Process for producing Phillips catalysts

The invention relates to a process for producing chromium catalysts, which process comprises the steps of:

a) applying one or more chromium compound(s) to a finely divided inorganic support to form a catalyst precursor,

b) thermally treating the catalyst precursor, the step of thermally treating the catalyst precursor being carried out for at least part of the time in an oxidizing atmosphere and in such a manner that a maximum temperature of from 350 ⁰ C to 1050 ⁰ C is not exceeded.

Thus, on the one side, the thermal treatment step is carried out for a predetermined time and, for at least part of this time, in an oxidizing atmosphere and, on the other side, the thermal treatment step is carried out at a predetermined temperature which is lower than or equal to a maximum temperature of 350°C to 105O ⁰ C.

Chromium catalysts have been used for decades for preparing ethylene polymers. They are generally produced by doping inorganic supports such as silica gels or aluminum oxides with chromium (catalyst precursors). To obtain a polymerization-active catalyst, this chromium-doped catalyst precursor has to be thermally treated at temperatures of from 300 to 1050 ⁰ C in an atmosphere which is oxidizing for at least part of the time.

The thermal treatment is carried out industrially in special activators, e.g. fluidized-bed activators. The activator is charged with the catalyst precursor and brought to the activation temperature over a heating-up period which is normally a number of hours. After the activation temperature has been reached, the catalyst is maintained at this temperature for a hold period, which is usually up to 12 hours. The catalyst is finally cooled and taken from the activator.

Since this treatment, frequently also referred to as activation or calcination, takes a given time, many efforts have been made to shorten this as much as possible in order to achieve as high as possible a plant capacity. However, these efforts are subject to limits, since the thermal activation has to ensure firstly sufficient oxidation of the chromium to chromium Vl and secondly a sufficient reduction in the OH groups of the support. US 2005/0255987 A1 has proposed, for example, providing two or more hold times in order to reduce the time for the thermal treatment.

The activation temperature(s) also determine(s) the polymerization properties of the chromium catalysts to a considerable extent. Thus, the activity of the catalysts generally increases with an increase in the activation temperature, while the molar mass of the polymerization product

decreases or the HLMI potential increases. To produce a product having a prescribed HLMI, it is therefore necessary to keep the activation temperature within narrow limits, since only limited alteration possibilities are provided by the polymerization conditions, in particular the temperature and the concentration of the reactants. A general presentation on the influence of the activation conditions on the catalyst properties may be found, for example, in Advances in Catalysis, Vol. 33, 48 - 98.

It is therefore an object of the present invention that of providing a process of the type mentioned at the outset by means of which the properties of the polymer product can be changed without changing or modifying the catalyst precursor.

It has now surprisingly been found that the catalyst properties can be varied in a targeted manner by means of a targeted increase in the hold time.

The present invention accordingly provides a process of the type mentioned at the outset in which the duration of the thermal treatment step at a temperature of above 300°C is at least 1500 minutes. The duration of the thermal treatment step above a temperature of 300 ⁰ C will hereinafter also be referred to as the activation time. During the activation, a reduction in the number of OH groups on the surface of the support and oxidation of the chromium to chromium Vl take place.

The present invention further provides a process for the polymerization of ethylene or of ethylene with further 1 -olefins at temperatures of from 30 to 150 ⁰ C and pressures of from 0.1 to 50 MPa, in which the catalyst used in the polymerization has been produced by the process of the invention.

As a result of the increase in the activation time, it is possible to produce catalysts which have the same HLMI potential as conventional products but differ in an advantageous manner from these in terms of other properties. Thus, it is possible to open up a significantly wider product range using the same catalyst precursor than is possible purely by variation of the polymerization conditions. In particular, it has been found that polymers having the same HLMI potential but also other advantageous product properties can be produced using the same catalyst precursor, while the use of an additional catalyst precursor was hitherto necessary. Furthermore, the activity of the catalysts can be increased by the process of the invention.

The HLMI potential of a catalyst is a measure of the HLMI which can be achieved under identical conditions. It is thus a relative parameter by means of which catalysts can be compared with one another in respect of the HLMI. The HLMI or MFR ₂ i is the melt flow rate determined in accordance with ISO 1133 at 190 ⁰ C under a load of 21.6 kg of the polymer prepared using the chromium catalyst. The use of the HLMI potential for defining the polymerization properties of catalysts is generally customary among those skilled in the art.

To produce the chromium catalysts, the supports are doped with chromium and, if appropriate, with further elements to form the catalyst precursor and are subsequently subjected to a thermal treatment step.

In step a), the chromium compound(s) is/are applied to the finely divided inorganic support to form the catalyst precursor, which will hereinafter also be referred to as doping. Doping can be carried out by all known methods, doping from homogeneous solutions being preferred.

Preferably, chromium compounds having a valence of less than 6, particularly preferably Cr(III) compounds, are used. Such compounds are, for example, chromium hydroxide and soluble salts of trivalent chromium with an organic or inorganic acid, e.g. acetates, oxalates, sulfates or nitrates. It is particularly preferred to use salts of acids which are converted essentially into chromium(VI) without leaving a residue during activation, e.g. chromium(lll) nitrate nonahydrate.

Furthermore, chelate compounds of chromium, e.g. chromium derivatives of β-diketones, β-ketoaldehydes or β-dialdehydes, and/or complexes of chromium, e.g. chromium(lll) acetylacetonate or chromium hexacarbonyl, or organometallic compounds of chromium, e.g. bis(cyclopentadienyl)chromium(ll), organic chromic(VI) esters or bis(arene)chromium(0), can likewise be used.

The chromium compound is preferably applied from a 0.05-15% by weight solution of a chromium compound which is converted under the conditions of the activation into chromium(VI) oxide in an organic polar solvent, preferably an organic protic solvent, particularly preferably a Ci-C ₄ -alcohol, the respective solvent preferably comprising not more than 5% by weight of water. Furthermore, loading of the support without solvents, for example by mechanical mixing, is also possible.

The chromium compound is present in a concentration of usually from 0.05 to 20% by weight, preferably from 0.1 to 15% by weight and particularly preferably from 0.5 to 10% by weight, based on the solvent. The amount of solution used during doping is preferably less than the pore volume of the support.

The chromium content of the finished catalyst is usually in the range from 0.1 to 5% by weight, preferably from 0.5 to 4% by weight, particularly preferably from 1 to 3% by weight, based on the support.

The chromium catalysts to be produced may comprise not only the element chromium but also further elements such as Mg, Ca, Sr, Ba, B, Al ₁ C, Si, P, Bi, Sc, Ti, V, Mn, Fe, Co, Ni, Cu, Zn, Zr, Nb, Mo, Ru, Rh, Pd, Hf, Ta and W ₁ either alone or in combination, and also, if appropriate, one or more activators. The mentioned elements may be constituents of the support matrix or be applied to the support before, during or after step a). The use of chromium catalysts in which no further

element in addition to chromium is applied to the support is preferred. Preference is also given to applying the elements Al ₁ Ti, Zr, or Zn in addition to chromium. The amount of further elements applied is preferably from 0.1 to 7% by weight, particularly preferably from 1 to 5% by weight, based on the support.

In case of co-support, preference is given to applying the chromium and further elements together from a homogeneous solution to the support.

Possible finely divided supports are all customary particulate inorganic supports, which are usually porous. Preference is given to oxidic support materials which may still contain hydroxy groups. The inorganic metal oxide can be spherical or granular. Examples of such solids which are also known to those skilled in the art are aluminum oxide, silicon dioxide (silica gel), titanium dioxide or their mixed oxides or cogels, or aluminum phosphate. Further suitable support materials can be obtained by modification of the pore surface, e.g. by means of compounds of the elements boron, aluminum, silicon or phosphorus. Preference is given to using a silica gel whose surface can, if desired, also be modified by means of aluminum phosphate. Preference is given to spherical or granular silica gels, with the former also being able to be spray dried.

Finely divided silica xerogels are preferred as support materials; these can be produced, for example, as described in DE-A 25 40 279. The finely divided silica xerogels are preferably produced by:

a) introducing a particulate silica hydrogel which comprises from 10 to 25% by weight solids (calculated as SiO ₂ ) and is largely spherical and has a particle diameter of from 1 to 8 mm and is obtained by:

a1) introducing a sodium or potassium water glass solution into a stream having a rotational motion of an aqueous mineral acid, both longitudinally and tangentially to the stream,

a2) spraying the resulting silica hydrosol as droplets into a gaseous medium,

a3) solidifying the sprayed hydrosol in the gaseous medium,

a4) freeing the resulting largely spherical particles of hydrogel from salts without prior ageing by washing,

b) extracting at least 60% of the water comprised in the hydrogel by means of an organic liquid,

c) drying the resulting gel until no weight loss occurs at 180 ⁰ C and a reduced pressure of 13 mbar over a period of 30 minutes (xerogel formation), and

d) adjusting the particle diameter of the xerogel obtained to from 20 to 2000 μm.

In the first step a) in which the support material is produced, it is important to use a silica hydrogel which has a relatively high solids content of from 10 to 25% by weight (calculated as SiO ₂ ), preferably from 12 to 20% by weight, particularly preferably from 14 to 20% by weight, and is largely spherical. This silica hydrogel has been produced in a special way, which is described in steps a1) to a4). The steps a1) to a3) are described in more detail in DE-A 21 03 243. Step a4), in which the hydrogel is washed, can be carried out in any desired way, for example according to the countercurrent principle using water which has a temperature of up to 80°C and is slightly ammoniacal (pH up to about 10).

The extraction of the water from the hydrogel (step b)) is preferably carried out using an organic liquid, which is particularly preferably miscible with water, from the group consisting of C ₁ -C ₄ - alcohols and/or that consisting of C ₃ -C ₅ -ketones. Particularly preferred alcohols are tert-butanol, i-propanol, ethanol and methanol. Among the group of ketones, acetone is preferred. The organic liquid can also consist of mixtures of the above-mentioned organic liquids, and in any case the organic liquid comprises less than 5% by weight, preferably less than 3% by weight, of water prior to the extraction. The extraction can be carried out in customary extraction apparatus, e.g. column extractors.

Drying (step c)) is preferably carried out at temperatures of from 30 to 140°C, particularly preferably from 80 to 110°C, and at pressures of preferably from 1.3 mbar to atmospheric pressure. Due to the vapor pressure, an increasing temperature should be accompanied by an increasing pressure and vice versa.

The adjustment of the particle diameter of the xerogel obtained (step d)) can be carried out in any desired way, e.g. by milling and sieving.

A further preferred support material is produced, inter alia, by spray drying of milled, appropriately sieved hydrogels which are for this purpose mixed with water or an aliphatic alcohol. The primary particles are porous, granular particles of the appropriately milled and sieved hydrogel which have a mean particle diameter of from 1 to 20 μm, preferably from 1 to 5 μm. Preference is given to using milled and sieved SiO ₂ hydrogels.

The desired mean particle size of the supports can be varied within wide ranges and can be matched appropriately to the use of the supports. The mean particle size of the supports can thus be set, for example, to values appropriate to various polymerization processes.

In general, the mean particle diameter of the support particles is in the range from 1 to 1000 μm, preferably in the range from 10 to 500 μm and particularly preferably in the range from 30 to 150 μm. The support particles, which can preferably be produced by means of spray drying, particularly preferably have a mean particle size in the range from 30 μm to 90 μm, more preferably in the range from 40 μm to 70 μm, even more preferably in the range from 40 μm to 50 μm and very particularly preferably in the range from 40 μm to 55 μm.

Particularly preferably from 70% by volume to 90% by volume of the support particles, preferably 80% by volume of the particles, based on the total volume of the particles, have a mean particle size in the range from ≥40 μm to ≤90 μm.

Support particles which are preferably usable for polymerization in slurry polymerization processes can preferably have mean particle sizes up to 350 μm. They preferably have a mean particle size in the range from 30 μm to 150 μm. Support particles which are preferably usable for polymerization in gas-phase fluidized-bed processes preferably have a mean particle size in the range from 30 μm to 120 μm. Support particles which are preferably usable for polymerization in suspension processes preferably have a mean particle size in the range from 30 μm to 300 μm, while support particles which are preferably usable for polymerization in loop processes preferably have a mean particle size in the range from 30 μm to 150 μm. Support particles which can be used, for example, for polymerization in fixed-bed reactors preferably have mean particle sizes of greater than or equal to 100 μm, preferably greater than or equal to 300 μm, more preferably in the range from 1 mm to 10 mm, particularly preferably in the range from 2 mm to 8 mm and more preferably in the range from 2.5 mm to 5.5 mm.

Preferably from 10% by volume to 90% by volume of the support particles, based on the total volume of the particles, have a particle size in the range from > 40 μm to < 120 μm; preferably from 30% by volume to 80% by volume of the particles, based on the total volume of the particles, have a particle size in the range from 30 μm to 70 μm. Particle sizes of the support particles in the range from 30 μm to 70 μm are preferred.

The support particles preferably have a particle size distribution in which at least 90% by volume of the particles, based on the total volume of the particles, comprises particles having a size in the range from 16 μm to 500 μm, at least 75% by volume of the particles comprises particles having a size in the range from 32 μm to 200 μm, and at least 30% by volume of the particles comprises particles having a size in the range from 48 μm to 150 μm.

After drying, the support particles advantageously have a low fines content. For the purposes of the present invention, the fines content of the support particles is the proportion of support particles which have a particle size of less than 25 μm, preferably less than 22 μm, particularly preferably less than 20.2 μm. After drying, it is advantageous for less than 5% by volume of the particles, based on the total volume of the particles, to have a particle size in the range from 0 μm to 25 μm, preferably in the range from 0 μm to 22 μm, particularly preferably in the range from 0 μm to 20.2 μm. Preferably less than 3% by volume, particularly preferably less than 2% by volume, of the particles, based on the total volume of the particles, have a particle size in the range from 0 μm to 25 μm, preferably in the range from 0 μm to 22 μm, particularly preferably in the range from 0 μm to 20.2 μm. Preferably less than 5% by volume, preferably less than 2% by volume, of the particles, based on the total volume of the particles, have a particle size in the range from 0 μm to 10 μm.

Furthermore, preferably less than 30% by volume, preferably less than 20% by volume, particularly preferably less than 15% by volume, very particularly preferably less than 10% by volume, of the particles, based on the total volume of the particles, have a particle size in the range from 0 μm to 35 μm, preferably in the range from 0 μm to 32 μm.

The mean average pore volume of the support material used is in the range from 0.1 to 10 ml/g, in particular from 0.8 to 4.0 ml/g and particularly preferably from 1 to 2.5 ml/g.

The support particles produced have a pore diameter which is preferably in the range below 20 nm; the support particles preferably have a pore diameter in the range below 15 nm, and the pore diameter is particularly preferably in the range from 5 nm to 13 nm. The support particles preferably have a mean pore diameter of from 8 to 25 nm, preferably from 9 to 21 nm and particularly preferably from 9.5 to 20 nm.

The surface area of the inorganic support can be varied within a wide range by means of the drying process, in particular by the process of spray drying. Preference is given to producing particles of the inorganic support, in particular a product from a spray dryer, which have a surface area in the range from 100 m ² /g to 1000 m ² /g, preferably in the range from 150 m ² /g to 700 rτι ² /g and particularly preferably in the range from 200 m ² /g to 500 m ² /g. Supports which can be used for polymerization preferably have a surface area in the range from 200 m ² /g to 500 m ² /g. The specific surface area of the support particles is the surface area of the support particles determined by means of nitrogen adsorption using the BET technique.

The specific surface area and the mean pore volume are determined by nitrogen adsorption using the BET method, as described, for example, in S. Brunauer, P. Emmett and E. Teller in Journal of the American Chemical Society, 60 (1938), pages 309-319.

The bulk density of the inorganic supports for catalysts is preferably in the range from 250 g/l to 1200 g/l, wherein the bulk density may vary as a function of the water content of the support. The bulk density of water-comprising support particles is preferably in the range from 500 g/l to 1000 g/l, more preferably in the range from 600 g/l to 950 g/l and particularly preferably in the range from 650 g/l to 900 g/l. In the case of supports which have no water or a very low water content, the bulk density is preferably from 250 g/l to 600 g/l.

The support material can be partially or completely modified prior to activation. The support material can, for example, be treated under oxidizing or non-oxidizing conditions at temperatures of from 200 to 1000 ⁰ C, if appropriate in the presence of fluorinating agents such as ammonium hexafluorosilicate. In this way, it is possible to vary, inter alia, the water and/or OH group content. The support material is preferably dried at from 100 to 200 ⁰ C under reduced pressure for from 1 to 10 hours before being used.

Suitable support materials are also known in commerce and are commercially available.

The step of thermally treating the catalyst precursor (step b)) is carried out in such a manner that a maximum temperature of from 350 to 1050 ⁰ C, preferably from 400 to 95O ⁰ C, more preferably from 450 to 850°C, particularly preferably from 500 to 800°C, is not exceeded. The thermal treatment of the catalyst is carried out for at least part of the time in an oxidizing atmosphere, the chromium compound applied being converted completely or partly into the hexavalent state, i.e. being activated. The time of the thermal treatment in an oxidizing atmosphere is preferably at least 180 minutes, in particular at least 300 minutes, in order to achieve substantial oxidation to Cr Vl.

The choice of activation temperature(s) is predetermined by the properties of the polymer to be prepared and the activity of the catalyst. An upper limit is imposed by the sintering of the support and a lower limit is imposed by an unsatisfactorily low activity of the catalyst. The activation is preferably carried out at a temperature which is at least 20-100 ⁰ C below the sintering temperature. The other influences of the calcination conditions on the catalyst are known in principle and are described, for example, in Advances in Catalysis, Vol. 33, page 48 ff. The catalyst precursor is advantageously heated to the appropriate calcination temperature in a water- free gas stream comprising oxygen in a concentration of above 10% by volume and maintained under these conditions. In addition to the oxidative activation, a preceding or subsequent thermal treatment in an inert atmosphere can also be carried out.

Various types of activators can be used for the thermal treatment step b). The activation can be carried out in a fluidized bed and/or in a stationary bed, without being restricted thereto.

Preferably a thermal activation is carried out in fluidized-bed reactors. The reactors can be operated continuously or batchwise.

The activation time is at least 25 hours in order to achieve the effect according to the invention. The activation time is preferably not more than 100 hours, more preferably from 26 hours to 99 hours, more preferably from 27 hours to 98 hours, more preferably from 28 hours to 97 hours, more preferably from 29 hours to 96 hours, more preferably from 30 hours to 95 hours, more preferably from 31 hours to 94 hours, more preferably from 32 hours to 93 hours, more preferably from 33 hours to 92 hours, more preferably from 34 hours to 91 hours, more preferably from 35 hours to 90 hours, more preferably from 36 hours to 89 hours, more preferably from 37 hours to 88 hours, more preferably from 39 hours to 87 hours, more preferably from 40 hours to 86 hours.

The activation time is particularly preferably from 1680 minutes (28 hours) to 50 hours, more particularly preferably from 1800 minutes (30 hours) to 45 hours.

When the hold time is increased, the activation temperature should be reduced correspondingly in order to produce a product having the same HLMI potential. The extent to which the activation temperature has to be adapted can and must be determined experimentally by means of a few experiments.

The overall thermal treatment step is preferably carried out in at least three steps:

i) heating up, ii) holding above one or more temperatures above 35O ⁰ C, and iii) cooling down.

The activation time is thus preferably made up of the time period of i) and iii) which is above 300 ⁰ C, and ii). In the case of a batch process, such reaction conditions can be met by carrying out the steps in subsequent manner. In the case of a continuous process, this can be achieved by the provision of appropriate zones in a continuously supplied reactor with appropriate residence time of the reaction material.

The heating-up period preferably comprises essentially two periods, viz. drying at temperatures below 300 ⁰ C (heating-up period I) and the time above a temperature of 300 ⁰ C which contributes to the activation (heating-up period II).

The heating-up period Il makes a given contribution to the activation time, in particular in the case of relatively large batches in the industrial production of chromium catalysts. The activation in batch operation is therefore preferably carried out using an activator charge of at least 50 kg,

more preferably at least 100 kg, more preferably at least 200 kg, particularly preferably at least 300 kg.

The heating-up period is generally from 15 minutes to 1400 minutes, in particular from 60 minutes to 960 minutes. It is more preferably from 120 minutes to 900 minutes, particularly preferably from 120 minutes to 600 minutes. It is generally the case that the larger the amount of catalyst, the longer the heating-up period.

The heating-up period I is generally from 5 minutes to 360 minutes, in particular from 30 minutes to 300 minutes. It is more preferably from 60 minutes to 270 minutes, particularly preferably from 120 minutes to 240 minutes. In general, a higher water content of the catalyst precursor makes a longer heating-up period necessary to ensure sufficient drying, while in the case of a lower water content, a shorter heating-up period is required. Furthermore, a higher hold temperature generally requires a longer heating-up period.

The heating-up period Il is generally from 10 minutes to 900 minutes, in particular from 60 minutes to 840 minutes. It is more preferably from 120 minutes to 780 minutes, particularly preferably from 300 minutes to 720 minutes.

Both the heating-up period I and also the heating-up period Il can take place in an inert or oxidizing atmosphere. It is preferable that the heating-up period I takes place in an inert atmosphere and to switch over to an oxidizing atmosphere in the heating-up period II. It is particularly preferable that both the heating-up period I and the heating-up period Il take place in an oxidizing atmosphere.

The subsequent holding for a predetermined hold period is the significant and preferably longest part of the activation. Here, the temperature can preferably be maintained at a fixed temperature for the entire time or a predetermined temperature profile can also be followed in a temperature range. Furthermore, it is possible to provide a plurality of hold times at different temperatures. The hold period can also be very short if slow heating up and cooling down again is carried out over the entire activation time, so that the temperature is above 300 ⁰ C for a long time. The hold period is generally at least 10 minutes and an upper limit is imposed only by the consideration that an excessively long hold time is inconvenient from an economic point of view. The hold period is preferably from 300 minutes to 2400 minutes, more preferably from 600 minutes to 2100 minutes, more preferably more than 900 minutes, more preferably from 900 minutes to 1800 minutes. The hold period is particularly preferably more than 1200 minutes, more preferably from 1200 to 1500 minutes.

The maximum temperature during the thermal treatment can be from 350 to 1050 ⁰ C; the sintering temperature of the support should not be exceeded. The maximum temperature is preferably from 450 to 850 ⁰ C, in particular from 500 ⁰ C to 800°C.

The maximum temperature generally corresponds to the hold temperature or the highest hold temperature when a plurality of hold times is used.

The hold period is followed by the cooling-down period. In the simplest case, particularly in the case of relatively small reactors, this can take place by unregulated release of energy to the environment. In the case of relatively large reactors in particular, a controlled cooling-down program can be employed so as to obtain better reproducibility. The cooling-down period is generally from 10 minutes to 600 minutes, preferably from 20 to 500 minutes. Also the cooling- down period can be divided into two periods, viz. one above 300 ⁰ C (cooling-down period I) which contributes to the activation time and one below 300 ⁰ C (cooling-down period II). After sufficient cooling, the finished catalyst can be taken from the reactor.

Also the cooling-down period can be carried out in an inert or oxidizing atmosphere. It is preferable that the cooling-down period I takes place in an oxidizing atmosphere and to switch over to an inert atmosphere in the cooling-down period II.

After activation, reduction of the activated chromium catalyst, for example by means of reducing gases such as CO or hydrogen, preferably at from 350 to 950 ⁰ C, can be carried out if appropriate in order to obtain the actual catalytically active species. However, the reduction can also be carried out only during the polymerization by means of reducing agents such as ethylene, metal alkyls and the like present in the reactor.

It has been found that the influence of the activation time on product properties other than the HLMI potential can be divided into two ranges:

In one range having relatively short activation times of less than 25 hours, it is possible to produce catalysts which give polymers having virtually identical product properties by varying the hold time and the activation temperature. On the other hand, in the range having prolonged activation times above 25 hours, products having different properties, in particular a narrower molar mass distribution, can be produced at a constant HLMI potential. It has been found that an increase in the hold time in combination with an appropriate decrease in the activation temperature enables the productivity of the catalyst to be increased, the molar mass distribution to be made narrower and the number average molar mass M _n to be increased considerably at the same HLMI potential.

The number average molar mass (and thus the tensile impact strength a _zk ) has hitherto been increased by addition of fluorine compounds during the activation. Since the process of the invention also opens up a new opportunity for increasing the number average molar mass, the two measures (an increase in the hold time and addition of fluorine compounds) are preferably combined in order to obtain products having an improved tensile impact strength (a _tn ). As an alternative, the process of the invention is suitable for producing chromium catalysts which can replace fluorinated catalysts.

Furthermore, products were hitherto produced using chromium catalysts with subsequent reduction by means of CO. The CO likewise effects a significant increase in the number average molar mass and thus an increase in the tensile impact strength, but the productivity is reduced at the same time. The process of the invention can represent an alternative to reduction by means of CO. The catalyst is therefore preferably not reduced prior to the polymerization.

In particular, the process can be combined with the polymerization or copolymerization of ethylene. As comonomers, preference is given to using C ₃ -C ₈ -I -olefins, in particular 1-butene, 1-pentene, 1-hexene and/or 1-octene. Particular preference is given to a process in which ethylene is copolymerized with 1-hexene or 1-butene.

The process can be carried out using all industrially known polymerization processes at temperatures in the range from 0 to 200 ⁰ C, preferably from 25 to 150°C and particularly preferably from 40 to 130 ⁰ C, and under pressures of from 0.05 to 10 MPa and particularly preferably from 0.3 to 4 MPa. The polymerization can be carried out batchwise or preferably continuously in one or more stages. Solution processes, suspension processes, stirred gas-phase processes and gas-phase fluidized-bed processes are all possible. Processes of this type are generally known to those skilled in the art. Among the polymerization processes mentioned, gas-phase polymerization, in particular in gas-phase fluidized-bed reactors, solution polymerization and suspension polymerization, in particular in loop reactors and stirred tank reactors, are preferred.

In suspension polymerizations, the polymerization is usually carried out in a suspension medium, preferably in an inert hydrocarbon such as isobutane or mixtures of hydrocarbons or else in the monomers themselves. Suspension polymerization temperatures are usually in the range from -20 to 115°C, and the pressure is in the range from 0.1 to 10 MPa. The solids content of the suspension is generally in the range from 10 to 80%. The process can be carried out either batchwise, i.e. in stirring autoclaves, or continuously, e.g. in tube reactors, preferably in loop reactors.

Suitable suspension media are all media which are generally known for use in suspension reactors. The suspension medium should be inert and be liquid or supercritical under the reaction

conditions and should have a boiling point which is significantly different from those of the monomers and comonomers used in order to allow recovery of these starting materials from the product mixture by distillation. Customary suspension media are saturated hydrocarbons having from 4 to 12 carbon atoms, for example isobutane, butane, propane, isopentane, pentane and hexane, or a mixture of these which is also known as diesel oil.

In loop reactors, the polymerization mixture is pumped continuously through a cyclic reactor tube. The pumped circulation firstly results in continual mixing of the reaction mixture and, at the same time, distributes the catalyst introduced and also the monomers fed in through the reaction mixture. Secondly, the pumped circulation prevents sedimentation of the suspended polymer. The pumped circulation also promotes the removal of the heat of reaction via the reactor wall. In general, these reactors consist essentially of a cyclic reactor tube having one or more ascending and one or more descending legs which are enclosed by cooling jackets for removing the heat of reaction and also horizontal tube sections which connect the vertical legs. The impeller pump, the catalyst and monomer feed facilities and the discharge device, in general the settling legs, are usually installed in the lower tube section. However, the reactor can also have more than two vertical tube sections, so that a serpentine arrangement is obtained.

The discharge of the polymer from the loop reactor is generally made discontinuously in so-called settling legs. These settling legs are vertical projections projecting from the lower part of the reactor tube in which the polymer particles can sediment. After the sedimentation of the polymer has reached a given extent, a valve at the lower end of the settling legs is opened briefly and the polymer which has settled is discharged discontinuously.

In a preferred embodiment, the suspension polymerization in a loop reactor is carried out at an ethylene concentration of at least 10 mol%, preferably 15 mol%, particularly preferably 17 mol%, based on the suspension medium. Here, the suspension medium is not the introduced suspension medium such as isobutane alone, but the mixture of the suspension medium introduced with the monomers dissolved therein. The ethylene concentration can easily be determined by gas- chromatographic analysis of the suspension medium.

Furthermore, the suspension polymerization process in a loop reactor is preferably carried out at solids concentrations of more than 53% by weight, based on the total mass of the contents of the reactor, in order to achieve a very high production capacity. In one embodiment, the high solids concentration is achieved by the diameter of the cyclic reactor tube varying by more than 10%, based on the predominant reactor tube diameter. The tube diameter should preferably vary by at least 20%, even better by at least 30% and very particularly preferably by at least 50%. A widening of the reactor tube in the region of the impeller pump for construction reasons should not be taken into account here, since such a widening serves predominantly to accommodate the

impeller in the reaction tube and highly turbulent flow prevails in this region in any case. Rather, the invention was based on the observation that, contrary to prevailing opinion, deliberate nonuniform flow of the polymerization mixture in the region of the reaction tube outside the impeller region, too, makes an increase in the solids concentration in the reactor possible. This effect appears, without being restricted to this hypothesis, to be due to more effective mixing of the heterogeneous reaction mixture. In particular, the monomer, e.g. ethylene, fed in clearly becomes more rapidly distributed in the reaction mixture, dissolves more quickly in the suspension medium and is available to an increased extent for the polymerization. The removal of the heat of reaction also appears to be aided, since the disturbance of the flow increases motion transverse to the flow direction, i.e. in the direction of the cooled reactor wall, which is the case to only a very limited degree for uniform plug flow.

The high solids concentrations can also be achieved without continuous discharge of the polymer product. One variant of the process of the invention comprises discontinuously discharging the polymer formed from the reactor and carrying out the polymerization at an average solids concentration in the reactor of more than 45% by weight, based on the total mass of the contents of the reactor. Under these conditions, the solids concentration in the reactor is preferably above 50% by weight, particularly preferably above 55% by weight.

A preferred polymerization process is a process performed in a horizontally or vertically stirred or fluidized gas-phase reactor.

Particular preferred is gas-phase polymerization in a gas-phase fluidized-bed reactor in which the circulated reactor gas is introduced at the lower end of a reactor and is taken off again at the upper end thereof. When α-olefins are employed for the polymerization, the circulated reactor gas is usually a mixture of the 1 -olefin to be polymerized, if desired a molecular weight regulator such as hydrogen, and inert gases such as nitrogen and/or lower alkanes such as ethane, propane, butane, pentane or hexane. Preferably propane, if appropriate in combination with further lower alkanes, is used. The velocity of the reactor gas has to be sufficiently high firstly to fluidize a mixed bed of particulate polymer which is present in the tube and serves as polymerization zone and secondly to remove the heat of polymerization in an effective manner (non-condensed mode). The polymerization can also be carried out in the condensed or supercondensed mode, in which part of the circulating gas is cooled to below the dew point and is recirculated as a two-phase mixture to the reactor, so that the enthalpy of vaporization is additionally used for cooling the reaction gas.

In gas-phase fluidized-bed reactors, it is advisable to work at pressures of from 0.1 to 10 MPa, preferably from 0.5 to 8 MPa and in particular from 1.0 to 3 MPa. In addition, the cooling capacity required depends on the temperature at which the (co)polymerization in the fluidized bed is

carried out. The process is advantageously carried out at temperatures of from 30 to 160 ⁰ C, particularly preferably from 65 to 125 ⁰ C, with temperatures in the upper part of this range preferably being set for copolymers of relatively high density and temperatures in the lower part of this range preferably being set for copolymers of relatively low density.

The temperature during steady-state operation of the reactor is particularly preferably in a range delimited by an upper limit given by equation I:

6d'

7^ = 170 + (I) 0.84 - d

and a lower limit given by equation II:

7 3d'

T _m = 173 + (II) ^ω 0.837 - d'

where the variables have the following meanings:

T _RH = highest reaction temperature in °C T _RL = lowest reaction temperature in ⁰ C d' = absolute value of the density d of the polymer to be prepared in g/cm ³ .

According to this definition, the reaction temperature for the preparation of a polymer having a predetermined density d must not exceed the value defined by equation I and not be below the value defined by equation II, but has to be between these limit values.

Furthermore, it is possible to use a multizone reactor in which two polymerization zones, which can be subject to different polymerization conditions, are linked to one another and the polymer is passed a number of times alternately through these two zones. Such a reactor is described, for example, in WO 97/04015 and WO 00/02929.

The different or identical polymerization processes can, if desired, be connected in series and thus form a polymerization cascade. A parallel reactor arrangement with two or more identical or different processes is also possible. However, the polymerization is preferably carried out in only one single reactor.

To determine the parameters used here, the following methods were used:

The density of the polymer samples was determined in accordance with DIN EN ISO 1183-1 , variant A.

The determination of the intrinsic viscosity η, which indicates the limit value of the viscosity number on extrapolation of the polymer concentration to zero, was carried out using an automatic Ubbelohde viscometer (Lauda PVS 1) using a concentration of 0.001 g/ml in decalin as solvent at 135°C in accordance with ISO 1628-1 :1998.

The determination of the molar mass distributions and the average values M _n , M _w and M _w /M _n derived therefrom was carried out by means of high-temperature gel permeation chromatography using a method based on DIN 55672 on a WATERS 150 C with the following columns connected in series: 3x SHODEX AT 806 MS, 1x SHODEX UT 807 and 1x SHODEX AT-G under the following conditions: solvent: 1 ,2,4-trichlorobenzene (stabilized with 0.025% by weight of 2,6-di- tert-butyl-4-methylphenol), flow: 1 ml/min, 500 μl injection volume, temperature: 135°C, calibration using PE standards. Evaluation was carried out using WIN-GPC.

The melt flow rate was determined in accordance with ISO 1133 at a temperature of 19O ⁰ C under a weight of 2.16 kg (MFR ₂ , Ml) or 21.6 kg (MFR ₂₁ , HLMI).

The determination of the tensile impact strength was carried out in accordance with ISO 8256 (1997) /1A at -30°C. The test specimen was produced from the pelletized material.

All documents cited are expressly incorporated by reference into the present patent application. All percentages in this patent application are by weight based on the total weight of the respective mixtures, unless indicated otherwise.

The invention is illustrated below with the aid of examples, without being restricted thereto.

The experiments were all carried out using the catalyst Avant C 273 (Basell). The activations were carried out in a steel activator (Procedyne Corp.) having a volume of 32 liters and a maximum loading of 2 kg, with the level of the activation temperature and the length of the hold time being varied so that comparable HLMI values were obtained in the subsequent polymerization. The same activation profile was always used for heating up and cooling down. The hold time comprised the hold at the maximum temperature. The heating-up period I had in each case a duration of 240 minutes, and the heating-up period Il had a duration of 150 minutes. The heating- up period I took place in an inert atmosphere (nitrogen), and the heating-up period Il and the hold period took place in an oxidizing atmosphere (air). The cooling-down period I took place in an oxidizing atmosphere down to 300 ⁰ C and the atmosphere was subsequently switched over to an

inert atmosphere (nitrogen) in cooling-down period II. The respective hold time and the activation time are shown in Table 1.

The polymerization was carried out in a 0.2 m ³ loop reactor. The setting of the melt flow rate (HLMI) and density was effected via the reactor temperature or the hexene concentration. lsobutane served as suspension medium. The internal reactor pressure was 3.9 MPa, and the reactor density was about 577 kg/m ³ . The centrifugal pump for mixing the contents of the reactor was operated at 2100 rpm. The output was about 26 kg/h. Stadis 450 was metered into the reactor as antistatic in a concentration of about 8 ppm based on the output. The polymers obtained were stabilized with 1500 ppm of Irganox B220.

The most important reactor and product data are shown in Table 1 , with the product data being determined after peptization. A catalyst (hold time 10 h at 600 ⁰ C) for preparing a 42..A product having an HLMI of 6.3 g/10 min served as a reference (Example C1). To set comparable HLMI values at predetermined extremely different hold times of 0 and 72 h, the following activation settings were required: on decreasing the hold time to 0 h (i.e. heating up and immediately cooling down again), the activation temperature had to be increased to 65O ⁰ C. Increasing the hold time to 72 h led to an activation temperature of 560 ⁰ C.

The preparation of a polyethylene 42..A (target ranges: HLMI: 5.4 - 6.8 g/10 min, density: 0.945 g/cm ³ ) using the three catalysts proceeded without problems. The reactor temperature was 104 ⁰ C, the ethylene concentration was 9.1 mol% and the hexene concentration was 0.33 mol%, which led to a product having a melt flow rate of 6.1 g/10 min and a density of 0.946 g/cm ³ . When a change was made to the two other catalysts, products having similar melt flow rates and densities could be produced as expected at comparable reactor parameters. The conditions and results are shown in Table 1.

Table 1 : Reactor and product data

* only heating up and cooling down

** after peptization in accordance with ISO1133 (190°C/21.6 kg)

The data show that, despite extremely different activation profiles, the three catalysts have the same HLMI potential.

The molecular data of Comparative Examples C1 and C2 are also comparable in respect of the molar mass distribution. Example 1 , however, which is prepared using the catalyst having a significantly increased hold time of 72 h at 56O ⁰ C, is distinguished by a quite high M _n (30 000 vs. 20 000 g/mol) and an associated narrow molar mass distribution (12 vs. 16 or 17).

In addition, an increase in the activation time leads to a significant increase in the productivity. This is surprising since, at a constant activation time, the activity is known to decrease significantly with a decrease in the activation temperature.