polymers with narrow molecular weight distribution of the invention.

High-pressure low-density polyethylene is usually processed into films by melt extrusion in an extruder with a screw having a length-to-diameter ratio greater than 21:1. In the industrial production of high-pressure low-density polyethylene films, the extrusion screw has a length-to-diameter ratio (L/D) of 24:1 or greater. These long extrusion screws use mixing devices to achieve maximum production speed and acceptable film quality (homogeneous temperature distribution in the melt). When an extrusion screw with a length-to-diameter ratio of 15:1 to 21:1 is used to extrude a high-pressure low-density polyethylene melt, the film quality is not acceptable due to the uneven temperature distribution. The film quality can be improved by using an extrusion screw with a length-to-diameter ratio of 15:1 to 21:1 if the extrusion screw is cooled. However, cooling the extrusion screw is not suitable from an industrial point of view, as it greatly reduces the production speed and increases the energy consumption per kilogram of resin. This can be illustrated by the data given in the following table (for the case of film production by blowing extruded tubes):

Table

Extrusion screw (L/D) Foil quality ³ ' Maximum •speed of extruder at full speed Energy consumption Consumption Cooling Type polyethyenergy limited to low •per 1 kg speed density 20 : 1 or >20 : 1 + 1.0 1.0 1.00 •1.0 high pressure 15:1 — 21:1 — 0.80 0.70 0.88 1.1 high pressure 15:1 — 21:1 + 0.50 0.80 1.60 1.0 high pressure (cooling) 15 : 1 — 21 : l ^b ' + 0.95 1.0 1.00 1.0 low pressure

aj Film quality marked — means that there is an obvious unevenness in the temperature distribution on the film bj For low density resins described in the substrates, extrusion screws with an L/D ratio of 15:1 to 21:1 provide satisfactory melt quality, defined by pressure variations in the extruder head.

The unique rheological behavior of low molecular weight distribution polyethylene resins produced, for example, by the processes of the above-mentioned US applications is manifested in the film forming by blowing the extruded tube in several ways. When these resins are processed in commercially available equipment, i.e. in an extruder with a screw having a length ratio greater than 21:1, especially 24:1, the energy consumption is high and the melt temperature is also high. This causes the production rate to be limited either by the power of the extruder or by the cooling of the blown bag to a value that is lower than in the case of processing high-pressure low density polyethylene resins.

However, when an extruder for extruding films from these narrow molecular weight distribution ethylene polymers uses an extruder screw with a length to diameter ratio ranging from about 15:1 to 21:1, the energy consumption and melt temperature are reduced and a film of good quality is obtained.

Films suitable for packaging applications must have a balance of certain key properties to meet the operational requirements important for wide application by the end customer. These properties include optical properties such as haze, gloss and transparency, and mechanical properties such as tear resistance, tensile strength, impact strength, stiffness and tear strength. When packaging perishable products, water vapour and gas permeability are important properties. Film properties such as coefficient of friction, adhesion, heat sealability and resistance to bending affect the performance of packaging and similar machines. High-pressure low-density polyethylene is widely used for the production of packaging in both the food and non-food industries. Low-density polyethylene is commonly used to make bags, suitable for use as transport bags, textile packaging, laundries and dry cleaners' packaging and garbage bags. Low-density polyethylene film can be used to line drums for various liquid and solid chemicals and as a filling material in wooden crates. Low-density polyethylene film can be used in a number of agricultural and horticultural applications, such as in plant and crop protection, such as silencing and for storing fruit and vegetables. Low-density polyethylene film can also be used in construction, as a barrier against moisture or water vapor. Low-density polyethylene film can also be coated and printed and can be used in the production of newspapers, books, etc.

Due to the unique combination of the above properties, films made of high-pressure low-density polyethylene are the most important thermoplastic packaging films. In packaging technology, they account for about 50% of the total consumption of these films. Films made of ethylene polymers, especially copolymers of ethylene and hydrocarbons according to the present invention have an improved combination of user properties and are particularly suitable for many of the applications where high-pressure low-density polyethylene has been used until now.

Improving any of the film properties or improving the extrusion resin properties or improving the film extrusion process itself is of great importance in many end-use applications, where such films can be used to replace low-density high-pressure polyethylene films.

Figures 1, 2 and 3 show a plot of the extensional viscosity versus logarithm of time for three types of low density polyethylene.

Figure 4 shows a fluidized bed reactor in which ethylene-hydrocarbon copolymers can be produced.

Fig. 5 shows a diagram of the cylinder of an extruder with an extrusion screw according to the invention.

Figures 8 and 7 illustrate the slow weakening of two low-pressure ethylene polymers of the invention under shear stress compared to two high-pressure polyethylene resins of comparable density and melt flow index.

According to the invention, it has been found that in the production of an extrudate, preferably in the form of a film, from an ethylene polymer which melts rapidly in the extruder and exhibits slow shear weakening, by extruding this polymer from the melt using an extruder whose screw has a length to diameter ratio in the range of about 15:1 to 21:1, the energy consumption and the melt temperature of the polymer are reduced compared to the production of the extrudate using an extruder screw with a length to diameter ratio greater than 21:1.

Using an extruder screw with a length to diameter ratio in the range of about 15:1 to 21:1 to extrude these ethylene polymers from the melt also results in higher machine productivity compared to a process using an extruder screw having a length to diameter ratio greater than 21:1. In addition, using an extruder screw with a length to diameter ratio in the range of about 15:1 to 21:1 provides an acceptable, homogeneous temperature distribution.

As ethylene polymers, ethylene homopolymers or copolymers containing a predominant molar proportion (i 90%) of ethylene and a minor molar proportion (slO°/o) of one or more C ₃ to C ₈ α-olefins, which should not contain any branches on any of their carbon atoms closer to the main chain of the copolymers than the fourth carbon atom, can be used in the process according to the invention. Preferred α-olefins are propylene, 1-butene, 1-hexene, 4-methyl-1-pentene and 1-octene.

Ethylene polymers have a melt flow ratio of 18 to 32, preferably from 22 to 32, inclusive. The melt flow ratio is another way of characterizing the molecular weight distribution of a polymer. A melt flow ratio (MFR) value of 22 to 32 (inclusive) corresponds to a Mw/Mn value in the range of about 2.7 to 4.1. The polymers used in the invention have a Mw/Mn value in the range of about 2.2 to 4.1.

The homopolymers have a density of from about 958 to 972, preferably from 961 to 968 kg/m ³ , inclusive.

The ethylene copolymers have a density of about 910 to 960, preferably 917 to 955 kg/m ³ and most preferably about 917 to 935 kg/m ³ , inclusive. The density of the copolymer can be controlled, for a given melt flow index, primarily by the amount of C ₃ to C ₈ comonomers to be copolymerized with the ethylene. In the presence of the comonomer, ethylene would homopolymerize with the catalyst of the present invention to form homopolymers having a density of about 0.96. The introduction of increasing amounts of comonomers into the polymers results in a gradual decrease in the density of the copolymer. Under the same reaction conditions, the amount of each of the different C ₃ to C ₈ comonomers required to achieve the same result varies from comonomer to comonomer.

To achieve the same results, i.e. the same density at a certain melt flow index, increasing molar amounts of comonomers are required in the series C ₃ > C ₄ > C ₃ >>> C ₇ > Cg.

The melt flow index of a homopolymer or copolymer is a reflection of its molecular weight. Polymers with relatively high molecular weight have relatively low melt flow indices. Ultra-high molecular weight ethylene polymers have a higher load melt flow index (HLMI) of about 0.0, and very high molecular weight polymers have a higher load melt flow index of about 0.0 to about 1.0 g/10 min. The polymers of the invention have a normal load melt index (MI) in the range of from about 0.0 to about 50, preferably in the range of about 0.5 to 35, and a high load melt index (HLMI) in the range of about 11 to about 950. The melt index of the polymers used in the process of the invention is dependent on a combination of the reaction temperature during polymerization, the density of the copolymers, and the hydrogen/monomer ratio in the reaction system. The melt index increases with increasing polymerization temperature and/or decreasing polymer density and/or increasing hydrogen/monomer ratio.

The polymers contain 1 and usually 0.1 to 0.3 C--C double bonds per 1000 carbon atoms and have a cyclohexane extractable component content of less than about 3 and preferably less than about 2% by weight.

The polymers contain as catalyst residues on the order of 20 ppm Ti inclusive at a productivity of 50,000 or more and 10 ppm Ti inclusive at a productivity of 100,000 or more and 3 ppm inclusive at a productivity of 300,000 or more. If the polymers have been produced using halogen-containing catalysts, where the halogen is chlorine, they have a chlorine residue content of 140 ppm inclusive at a productivity of 50,000 or more, 70 ppm inclusive at a productivity of 100,000 or more and 21 ppm at a productivity of 300,000 or more. Ethylene polymers are readily produced at productivity up to about 300,000.

The copolymers are granules of material having a mean particle size on the order of about 0.127 to about 1.524 mm, and preferably about 0.5 to about 1 mm. The particle size is important for ease of fluidization of the polymer particles in a fluidized bed reactor, as discussed below. The copolymers have a bulk density after settling of about 243.2 to about 518.7 kg/m ³ .

Homopolymers and copolymers are suitable for the production of films.

For the production of films, those copolymers are preferably used which have a density of about 917 to 924, inclusive, a molecular weight distribution of 2.7 g Mw//Mn g 3.6, preferably 2.8 g Mw/Mn g 3.1 and a standard melt flow index (MFI) of 0.5 < < MFI g 5.0, preferably about 1.0 g MFI g 4.0. The films have a thickness of from 0 to 0.25 mm, preferably up to 0.12 mm.

The compounds used for the preparation of a highly active catalyst, which is used for the production of polymers, include at least one titanium compound, at least one magnesium compound, at least one electron donor compound, at least one activating compound and at least one inert substance, such as a carrier. The individual components are defined below.

The titanium compound has a structure corresponding to the general formula

Ti(OR) _a X _b where

R represents a _C1 to _C14 aliphatic or aromatic hydrocarbon radical or a radical of the general formula COR', where R' represents a C1 to C14 aliphatic or aromatic hydrocarbon radical,

X represents chlorine, bromine, iodine or a mixture thereof, a represents the number 0 or 1, b represents the number 2 to 4, inclusive, with the sum of a and b having the value 3 or 4.

Titanium compounds can be used individually or in mixtures and include compounds of the formulas

T1C1 ₃ , TiCl ₄ , Ti(OCH ₃ )Cl ₃ , Ti(OC ₆ H ₅ )Cl ₃ , Ti(OCOCH ₃ )Cl ₃ and Ti(OCOC ₆ H ₅ )C1 ₃ .

The magnesium compound has a structure corresponding to the general formula

MgX ₂ where

X represents chlorine, bromine or iodine.

These magnesium compounds can be used individually or in combinations and include compounds of the formulae MgCl ₂ , MgBr ₂ and Mgl ₂ . Of the magnesium compounds, anhydrous magnesium chloride is particularly preferred.

In the preparation of catalysts, about 0.5 to 56, preferably 1 to 10, moles of magnesium compound are used per 1 mole of titanium compound.

The titanium compounds and the magnesium compounds are advantageously used in a form that facilitates their dissolution in the electron donor compound, as described below.

An electron donor compound is an organic compound that is liquid at 25 °C and in which a titanium compound and a magnesium compound are partially or completely soluble. Electron donor compounds are known either by this name or by the name. Lewls bases.

The term electron donor compounds refers to compounds such as alkyl esters of aliphatic and aromatic carboxylic acids, aliphatic ethers, cyclic ethers and aliphatic ketones. Of these electron donor compounds, preferred are alkyl esters of C1 to _C4 saturated aliphatic carboxylic acids, alkyl esters of _C7 to _C8 aromatic carboxylic acids, _C2 to _C8 and preferably _C3 to _C4 aliphatic ethers, _C3 to _C4 cyclic ethers and preferably _C4 cyclic mono- or di-ethers, _C3 to _C1 and preferably _C3 to _C4 aliphatic ketones. The most preferred of these electron donor compounds are methyl formate, ethyl acetate, butyl acetate, ethyl ether, hexyl ether, tetrahydrofuran, dioxane, acetone and methyl isobutyl ketone.

Electron donor compounds can be used individually or in mixtures.

About 2 to 85 and preferably about 3 to 10 moles of electron-withdrawing compound are used per mole of titanium.

The activating compound has a structure corresponding to the general formula

Al(R“) _c X _d 'H _c where

X* represents chlorine or a group of the general formula OR"' and each of the symbols

R" and R"', which may be the same or different, represent a C ₁ to C ₁₄ saturated hydrocarbon radical, d represents a number from 0 to 1.5, e represents a number 1 or 0, with the sum of c + d + e being 3.

The activating compounds can be used individually or in combinations and examples thereof include substances of the formula

A1(C ₂ H ₅ ) ₃ , A1(C ₂ H ₅ ) ₂ C1, A1(í-C ₄ H ₉ ) ₃ , A1 ₂ (C ₂ H ₅ ) ₃ C1 ₃ , A1(í-C ₄ H ₉ ) ₂ H, A1(C ₆ H ₁₃ ) _3i A1(C ₈ H ₁₇ ) ₃ , A1(C ₂ H _s ) ₂ H a

A1( _C2H5 ) _{2 ( OC2H5} ) _.

In activating the catalysts, about 10 to 400 and preferably about 10 to 100 moles of activating compound are used per 1 mole of titanium compound.

The materials used as supports are particulate solids which are inert to the other components of the catalyst composition and to the other active components of the reaction system. These supports include inorganic materials such as silica and alumina and molecular sieves and organic materials such as polyolefins, for example polyethylene. The supports are used in the form of dry powders having an average particle size of about 10 to 250, preferably about 50 to 150 μΐη. These materials are preferably also porous and have a specific surface area of at least 3 and preferably at least 50 m ² /g. The support should be dry, i.e. free of absorbed water. This is normally done by heating or pre-drying the support with a dry inert gas before use. The inorganic support can also be further activated by treating it with about 1 to 8% by weight of one or more of the alkylaluminum compounds described above.

The catalyst is prepared by first preparing a precursor from a titanium compound, a magnesium compound and an electron donor compound as described below. The support may then be impregnated with the precursor and then treated with an activating compound in one or more steps as described below. Alternatively, the precursor may be treated with a support and an activating compound in one or more steps as described below.

The precursor is prepared by dissolving the titanium compound and the magnesium compound in the electron donor compound at a temperature of from about 20°C to about the boiling point of the electron donor compound. The titanium compound may be added to the electron donor compound before, after, or simultaneously with the addition of the magnesium compound. Dissolution of the titanium compound and the magnesium compound may be aided by stirring and, in some cases, by heating the two compounds in the electron donor compound to reflux. After dissolution of the titanium compound and the magnesium compound, the precursor is isolated by crystallization or precipitation with a C ₅ to C ₈ aliphatic or aromatic hydrocarbon, such as hexane, isopentane, or benzene.

The crystallized or precipitated precursor is isolated in the form of fine, freely moving particles with an average size of about 10 to 100 microns and a bulk density of about 292 to about 535 kg/m ³ (after settling).

The precursor prepared in this way has a composition corresponding to the general formula

Mg _m Ti ₁ (ORj _n X _p [EDj _q where

ED represents an electron donor compound, m represents a number from 0.5 to 56, preferably from 1.5 to 5, inclusive, n represents a number 0, 1 or 2, p represents a number from 2 to 116, preferably from 6 to 14, inclusive, q represents a number from 2 to 85, preferably from 4 to 11, inclusive,

R represents a C1 to C4 aliphatic or aromatic hydrocarbon radical or a radical of the general formula COR ¹ , where R' represents a C1 to C4 aliphatic or aromatic hydrocarbon radical and

X represents chlorine, bromine or iodine or a mixture thereof.

The precursor can then be impregnated into the carrier in a weight ratio corresponding to about 0.033 to 1, and preferably 0.1 to 0.33, parts of precursor per part of carrier.

The impregnation of the dried (activated) carrier with the precursor can be carried out by dissolving the precursor in an electron donor compound and then mixing the resulting solution with the carrier. When the carrier is impregnated with the precursor, the solvent is evaporated by drying at a temperature of up to 70 °C.

The carrier can also be impregnated with the precursor by adding it to the solution of the raw materials used for the preparation of the precursor in the electron donor compound. In this case, the precursor is not isolated from the solution. The excess electron donor compound is then removed by drying at a temperature of up to 70 °C.

Alternatively, the precursor may be diluted with a carrier.

Dilution of the precursor can be carried out before the precursor is partially or fully activated as described below, or simultaneously with such activation. Dilution of the precursor is carried out by mechanical mixing or mixing about 0.033 to 1, and preferably about 0.1 to 0.33 parts of the precursor with 1 part by weight of the carrier.

In order to be able to use it, the precursor must be fully or partially activated, i.e. it must be treated with a sufficient amount of an activating compound to convert the titanium atoms contained in the precursor into an active state.

In preparing a useful catalyst, it is necessary to carry out the activation so that at least the last stage of activation takes place in the absence of solvent, so that the fully active catalyst does not have to be dried to remove the solvent. Two activation procedures are described below, namely activation of the impregnated precursor (procedure A) and activation of the precursor diluted with a carrier.

In process A), the activation is carried out in at least two stages. In the first stage, the precursor impregnated in silica is partially activated by partial reduction with such an amount of activating compound as to obtain a partially activated precursor in which the molar ratio of activator to titanium is about up to 10:1 and preferably about 4 to 8:1. This partial activation is preferably carried out in a suspension of a hydrocarbon solvent. The resulting product is then dried to remove the solvent at a temperature of 20 to 80, preferably 50 to 70°C. In this partial activation, the activating compound can be used in absorbed form on the carrier used as a carrier for the precursor. The resulting product has the character of a freely mobile substance in the form of particles and can be easily introduced into the polymerization reactor. The partially activated impregnated precursor is preferably weakly effective as a polymerization catalyst in the process according to the invention. In order to activate the partially activated precursor for the polymerization of ethylene, an additional activator must be introduced into the polymerization reactor. This completes the activation of the precursor in the reactor. The additional activating compound and the partially activated impregnated precursor are preferably introduced into the reactor via separate lines. The additional activating compound can be injected into the reactor in the form of a solution in a hydrocarbon solvent, such as isopentane, hexane or mineral oil. This solution usually contains about 2 to 30% by weight of the activating compound. The additional activating compound is added to the reactor in such an amount that, together with the amount of activating compound and titanium compound introduced in the form of partially activated precursor, an overall Al/Ti molar ratio of about 10 to 400, and preferably about 15 to 60, is achieved in the reactor. The additional amount of activating compound introduced into the reactor reacts with the partially activated impregnated precursor and completes the activation of the titanium compound directly in the reactor.

For procedure Bj, i.e. activation of the precursor diluted with a carrier, two variants were developed:

According to the first variant, the precursor is completely activated outside the reactor in the absence of solvent by dry mixing the precursor with the activating compound. In this dry mixing, the activating compound is preferably used in an impregnated state in a carrier. However, this process has the disadvantage that the resulting fully activated catalyst is pyrophoric when it contains more than 10% by weight of the activating compound.

According to the second preferred variant, the precursor is partially activated outside the polymerization reactor. This partial activation is carried out in a suspension of a hydrocarbon solvent. The resulting product is then dried to remove the solvent. The partially activated precursor is introduced into the polymerization reactor, where the activation is completed with an additional activator.

In the dry-mixing catalyst preparation, a solid precursor in particulate form is uniformly mixed with solid particles of a porous carrier in which the activating compound is absorbed. The activating compound is absorbed into the carrier from its solution in a hydrocarbon solvent so that 10 to 50% by weight of the activator is deposited in 90 to 50% by weight of the carrier. The amounts of precursor, activator and carrier are selected to achieve the desired Al/Ti molar ratio and to obtain a final catalyst with a precursor to carrier weight ratio of less than about 0.50, preferably less than about 0.33. This amount of carrier provides the necessary dilution of the activated catalyst to allow the desired control of the polymerization activity of the catalyst in the reactor. When the final catalyst contains more than 10% by weight of the activator, it is pyrophoric. During dry mixing, which can be carried out at ambient temperature (25°C) or lower, the dry mixture is well mixed to prevent heat build-up during the subsequent activation, which is initially exothermic. This results in a catalyst that is fully reduced and activated and can be used as such in the polymerization reactor. It has the nature of a free-flowing (free-flowing) substance.

According to a second preferred process variant, the catalyst is activated in at least two stages. In the first stage, the solid precursor in particulate form diluted with a carrier is partially activated by reaction with such an amount of activating compound as to obtain a partially activated precursor in which the molar ratio of activator to titanium is about 1 to 10:1 and preferably about 4 to 8:

: 1. This partial reduction is preferably carried out in a suspension of a hydrocarbon solvent. The resulting product is then dried to remove the solvent at a temperature of 20 to 80, preferably 50 to 70°C. In this partial activation, the activating compound can be used in absorbed form on the carrier used to dilute the activating compound. The resulting product has the character of a freely mobile particulate substance and can be easily introduced into the polymerization reactor. The partially activated precursor is preferably weakly effective as a polymerization catalyst in the process according to the invention. In order to activate the partially activated precursor for the polymerization of ethylene, an additional activator must be introduced into the polymerization reactor. This completes the activation of the precursor in the reactor. The additional activating compound and the partially activated precursor are preferably introduced into the reactor via separate lines. The additional activating compound can be injected into the reactor in the form of a solution in a hydrocarbon solvent, such as isopentane, hexane or mineral oil. This solution usually contains about 2 to 30% by weight of the activating compound. The activator can also be added to the reactor in solid form, absorbed in a carrier. The carrier usually contains 10 to 50% by weight of the activator. The additional activating compound is added to the reactor in such an amount that, together with the amount of activating compound and the titanium compound introduced in the form of a partially activated precursor, a total molar ratio of Al/Ti of about 10 to 400 and preferably about 15 to 60 is achieved in the reactor. The additional amount of activating compound introduced into the reactor reacts with the partially activated precursor and completes the activation of the titanium compound directly in the reactor.

In a continuous gas phase process, such as the fluidized bed process described below, individual discrete doses of partially or fully activated precursor or discrete doses of partially activated precursor impregnated in a carrier are continuously fed into the reactor during the polymerization, optionally with the simultaneous introduction of discrete doses of additional activating compound required to complete the activation of the partially activated precursor, thus replacing the active sites of the catalyst that have been consumed during the reaction.

The polymerization is carried out by contacting a stream of ionomers in the gas phase, for example by the fluidized bed process described below, substantially in the absence of catalytic poisons such as moisture, oxygen, carbon monoxide, carbon dioxide and acetylene, with a catalytically effective amount of a fully activated precursor (catalyst), which is optionally impregnated on a support, at a pressure and temperature sufficient to initiate the polymerization.

In order to achieve the desired density range of the copolymers, it is necessary to copolymerize with ethylene a sufficient amount of comonomers having a carbon chain of 3 or more carbon atoms such that the content of C; ₅ to _C8 comonomers in the copolymers is 0 to 10 mole %. The amount of comonomer required to achieve this result depends on the particular comonomer or comonomers used.

The following table lists several different comonomers that must be used for copolymerization with ethylene in order to produce polymers having a density within the desired range at any given melt flow index. The table lists both the molar amounts of these comonomers in the polymer and the relative molar ratio of these comonomers to ethylene that must be maintained in the recycle monomer gas stream under equilibrium reaction conditions in the reactor.

Propylene comonomer

1-butene

1-penten

1-hexene

4-methyl-1-penten

1-heptene

1-octene

Molar ratio of comonomer//ethylene in the gas phase under equilibrium conditions

0.2 to 0.9 0.1 to 0.5 0.05 to 0.2 0.02 to 0.15 0.02 to 0.15 0.017 to 0.10 0.015 to 0.03

Molar ratio of comonomer//ethylene in the polymer

0.01 to 0.09 0.006 to 0.08 0.005 to 0.07 0.004 to 0.06 0.004 to 0.06 0.003 to 0.04 0.003 to 0.04

In order to maintain the above ratio during the reaction, the homomers are introduced into the reactor in a higher ratio than their required concentration in the gas phase.

A fluidized bed reaction system that can be used is illustrated in Figure 4. Reactor 10 consists of a reaction zone 12 and a rate-reducing zone 14.

The reaction zone 12 contains a bed of growing polymer particles, formed polymer particles, and a minor amount of catalyst fluidized by a continuous flow of polymerizable and modifying gaseous components formed by make-up gas and gas recycled to the reaction zone. In order to maintain the fluidized bed in a fluidized state, the mass flow rate of the gas through the bed must be greater than the minimum flow rate required for fluidization. Preferably, the mass flow rate of the gas through the bed is 1.5 to about 10 times greater, and preferably about 3 to 6 times greater, than the minimum mass flow rate of the gas required for fluidization (Gmf). The designation Gmf for the minimum mass flow rate of the gas required to achieve fluidization was proposed by C. Y. Wen and Y. H. Yu, "Mechanics of Fluidization", Chemical Engineering Progress Symposium Series, vol. 62, pp. 100-111 (1966).

It is important that the bed always contains particles to prevent the formation of "hot spots" and to trap and distribute the powdered catalyst in the reaction zone. At start-up, polymer particles are usually introduced into the reaction zone as a base before the gas flow is introduced. These particles may be identical to the polymer being produced or they may be different. If they are different, they are withdrawn with the polymer particles as the leading product. Finally, the fluidized bed of particles of the desired polymer displaces the starting bed.

The partially or fully activated precursor (catalyst) used in the fluidized bed is preferably stored in a standby tank 32 under an atmosphere of an inert gas, such as nitrogen or argon.

Fluidization is achieved by a high velocity of the gaseous recycle introduced into and through the bed, which is typically on the order of about 50 times higher than the velocity of the make-up gas. A fluidized bed usually appears as a dense mass of moving particles flowing in a free vortex created by the passage of the gas through the bed. The pressure drop across the bed is equal to or slightly higher than the weight of the bed divided by the cross-section of the bed, and is therefore dependent on the geometry of the reactor.

The make-up gas is introduced into the bed at the same rate as the rate of withdrawal of the polymer product particles. The composition of the make-up gas is determined by means of an analyzer 16 located above the bed. The gas analyzer determines the loss of components in the gas being recycled and the make-up gas composition is adjusted accordingly to maintain the composition of the gas mixture in the reaction zone at a substantially steady state.

To ensure complete fluidization, recycle gas and optionally also a portion of the make-up gas are introduced into the reactor through a feed 18 below the bed. A gas distribution plate 20 is located above the feed 18 to assist in fluidizing the bed.

The portion of the gas stream that does not react in the bed forms the recycle gas, which is removed from the polymerization zone preferably by passing it through a zone 14 located above the bed, in which the velocity is reduced and the entrained particles fall back into the bed. Particle separation can be assisted by the inclusion of a cyclone 22, which can either be part of the velocity reducing zone 14 or be external to this zone. If necessary, the recycle gas can then be passed through a filter 24, which is adapted to separate small particles at high flow rates and - in order to prevent - dust deposits on the heat transfer surfaces and on the compressor blades.

The recycle gas is then compressed in compressor 25 and passed through heat exchanger 26 where it is stripped of the heat of reaction before being returned to the bed. By constantly removing the heat of reaction, there is no noticeable temperature gradient at the top of the bed. A temperature gradient exists in the lower 152 to 304 mm of the bed, between the temperature of the entering gas and the temperature of the rest of the bed. It has been observed that the bed almost immediately adjusts the temperature of the recycle gas above this 152 to 304 mm zone of the bed to the temperature of the bed, so that under steady-state conditions the temperature of the bed is essentially constant. The recycle gas is introduced into the bottom of the reactor through inlet 18 and into the fluidized bed through distributor plate 20. Compressor 25 may also be located upstream of heat exchanger 26 in the direction of flow.

The monomers are introduced into the reactor via inlet 18. In order to avoid condensation of the _C5 to _C8 comonomers in particular in the reaction system, it is expedient to use the _C5 to _C8 comonomers in a gaseous mixture. The temperature of the gaseous mixture in the reaction system must be above the dew point of such a mixture. This is achieved in particular by maintaining the temperature of the gaseous mixture at least about 3 to 10°C above the dew point by means of a suitable adjustment of the temperature of the coolant in the heat exchanger 26. If condensation of any of the C5 to _C8 comonomers occurs on the cooled surface of the heat exchanger 26, these comonomers are readily re-vaporized by being brought into contact with part or all of the recycled gas stream which is maintained at a temperature above the dew point.

The distributor plate 20 plays an important role in the operation of the reactor. The fluidized bed contains both the growing and formed polymer particles and the catalyst particles. Since the polymer particles are hot and possibly active, their settling must be prevented, since if any quiescent mass were to form, the active catalyst contained in these particles could cause the reaction to continue and the particles would melt due to the heat generated. It is therefore important that the recycle gas is passed through the bed at a rate sufficient to fluidize the particles at the bottom of the bed. The bottom of the bed is formed by a distribution plate 20, and this plate may be a screen, a slotted plate, a perforated plate, a plate with caps (analogous to the trays of a distillation column), etc. All parts of the plate may be stationary, or the plate may be of the moving type, such as the plate described in U.S. Patent No. 3,298,792. Whatever the design of the plate, the plate must disperse the recycle gas among the particles at the bottom of the bed and maintain them in a fluidized state. The function of the plate is also to support a quiescent bed of resin particles when the reactor is not in operation. The moving parts of the plate may be used to dislodge polymer particles trapped on or in the plate.

Hydrogen can be used as a chain transfer agent. The hydrogen/ethylene molar ratio that can be used is in the range of about 0 to about 2.0 moles of hydrogen per mole of monomer in the gas stream.

Any gas that is inert to the catalysts and reactants may also be present in the gas stream. The activating compound is preferably added to the gas recycle system at its hottest point. Therefore, the activator is preferably introduced into the recycle line downstream of the heat exchanger, such as the reservoir 27 via line 27A.

As agents for regulating molecular weight, i.e. as chain transfer agents, in addition to hydrogen, compounds of the general formula can also be used together with catalysts.

Zn(R _a J(R _b ) where

R _a and R _b are the same or different C ₁ to C ₁₄ aliphatic or aromatic hydrocarbon radicals.

This achieves a further increase in the melt flow index values of the produced copolymers. About 0 to 50, and preferably about 20 to 30 moles of zinc compound (calculated as Zn) per mole of titanium compound (calculated as Ti) may be added to the gas stream fed to the reactor. The zinc compound may be fed to the reactor preferably in the form of a dilute solution (concentration 2 to 30% by weight in a hydrocarbon solvent or in absorbed form in a solid diluent such as silica of the type described above, in an amount of about 10 to 50% by weight ^. These substances tend to be pyrophoric. The zinc compound may be fed either alone or optionally together with an activating compound from a metering device, not shown in the figure, which may be located next to the metering device 27. It is fed in the vicinity of the hottest part of the gas recycle system.

The fluidized bed reactor must operate at a temperature below the sintering temperature of the polymer particles. In order to prevent sintering of the particles, it is necessary that the operating temperature is below the sintering temperature. In the production of ethylene copolymers, it is preferable to operate at a temperature of about 30 to about 115 °C, most preferably at a temperature of about 75 to 95 °C. A temperature of about 75 to 95 °C is used for the production of products with a density of about 910 to 920 kg/m ³ and temperatures of about 80 to 100 °C are used for the production of products with a density of about 920 to 940 kg/m ³ and temperatures of 90 to 115 °C are used for the preparation of products with a density above 940 kg/m ³ .

The fluidized bed reactor operates at pressures up to about 6.87 MPa, preferably at pressures from about 1.03 to 2.41 MPa, with higher pressures in this range facilitating heat transfer since the increase in pressure results in an increase in the unit volume heat capacity of the gas.

The partially or fully activated precursor is injected into the bed at a rate equal to the rate at which it is consumed through a feed 30 located above the distributor plate 20. Preferably, the catalyst is injected at a location located about 6.35 to 19.05 mm above the plate. An important feature is that the catalyst is injected at a location above the distributor plate. Since the catalysts used are highly active, injection into the space below the distributor plate could result in polymerization starting at that location, which could ultimately lead to clogging of the distributor plate. Injecting the catalyst into the fluidized bed, on the other hand, helps distribute the catalyst throughout the bed and prevents the formation of high catalyst concentration areas that could lead to hot spots.

A gas that is inert to the catalyst, such as nitrogen or argon, is used to introduce the partially or fully reduced precursor and any activating compound or transfer agent, preferably not in the gaseous state, into the bed.

The productivity of the bed is controlled by the catalyst injection rate. The productivity of the bed can be increased simply by increasing the catalyst injection rate and decreased by decreasing the catalyst injection rate.

Since any change in the catalyst injection rate results in a change in the rate of heat evolution, the recycle gas temperature is adjusted up or down to match the rate of heat evolution. This ensures that the bed temperature is maintained at a substantially constant level. Both the fluidized bed and the recycle gas cooling system must, of course, be equipped with appropriate instrumentation to detect changes in the bed temperature and the cooling system so that the operator can make appropriate adjustments to the recycle gas temperature.

Under certain combinations of operating conditions, the fluidized bed is maintained at substantially the same height by withdrawing a portion of the bed as product at the same rate as the rate of formation of polymer product particles. Since the rate of heat evolution is directly proportional to the rate of formation of the product, the rate of formation of polymer particles at a constant gas velocity is determined by measuring the rise in gas temperature along the reactor (i.e., the temperature difference between the temperature of the entering gas and the temperature of the leaving gas).

The polymer product particles are preferably withdrawn continuously through line 34 or near the distributor plate 20. The particles are withdrawn in suspension with a portion of the gas stream being vented before the particles settle to prevent further polymerization and caking before the particles reach the zone where they are collected. The suspension gas, as already mentioned, can also be used to transfer the product from one reactor to another.

The polymer product particles are conveniently and preferably withdrawn by means of two sequentially operating timed valves 38 and 38, between which a separation zone 40 is provided. When valve 38 is closed, valve 36 is opened, allowing the gas/product mixture to pass into zone 40. Valve 38 is then opened to transfer the product to an outer isolation zone. Valve 38 is then closed and remains closed until the next product withdrawal.

Finally, the fluidized bed reactor is equipped with a suitable venting system that allows the bed to be vented at start-up and at shutdown. The reactor does not require the use of any stirring devices and/or wall scraping devices.

The highly active supported catalyst system provides a product with a mean particle size of approximately 0.127 to

1.32 mm, preferably between approximately 0.50 to 1.0 mm.

The feed stream of gaseous monomer, optionally with inert gaseous diluents, is introduced into the reactor at a rate such that a space-time yield of about 32.4 to about 162.2 kg/h/m ³ of bed volume is achieved.

The term "crude resin" or polymer herein refers to the polymer in granular form as obtained from the polymerization reactor.

Films produced by the method of the invention have a thickness of from about 0.00254 mm to about 0.508 mm, preferably from about 0.00254 to about 0.254 mm, and most preferably from about 0.00254 to about 0.1016 mm. Films with a thickness of 0.00254 to 0.1524 mm have the following properties: a tear strength of greater than about 31.17 J/mm, an elongation of greater than about 400%, a heat shrinkage of less than 3% after heating to 105 to 110 degrees Celsius and cooling to room temperature, an impact tensile strength of about 41.5 to about 166.77 J/cm ³ and a tensile strength of about 13.74 to about 48.09 MPa.

Various conventional additives, such as glidants, anti-blocking agents and antioxidants, may be added to the films in accordance with common practice.

Films can be produced by extrusion blow molding. In this process, a molten ethylene-hydrocarbon copolymer with a narrow molecular weight distribution is extruded using an extruder with an extrusion screw having a length to diameter ratio in the range of 15:1 to 21:1, preferably 15:1 to 18:1, most preferably 18:1. The extrusion screw includes a metering, transition and metering zone. The extrusion screw optionally includes a mixing zone, such as the screw described in U.S. Patent Nos. 3,486,192, 3,730,492, 3,756,574. The mixing zone is preferably located at the end of the screw.

For example, an Egan 24/1 extruder can be used. This extruder typically has a screw with a length to diameter ratio of 24:1. The screw according to the invention can be directly replaced by an extrusion screw with an L/D ratio of 2.4/1. The space remaining in the extruder barrel after replacing the screw can, for example, if an 18/1 screw is used instead of an L/D ratio of 24/1, be partially filled with various inserts, torpedoes or stops to reduce the residence time of the polymer melt. The extruder barrel can also be directly adapted to an extrusion screw with a length to diameter ratio of 18/1.

Fig. 5 shows a diagram of an extruder barrel containing an extrusion screw which can be used in the method according to the invention. The extruder barrel 2 contains an extrusion screw 3 and has a mixing head 4. The extrusion screw has a length A and a diameter B. The ratio of the length of this extrusion screw to the diameter is in the range of about 15:1 to 21:1. The insert 7 in the extruder barrel contains a stop 5. The hopper 1 and the vent plate 6 are also shown.

The molten polymer may be extruded through an annular die having a gap wider than about 1.27 mm and narrower than about 3.05 mm, as described in U.S. Patent Application No. 892,324 filed March 31, 1978 and refiled February 16, 1979 under Serial No. 012,795 (W. A. Fraser et al.: Process for Making Low Density Ethylene Hydrocarbon Copolymer Film).

The polymer is extruded at a temperature of about 163 to about 260 °C in the form of a tube in a vertical direction from the bottom up. However, it can also be extruded in a top-down direction or sideways. After the molten polymer has been extruded through an annular die, the tube is blown to the desired size, cooled or allowed to cool, and the bag is compressed. The compressed bag is compressed by guiding the bag through a flattening frame and a series of pressure rollers. These rollers are driven and thereby ensure that the tubular film is pulled away from the annular die.

An overpressure of gas, for example air or nitrogen, is maintained inside the cylindrical bag. As is known from conventional film production, the gas pressure is adjusted so as to achieve the desired degree of expansion of the bag. The degree of expansion is defined as the ratio of the circumference of ^the fully inflated bag to the circumference of the annular nozzle and is in the range from 1/1 to 6/1 and preferably from 1/1 to 4/1. The hose is cooled in a conventional manner, for example by air cooling, water cooling or by means of a cooling mandrel.

The properties of the ethylene-hydrocarbon copolymer during stretching are excellent. The stretching ratio, defined as the ratio of the gap width to the product of the film thickness and the blowing ratio, is maintained in the range of about 2 to 250, and preferably about 25 to 150. Very thin films can be produced from these ethylene-hydrocarbon copolymers at a high stretching ratio even when the copolymer is highly contaminated with foreign particles and/or gel. Thin films with a thickness of more than 0.0127 mm can have an elongation in the longitudinal direction of from about 400 to about 700% and in the transverse direction of from about 500 to about 700%. In addition, the films obtained are not frangible. "Frangibility" is a qualitative property describing the tearing behavior of a notched film at a high deformation rate. Frangibility reflects the speed of crack propagation. This is a user characteristic of a certain type of foil and does not result from basic knowledge in this field.

As the ethylene-hydrocarbon copolymer exits the annular die, it begins to cool, its temperature drops below the melting point, and the polymer solidifies. The optical properties of the extruded polymer change as it crystallizes, forming a solidification line. The position of the solidification line above the annular die is a measure of the cooling rate of the copolymer film. The cooling rate has a very significant effect on the optical properties of the film.

Films can also be produced by slot extrusion with a flat die. This method of film extrusion is well known and involves extruding a sheet of molten polymer through a flat slot die and then cooling the extruded sheet, for example, with a cooled roll or water bath. When using a cooling roll, the film can be extruded horizontally and guided to the top of the cooling roll or vertically downwards and guided to the bottom of the cooling roll. The cooling rate of the extruded material in slot extrusion is very high. The cooling by the cooling roll or water bath is so rapid that as the extruded material cools below its melting point, crystallites nucleate very quickly, supramolecular structures have a short time to grow and the size of the spherulites is therefore very small. The optical properties of films produced by slot extrusion are much better than those obtained using lower cooling rates in extrusion blow molding. The temperature of the extruded material in slot extrusion is usually much higher than the typical temperatures used in blow molding. Melt strength is not a limiting factor in this extrusion process. Both shear and extension viscosity are lower. Films can usually be extruded at higher speeds than in blow molding. Higher temperatures reduce the shear stress in the die and allow for increased throughput without melt breakage.

When high-pressure polyethylene is extruded at high stretch ratios, a property of the melt of high-pressure low-density polyethylene, known as solidification under deformation, causes significant molecular orientation. Films of high-pressure low-density polyethylene produced in this manner can have very unbalanced mechanical properties. The ethylene-hydrocarbon copolymers with narrow molecular weight distributions used in the present invention, as already mentioned, exhibit less melt solidification during deformation. As in the production of blow molded tubing, these materials can be extruded at higher stretch ratios in slot extrusion without excessive molecular orientation. The balance of mechanical properties of these films does not change significantly with increasing stretch ratios.

The properties of the polymers produced in the examples were measured by the following test methods:

Density:

For substances with a density below 940 kg/ ^m3, the procedure according to ASΓΜ—1505 is used, whereby the plate is tempered for 1 hour at 100 °C to approach the state of equilibrium crystallinity. For substances with a density of 940 kg/ ^m3 and above, a modified procedure is used, whereby the test plate is tempered for 1 hour at 12 °C to approach the state of equilibrium crystallinity and then rapidly cooled to room temperature. All density data are given in kg/ ^m3 . All measurements were made in a density gradient column.

Melt flow index (MI)

ASTM D—1238—Condition E measurement is performed at 190°C, data in g per 10 minutes.

Flow rate (High Load Melt Flow Index) (HLMI)

ASTM D—1238—Condition F

The measurement is performed using 10 times the load than when measuring the melt flow index, the data is in g/10 min.

Flow rate ratio = flow rate_ melt flow index ' Productivity

A sample of the resin product is ashed and the weight % of ash is determined, since the ash is essentially composed of catalyst. The productivity is therefore given in kg of polymer produced per kg of total catalyst consumed. The amounts of titanium, magnesium and chlorine in the ash are determined by elemental analysis.

n-Hexane extractables (FDA test used for polyethylene films intended to come into contact with food). A film sample with a thickness of 0.038 mm and an area of 1290 ^cm2 is cut into strips measuring 2.54 x 15.24 cm and weighed to the nearest 0.1 mg. The strips are placed in a container and extracted with 300 ml of n-hexane at a temperature of 50 + 1 °C for 2 hours. The extract is then decanted into a balanced culture dish. After drying the extract in a vacuum desiccator, the culture dish is weighed to the nearest 0.1 mg. The amount of extract converted to the original weight of the sample is then reported as the mass fraction of n-hexane extractables.

Volumetric weight

The resin is poured into a 100 ml graduated cylinder using a funnel with a stem diameter of 9.5 mm up to the 100 ml mark. The cylinder is not shaken while filling. The difference in weight is determined by weighing.

Molecular weight distribution (Mw/ /Mn)

It is determined by gel chromatography, in the case of resins with a density below 940 kg/m ³ ^ÉLmagnesium and titanium still exist in the form of salts. Styrogel with a pore size sequence of ΙΟ ⁶ , ΙΟ ⁴ , 10 ³ , 10 ² and 6 nm is used as a stationary phase and perchloroethylene as a solvent, while working at 117 °C. In the case of resins with a density of 940 kg/m ³ or higher, Styrogel with a pore size sequence of ΙΟ ⁶ , ΙΟ ⁵ , ΙΟ ⁴ , 10 ^θ , β nm is used. o-dichlorobenzene is used as a solvent at 135 °C. Detection is carried out in all cases by infrared radiation at 3.45 μνη.

The following examples illustrate the process of the invention but do not limit its scope in any way.

Example I and

Preparation of impregnated precursor

In a 12 L flask equipped with a mechanical stirrer, 41.8 g (0.439 mol) of anhydrous magnesium chloride and 2.5 L of tetrahydrofuran (THF) are placed. To this mixture, 27.7 g (0.184 mol) of titanium tetrachloride are added dropwise over a period of 1/2 hour. It may be necessary to heat the mixture to 60 °C for about 1/2 hour to completely dissolve the substance.

To the above solution, 500 grams of porous silica are added and the mixture is stirred for 1/4 hour. The mixture is dried by blowing nitrogen at 60°C for about 3 to 5 hours. A dry, free-flowing powder is obtained, which has the same particle size as the silica particles used. The absorbed precursor has a composition corresponding to the formula

TiM _g3 , ₀ Cl ₁₀ (THF) ₆ , ₇ Example I b

Preparation of impregnated precursor from pre-prepared precursor.

In a 12 L flask equipped with a mechanical stirrer, 146 g of the pre-sodium salt are dissolved in 2.5 L of dry THF. The solution may be heated to 60°C to facilitate dissolution. 500 g of porous silica are added and the mixture is stirred for 1/4 hour. The mixture is then dried by purging with a stream of nitrogen at a temperature not exceeding 60°C for about 3 to 5 hours. A dry, free-flowing powder is obtained which has a particle size similar to that of the silica particles.

The precursor used in process Ib is obtained in the same manner as in example Ia except that it is isolated from a solution in tetrahydrofuran by crystallization or precipitation. The precursor may be analyzed at this stage for magnesium and titanium content, since some amount of magnesium or/and titanium compound may have been lost during the isolation of the precursor. The empirical formulas used herein to characterize the composition of the precursor are derived assuming that the amounts originally added to the electron donor compound are the same. The amount of electron donor compound is determined chromatographically.

Example II

Activation procedure

The required weight of the impregnated precursor and the activating compound is placed in a mixing drum along with a sufficient amount of an anhydrous aliphatic hydrocarbon diluent, such as isopentane, to form a suspension.

The activating compound and the precursor are used in such amounts as to obtain a partially activated precursor having an Al/Ti ratio greater than 0 but less than or equal to 10. The preferred ratio is 4 to 8:1.

The suspension is stirred thoroughly at room temperature and atmospheric pressure for 1/4 to 1/2 hour. The resulting suspension is dried by a stream of dry inert gas, such as nitrogen or argon, at atmospheric pressure and at a temperature of 65 ± 10 °C, to remove the hydrocarbon diluent. This usually takes about 3 to 5 hours. The resulting catalyst is in the form of a partially activated precursor impregnated into the pores of the silica. The substance is in the form of a freely moving powdery material. The particles correspond in size and shape to the particles of silica. The product is not pyrophoric if the alkylaluminum content is not more than 10% by weight. It is kept under dry inert gas, such as nitrogen or argon, until it is used. It is ready for introduction into the polymer reactor, where complete activation is then carried out.

When additional activating compound is introduced into the polymerization reactor to complete the activation of the precursor, it is fed to the reactor in the form of a dilute solution in a hydrocarbon solvent, such as isopentane. These dilute solutions contain about 5 to about 30% by volume of the activating compound.

The activating compound is introduced into the polymerization reactor in such an amount that the Al/Ti ratio in the reactor is maintained at a value of from 10 to 400:1, preferably from 15 to 60:1.

Example III

Preparation of the precursor

In a 51-liter flask equipped with a mechanical stirrer, 16.0 g (0.163 mol) of anhydrous magnesium chloride are mixed under nitrogen with 850 ml of pure tetrahydrofuran. The mixture is stirred at room temperature (about 25 °C) and 13.05 g (0.069 mol) of titanium tetrachloride are added dropwise. After the addition is complete, the contents of the flask are refluxed for 1/2 to 1 hour to dissolve the solids. The system is cooled to room temperature and then 3 liters of pure n-hexane are slowly added over 1/4 hour. A yellow solid precipitates. The supernatant is decanted and the solid is washed three times with one liter of n-hexane. The solids are then filtered off and dried in a rotary evaporator at 40 to 60 °C. 55 g of solid precursor are obtained.

The precursor can be analyzed for magnesium and titanium at this point, since some amounts of magnesium and/or titanium compound may have been lost during the isolation of the precursor. The empirical formulas used herein to characterize the precipitation of the precursor are derived assuming that the magnesium and titanium still exist in the form of compounds that were originally added to the electron donor compound, and that the remaining mass of the precursor is made up of the electron donor compound.

Analysis of the solid reveals the composition of Mg:

6.1%, Ti: 4.9%, which corresponds to the formula: o..,!! i ikj.,,.

where

THE represents tetrahydrofuran.

Example 1

Activation procedures

Procedure A

This activation process involves a multi-stage activation of the precursor. In this process, the activation is carried out in such a way that the precursor is only partially reduced before being introduced into the polymerization reactor and the reduction is completed in the reactor.

The desired weight of dry inert support is placed in a mixing vessel or drum. In the examples described herein, about 500 g of silica or about 1000 g of polyethylene support are used as the support. The inert support is then mixed with a sufficient amount of anhydrous aliphatic hydrocarbon diluent, such as isopentane, to form a suspension. Typically, about 4 to 7 ml of diluent is used per 1 g of inert support. The desired weight of precursor is then placed in the mixing vessel and thoroughly mixed with the suspension. In the cases described herein, about 80 to 135 g of precursor are used to produce catalysts containing 1 ± 0.1 mmol of titanium per gram of precursor.

The required amount of activator required for partial activation of the precursor is added to the contents of the mixing vessel. Such an amount of activator is used so that the Al/Ti ratio in the partially reduced precursor is up to 10:1, and preferably 4 to 8:1. The activator is added to the mixing drum in the form of a solution containing about 20% by weight of the activator (in these examples triethylaluminum) in an inert aliphatic, hydrocarbon solvent (in these examples hexane). Activation is carried out by thorough mixing and bringing the activator into contact with the precursor. All of the above-described operations are carried out at room temperature and at atmospheric pressure under an inert atmosphere.

The resulting suspension is dried by a stream of dry inert gas, such as nitrogen or argon, at atmospheric pressure and at a temperature not exceeding 60°C to remove the hydrocarbon diluent. This usually takes about 3 to 5 hours. The substance is in the form of a freely moving powdery material, consisting of a uniform mixture of the activated precursor with an inert carrier. The dried, non-pyrophoric product is stored under inert gas.

When additional activator is introduced into the polymerization reactor according to Process A to complete the activation of the precursor, the activator may first be absorbed in an inert carrier such as silica gel or polyethylene, but most preferably the activator is introduced into the reactor as a dilute solution in a hydrocarbon solvent such as isopentane.

When the activator is to be absorbed in a silica gel carrier, the two substances are mixed together in a vessel containing about 4 ml of isopentane per gram of carrier. The resulting suspension is then dried for about 3 to 5 hours under a stream of nitrogen at atmospheric pressure at a temperature of 65 ± 10°C to remove the hydrocarbon diluent.

When the activator is injected into the polymerization system in the form of a dilute solution, a concentration of about 5 to 10% by weight is preferably used.

Regardless of the method of introducing the activator into the polymerization reactor to complete the activation of the precursor, the activator is used in such an amount that the Al/Ti ratio in the polymerization reactor is maintained at a value of from 10 to 400:1 and preferably from 10 to 100:1.

Before use, the silica is dried for at least 4 hours at at least 600 °C. Procedure B

In this process, complete activation of the precursor is achieved by bringing the precursor into contact with the activator absorbed in an inert carrier during mixing.

The activator is absorbed in an inert carrier by suspending it together with the carrier in an inert hydrocarbon solvent and then drying the suspension to remove the solvent. The mixture obtained contains about 10 to 50% by weight of the activator. The procedure is as follows: 500 g of silica (previously dried for 4 hours at 800 degrees Celsius) are placed in a mixing vessel. The required amount of activator is then added to the mixing vessel in the form of a 20% (by weight) solution in a hydrocarbon solvent such as hexane and the mixture is mixed, suspended with the inert carrier at room temperature and atmospheric pressure. The solvent is then evaporated by drying the resulting suspension at 65 ± 10 °C for about 3 to 5 hours at atmospheric pressure with a stream of dry inert gas such as nitrogen. The dried mixture is in the form of loose particles that have the same size as the particles of the carrier used.

About 500 g of dried silica, on which the activator is applied (in a weight ratio of 50/50), is placed in a mixing vessel and the required amount of precursor (80 to 100 g) is added. The substances are mixed thoroughly for 1 to 3 hours at room temperature and atmospheric pressure under a dry inert gas, such as nitrogen or argon. The resulting mixture is in the form of a physical mixture of loose particles, which have a size of the order of 10 to 150 μΐη. During mixing, the activator on the support comes into contact with the precursor and activates it completely. During the exothermic reaction that occurs, the temperature of the catalyst should not exceed 50 °C to prevent any deactivation of the catalyst. The resulting activated catalyst has an Al/Ti ratio of about 10 to 50 and, if it contains more than 10% by weight of activator, may be pyrrhotic. Before being introduced into the reactor, it is stored under a dry inert gas, such as nitrogen or argon.

Example 1

Preparation of copolymers

Ethylene is copolymerized with propylene or 1-butene (in Examples 1 to 2 with propylene and in Examples 3 to 14 with 1-butene). In all examples, the catalyst prepared as described above and activated by Activation Procedure A is used. Polymers are produced which have a density of 940 kg/m ³ or less. In all cases, the partially activated precursor has an Al/Ti molar ratio of 4.4 to 5.8. The activation of the precursor in the polymerization reactor is completed with triethylaluminum, which is used either in the form of a 5% by weight isopentane solution (Examples 1 to 3 and 4 to 14) or in absorbed form on silica, 50/50 by weight (Examples 4 and 5) in such an amount as to obtain a fully activated catalyst in the reactor with an Al/Ti molar ratio of about 29 to 140.

Each of the reactions is continuously carried out in a fluidized bed reaction system for at least 1 hour after reaching equilibrium at a pressure of 2.06 MPa at a gas velocity corresponding to about five to six times the Gmf value and a space-time yield of 48.7 to 97.3 kg/m ² /h. The reaction system described above is used as shown in the figure. Its lower part is 3.05 m high and its inner diameter is 0.343 m. The upper part is 4.88 m high and its inner diameter is 0.597 m.

In several examples, diethylzinc (as a 2.6% by weight solution in isopentane) was added during the reaction to maintain a constant Zn/Ti molar ratio. In the case of diethylzinc, triethylaluminum was also added as a 2.6% by weight solution in isopentane.

The following Table A lists the various operating conditions used in Examples 1 to 14, i.e. the precursor content in the precursor-silica mixture (% by weight); the Al/Ti ratio in the partially activated precursor, the Al/Ti ratio maintained in the reactor; the polymerization temperature; the ethylene content in the reactor (% by volume); the hydrogen/ethylene molar ratio; the comonomer (C _x ) to ethylene (C ₂ ) molar ratio in the reactor, the catalyst productivity and the Zn/Ti molar ratio.

Table B shows the properties of the granular raw resins produced in Examples 1 to 14, i.e. density, melt flow index (MI), melt flow rate ratio (MFR), bulk density, mean particle size, ash content in wt% and titanium content (ppm).

Table A

Reaction conditions in examples 1 to 14

Example Precursor content [% by weight] Molar ratio of Al/Ti in partially activated precursor Molar ratio of Al/Ti in the reactor Temperature PC] Contents Molar mass Molar mass ethylene rate H ₂ / [%, vol.] /ethylene como·ncmer/ /ethylene 1 8.3 5.8 40.5 90 41.7 0.492 0.486 2 8.3 5.8 50.8 90 39.7 0.566 0.534 3 20.1 4.50 88.3 85 56.3 0.148 0.450 4 19.8 4.40 26.7 85 54.1 0.350 0.350 5 19.8 4.40 26.7 80 54.1 0.157 0.407 6 6.9 5.08 42.0 85 49.2 0.209· 0.480 7 6.9 5.08 33.6 85 46.5 0.208 0.482 8 6.9 5.08 28.8 85 42.1 0.206 0.515 10 8.3 5.8 124.6 90 45.1 0.456 0.390 11 8.3 5.8 80.8 90 42.7 0.365 0.396 12 8.3 5.8 52.0 90 48.4 0.350 0.397 13 8.3 5.8 140.1 90 42.6 0.518 0.393 14 8.3 5.8 63.5 90 40.8 0.556 0.391 Table l B Features polymers produced in examples 1 to 14 Example Density Flow Index Speed ratio Volumetric Medium size [kg/m ^:! ] melts melt flow weight particles [mm] [g/10 min) [kg/m ^:! ] 1 927 22.0 24.4 272.5 0.58 2 929 24.0 23.4 283.8 0.58 3 925 0.61 27.1 272.5 0.76 4 931 12.0 26.7 272.5 0.70 5 923 1.47 28.2 253.0 1.03 6 919 3.41 25.9 272.5 1.40 7 925 2.90 24.5 283.8 1.50 8 919 3.10 24.6 262.8 1.45 10 929 16.0 24.1 280.6 0.58 11 929 15.3 24.0 269.2 0.59 12 928 11.5 24.1 270.9 0.63 13 929 20.7 24.3 280.6 0.65 14 929 29.2 26.1 272.5 0.52

Appendix 1 and 2

The copolymer of ethylene and 1-butene prepared according to Example 1, which has a density of 920 kg/ ^m3 and a melt flow index of 2 g/10 min, is processed into a film with a thickness of 0.038 mm by blowing an extruded tube using a 24/1 extruder with an extrusion screw with a diameter of 63.5 mm and an L/D ratio of 18:1. The extrusion screw has a metering zone with a length of 317.5 mm, a transition zone ¹ with a length of

190.5 mm, a metering zone with a length of 508 mm and a mixing zone with a length of 127 mm. The mixing part is grooved and has the following characteristics: diameter 63.5 mm, channels 76.2 mm, channel radius 13.74 mm, width of the mixing field of the thread 6.35 mm, width of the wiping field of the thread 5.08 mm and length of the mixing cap 114.3 mm. The empty space in the cylinder is filled with an insert with a diameter of 63.25 mm and a length of 279.4 mm, which contains a stop with a length of 228.6 mm and a diameter of 25.4 mm.

The temperature in the extruder barrel is about 177 °C. A 20/60/20 mesh screen and an extrusion die for the production of foils by blowing with a diameter of 152.4 mm with a spiral of 1 ka I hunters needle are used. The die temperature is set at about 177 °C. The production rate is 21.79 kg/h. The height of the pressure roll is about 4.57 au. Cooling is carried out by means of a Venturi-type air ring. All foils are produced at a blowing ratio of 2 : 1 [ratio of the circumference of the bag to the circumference of the die]. The speed of the extruder screw [min ^-1 ], the output in kg/h per 1 revolution, the shaft power in kW, the efficiency in kWh/kg and the temperature of the extrudate are given in Table I.

Example 3

The procedure of Example 2 is repeated in all details, except that the production rate is 31.78 kg/h. The results are given in Table I.

Example 4

The procedure of Example 2 is repeated in all details, except that the production rate is 43.58 kg/h. The results are given in Table I.

Example 2

Production rate [kg/hj 21.79

Speed [min ^-1 ] 30

Power per revolution [kg/h/rev.h 0.73

Shaft power [kW] 2.39

Efficiency [kWh/kg] 0.110

Temperature [°C] 196.3

Example 5

The procedure according to example 2 is repeated exactly, with the difference that the production speed is

31.78 kg/h and an extruder screw with a L/D ratio of 24:1 is used instead of the extruder. Table II

Example 3

L/D ratio of the extruder screw Production rate [kg/h]

Speed [min ^-1 ]

Power output [kg/h/rev.h Shaft power [kW]

Efficiency [kWh/kg]

Temperature [°C] : 1 31.78 44 0.73 4.39 0.136 221.3

31.78 43.58

60

0.73 0.73

4.25 6.56

0.133 0.151

210.2 220.2 whose screw has an L/D ratio of 18:1. The results using this production rate and extruder screw are compared with Example 3, in which the production rate is 31.78 kg/h and the extruder screw used has an L/D ratio of 18:1. The results are shown in Table II.

: 1

31.78 41

0.75

4.71

0.151

235.2

The data in Table II show the advantages achieved by using an extruder screw with an L/D ratio of 18:1, i.e. lower energy requirements at the same production rate, better energy utilization (as shown by lower kWh/kgj values) and lower extrudate temperature.

Rapid melting in extrusion machines

The shear melting mechanism is accepted as the melting mechanism for all crystalline and glassy thermoplastic polymers.

Shear melting can be described as a mechanism for the development of heat in a thin film that occurs between the unmelted solid polymer and the extruder barrel. The film that forms in the extruder channel is typically 0.381 to 0.508 mm thick and shear rates of up to 5000 s~ ^l can occur in the thin film of melt. As the channel progresses to the extruder outlet, the amount of solid polymer in the channel decreases due to heating and melting in the thin film that is under shear stress. For high-density polyethylene, this melting continues along 50 to 80% of the length of the extruder. The polymers used in the invention were not observed to exhibit this slow melting mechanism seen in high-density polyethylene in the series of tests described below.

The melting mechanism of the polymers used in the invention is faster, as evidenced by the fact that it occurs along less than 40% of the length of the screw (which has a length to diameter ratio of 15:1 to 21:1). In some cases, melting occurs only over a length corresponding to 2 to 5 screw diameters from the point where melting begins.

The differences between fast melt and slow melt high pressure polyethylene resin are threefold:

1) Melting is much faster

2) Melting is not limited to the thin film between the solid polymer and the cylinder

3) The melt tends to form on the trailing edge instead of the leading edge of the extrusion screw screw.

The melting mechanism described above, referred to as so-called rapid melting, applies as has been demonstrated for feedstock in the form of pellets and granules.

Fast melting is responsible for the ability of a short screw to provide a quality melt at a short length and at a high film extrusion speed. In film extrusion, quality, i.e. uniform temperature distribution, is important. To achieve this uniform temperature distribution, tempering of the melt is necessary. In the case of a short screw, melting occurs quickly and tempering occurs over the remaining length of the screw.

Method for determining melting rate

A 2.5 inch (63.5 mm) extruder with a rapid cool or heat-up barrel is used, i.e., with the ability to heat to 150°C from room temperature (about 25°C) at a rate of about 15°C/min. Cooling is accomplished by means of a cooling channel located in the extruder barrel, which can be programmed to continuously pass cold water through the cooling channel. For rapid heating, high-intensity heaters are used in combination with steam that is passed through the cooling channel.

At the rear of the extruder is a large air cylinder that is used to apply a large force to the extruder screw (about 549.6 kPa per 304.8 mm piston, which corresponds to a total force of about 39,240 N.

The melting rate is determined by freezing the molten polymer in the cylinder of an extruder and then pushing the “frozen” sample out.

The frozen sample is obtained as follows:

The extruder loaded with polymer is run at the desired conditions (i.e. 60 rpm, barrel temperature 149°C), then the cooling water is suddenly turned on and the speed is reduced to zero. Cooling of the barrel is continued until the barrel is at room temperature. Wait until the polymer is at room temperature, which usually takes at least 4 hours, and then the electric heater is turned on and steam is introduced into the barrel to quickly loosen the liner. Then, during the heating, a large force is applied to the screw. After an interval of 3 to 10 minutes, the screw is driven out of the extruder by applying force to the rear of the screw. The heating melts as little polymer as possible at the barrel wall. After the screw is removed, the solid polymer is removed from the screw and divided into portions depending on the ratio of screw length to diameter to determine the relative amount of unmelted polymer compared to the molten polymer under the extrusion conditions existing at the time of cooling.

A fast-melting polymer is a polymer that, when passed through the extrusion machine described above, melts completely within a screw length range corresponding to 2 to 5 screw diameters from the start of melting. Slow weakening under shear stress

The term slow attenuation means that when the shear rate of the polymer under test is measured in a capillary plastometer at 204.7°C over a shear rate range of 1 to 5000 s- ¹ , the viscosity of the polymer under test, measured in Pa.s (lb.s/ ^in.2 ), decreases consistently over the test range, but is consistently higher throughout the test range than the viscosity of a high-pressure ethylene homopolymer having the same density (+5 kg/ ^m3 ) and melt flow index (+0.3 g/10 min.).

A comparison of the shear strain rates of two such groups of polymers is shown in Figures 6 and 7. Figure 6 shows a comparison between the shear strain rates of a low-pressure copolymer of ethylene and 1-butene used in the process of the invention (low-pressure polyethylene) having a melt flow index of 1 dg/min and a high-pressure homopolymer of ethylene (high-pressure polyethylene) having a melt flow index of 0.7 dg/min. Both polymers tested have a density of 920 kg/ ^m3 . The curves for the individual resins will be essentially the same regardless of the density of the polymers (ranging from about 910 to 960 kg/ ^m3 ). However, the position of the curves on the graph will change as the melt flow index of the polymer tested changes. As the melt flow index increases, the curves are positioned higher on the graph.

Fig. 7 shows a comparison between the shear strain rate of a low-pressure copolymer of ethylene and 1-butene used in the process according to the invention (low-pressure polyethylene), which has a melt index of 2.0 dg/min, with a high-pressure homopolymer of ethylene (high-pressure polyethylene), which also has a melt index of 2.0 dg/min. Both polymers tested have a density of 920 kg/m ³ .

Although the extrusion process described in these references has been explained in the case of polymers produced according to the earlier US patent applications mentioned above, the use of these polymers is not limited to them. Other ethylene polymers (homopolymers and copolymers with one or more C ₃ to C ₈ comonomers) which melt rapidly and which exhibit slow shear weakening, regardless of the molecular weight distribution of these polymers, can also be used in the extrusion process of the invention. These polymers can therefore have a relatively wide molecular weight range of about 2 to 10 and preferably about 2 to 4.

The process of the invention is generally applicable to resins having a melt flow index in the range of about 0.1 to 10 dg/min, preferably in the range of about 0.1 to 5 dg/min. These resins should also have a bulk density of about 910 to 940 kg/m ³ .

Claims (1)

A method for producing an extrudate from a fast melting ethylene polymer that exhibits a slow shear attenuation by melt extrusion in an extrusion BACKGROUND OF THE INVENTION A machine comprising an extruder screw, characterized in that an extruder screw having a length to diameter ratio in the range of 15: 1 to 21: 1 is used.