Method of and apparatus for polymerising alpha-olefins in gaseous phase

The subject of the invention is a gas-phase olefin polymerization process carried out in a device with two connected polymerization zones into which one or more α-olefins of the formula CH2 = CHR are introduced under polymerization conditions in the presence of a catalyst and from which the produced polymer is collected . In the process according to the invention, the growing polymer particles flow through the first polymerization zone under rapid fluidization conditions, leave this first zone and enter the second polymerization zone, through which they flow in concentrated form, under the influence of gravity, leave the second zone and are re-introduced into the zone the first, thereby establishing polymer circulation between these two zones.

The development of Ziegler-type Natta catalysts with high activity and selectivity, and in an ever-growing number of applications - of metallocene type, has led to a wide spread of the industrial scale process in which gas phase olefin polymerization is carried out in the presence of a solid catalyst. Compared to liquid suspension, more conventional technology (monomer or monomer / solvent mixture), this technology presents the following benefits.

Operational flexibility. The reaction parameters can be optimally selected for the properties of the catalyst and product and are not limited by the physico-chemical properties of liquid mixtures of reaction components (usually containing hydrogen as a chain transfer agent).

Production range extension. The effect of swelling of growing polymer particles and dissolution of the polymer fraction in a liquid medium significantly limits the range of production for all types of copolymers.

Minimization of operations following polymerization. The polymer is obtained directly from the reactor as a dry solid and requires simple operations to remove dissolved monomer and deactivate the catalyst.

All technologies invented to date for the polymerization of α-olefins in the gas phase provide for the maintenance of a polymer bed through which reaction gases flow, the bed is kept in suspension either by mechanical mixing (mixed bed reactor) or by fluidization obtained by recycling the reaction gases as such ( fluidized bed reactor). In both types of reactors, the monomer composition around the polymer particles during the reaction is kept substantially constant by the agitation produced. These reactors are very close to the ideal behavior of the reactor referred to as "continuous stirred tank reactor" (CSTR), making it relatively easy to control the reaction and thus ensuring uniformity of product quality when operating under set conditions.

As for the most widely existing industrial installations to date, fluidized bed reactors operate in "bubbling" conditions. The polymer is compacted in a vertical cylindrical zone. The reaction gases coming out of the reactor are compressed with a centrifugal compressor, cooled and recycled together with prepared monomers and appropriate amounts of hydrogen, from the bottom of the bed through the feeding device. The entrainment of solid particles by gas is limited by appropriate dimensioning of the upper part of the reactor (free space, i.e. the space between the bed surface and the gas collection), where the gas velocity is reduced and, in several embodiments, by supplying cyclones in the output gas line.

The circulating gas flow rate is set to provide a fluidization rate within the appropriate range, below the minimum fluidization rate (critical speed) and below the "transport rate" - particle entrainment rate. The heat of reaction is only removed by cooling the circulating gas. Catalyst components are introduced continuously. The composition of the gas phase controls the composition of the polymer. The reactor works under constant pressure, usually in the range of 1-3 MPa. The reaction kinetics are controlled by the addition of an inert gas.

A significant contribution to the reliability of α-olefin polymerization technology in the fluidized bed reactor was the introduction of a suitably prepared spheroidal catalyst with controlled dimensions and the introduction of propane as a diluent (WO 92/21706).

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Fluidised bed technology has limitations, several of which are discussed below.

Removal of reaction heat.

The maximum gas velocity during fluidization is somewhat narrow (which is already set by the volumes of separation reactors that are equal to or greater than those established for the fluidized bed). Depending on the heat of reaction, polymer dimensions and gas density, the reactor efficiency limit (expressed as hourly output per unit of reactor section) is inevitably achieved when working with inlet gas temperatures higher than the desired dew point of the gas mixture. This limit may lead to a reduction in the efficiency of the installation, in particular during the copolymerization of ethylene with higher α-olefins (hexene, octene), carried out in the presence of conventional Ziegler-Natta catalysts, requiring a gas composition rich in such olefins.

A number of ways have been proposed to overcome these limitations, traditional technologies have been proposed in the field of heat removal, based on the partial condensation of circulating gases, and the use of latent heat of condensate evaporation to control the temperature of the reactor (European patent EP-89691, US patent 5,352,749, WO 94 / 28032). Despite the technical importance of the role, all of these systems proposed for the introduction of rules for the implementation of fluidized bed reactor operations are critical.

In particular (and regardless of the problems associated with the distribution of the wet solid into the space below the scattering grid), the technology used in EP-89691 and US 5,352,749 reduces turbulence generated by the grid to disperse liquid over the polymer. A possible symptom of coalescence in space can cause an increase in the uncontrolled symptom of poor liquid dispersion including the formation of agglomerates that may not be dispersed, especially in the case of polymers that tend to stick together. The distinctive criteria given in US 5,352,749 reflect situations under steady-state conditions, but do not provide any practicable indications in situations of transient "quenching of the reaction" which may lead to loss of fluidization, with further decline of the reactor bed.

The method described in the reference WO 94/28032 relates to the separation of condensate and its dispersion above the grate by means of special appropriately arranged nozzles. In fact, the condensate inevitably contains solid particles under reaction conditions, the concentration of which can be very high in a small amount of condensate. Moreover, the inherent difficulty of evenly dispersing the suspension over a series of nozzles may compromise the operation of several of them, and blocking one nozzle may adversely affect the dispersion of the evaporation liquid in the reactor section in question. It is also clear that the efficiency of work depends on the intensive circulation of solid particles in the reactor and below the injection points, it is reduced by the image of the gas flow rate caused by large amounts of condensate. Further, any maintenance on the nozzle requires complete shutdown of the reactor. Molecular Weight Spread.

As already stated, the fluidized bed exhibits behavior directly comparable to an ideal stirred reactor (CSTR). It is well known that during continuous polymerization of α-olefins in a single mixing step (which also requires a constant monomer composition and chain transfer agent, usually hydrogen) in the presence of Ziegler-Natta type catalysts, polyolefins with a relatively narrow molecular weight distribution are obtained. This property manifests itself even more if metallocene catalysts are used. The range of molecular weight distribution affects both the rheological behavior of the polymer (and hence the ability to process for the alloy) as well as the final mechanical properties of the product and is a particularly important property for ethylene (co) polymers.

In order to extend the range of molecular weight distribution, the process carried out in a series of reactors arranged in a series, in which each of them is able to work with at least different hydrogen concentration, gained industrial significance. The problem that is usually to be expected when using these processes, if a very wide molecular weight distribution is required, is insufficient product homogeneity. In particular, the homogeneity of the material in blow-molding processes and in the production of C6 is critical

187 165 foils, in which the presence of even small amounts of heterogeneous material leads to the occurrence of unmelted particles in the film ("fish eyes"). In EP-574,821, a system of two reactors was proposed, operating under different polymerization conditions with two-sided recirculation of the polymer between the reactors. Even if this idea is appropriate to solve the problem of product heterogeneity, the results of the experiment have shown that such a system entails investment costs and a number of operational complications. In other cases, a polymer with a high molecular weight distribution is obtained, using mixtures of different Ziegler-Natta catalysts in a single reactor, each catalyst being prepared to give different hydrogen sensitivity. It is clear that at the exit of the reactor a mixture of granules is obtained, each of which is its own individuality. This way it is difficult to achieve product homogeneity.

Product embedding.

The technology of polymerization of α-olefins in gas phase reactors has developed rapidly in recent years and the number of types of polymers obtainable by this method has increased significantly. In particular, in addition to ethylene and propylene homopolymers, a wide range of copolymers can be industrially produced, for example, statistical polymers of propylene / ethylene, ethylene / propylene / higher α-olefins and propylene / higher α-olefins, low and very low density polyethylene (LLDPE, VLDPE ), modified with higher α-olefins containing 4 to 8 carbon atoms, heterophasic high-impact copolymers, obtained by growth on active catalyst centers, in subsequent stages, one or more of the polymers listed above and ethylene propylene or ethylene butene rubbers, and EPR (ethylene propylene) and EpDm ethylene propylene diene rubbers).

For polymers produced in the gas phase, the modulus of elasticity ranges from 2300 MPa to below 100 MPa, and the xylene soluble fraction ranges from 1% to 80%. Flowability, compactability and adhesive properties change extremely much as a function of degree of crystallinity, molecular weight and composition of different polymer phases. Many of these products remain in the form of granules and in a flowable form (and hence suitable for processing), as long as they are kept in the fluid phase or stream, which is conditions in which static forces between individual solid particles have no effect . They tend to be more or less violent to clump together and form agglomerates, if they are left to sink or if they are concentrated in stationary zones, this phenomenon is particularly intensified under reaction conditions, where due to the combined effect of temperature and a large amount of dissolved hydrocarbon, the polymer is particularly soft, able to thicken and compact and stick. The description of soft and sticky polymers is effectively provided in EP-348,907 or US 4,958,006.

The simplest solution for discharging the polymer from the reactor is direct discharging from the fluidized bed through an adjustable valve. This type of discharge combines simplicity with the advantage of not creating a stationary zone. In the area of ​​sufficiently low pressure (in the range of 0.5 · 10⁵-3Τ0⁵ Pa) kept below the unloading valve, the reaction actually stops either by lowering the temperature, caused by the evaporation of monomers dissolved in the polymer, or also by causing the low partial pressure of the monomers in the gas: thus avoiding any risk in the equipment for the downstream reactor.

Nevertheless, it is known that the amount of gas discharged with the polymer from the fluidized bed through the outlet reaches very high values ​​as a function of reactor pressure, fluidization rate, solid phase density in the bed and so on (for example: Massimilla, "Flow properties of the fluidized dense phase "- this is" Flow properties of a dense fluidization phase, "in Fluidyzation pp. 651-676, Dawidson & Harrison, Academic, New York 1971). Large amounts of gas discharged from the polymer reactor represent both investment and operating costs, as it is necessary to re-compress this gas to the value of the pressure in the reactor from the receiving pressure. In many industrial applications, discontinuous collection systems are installed, including at least two hoppers

187 165 .7 dumpers working alternately. For example, US 4,621,952 describes an unloading system in which the polymer is transferred discontinuously at a high pressure difference from the reactor to the sludge tank. The momentum of the polymer, which during the filling phase, collides first with the walls of the sludge tank and then, with the polymer bed, compacting the material, which loses its flowability. During the filling phase, the pressure in the collecting tank increases rapidly to the value of the reactor pressure and the temperature does not change significantly. The reaction proceeds adiabatically at high speed. For soft and sticky products, this easily leads to the formation of agglomerates that cannot be granulated, with the resulting difficulties of unloading into the receiving tank below. Similar observations apply to US 4,703,094.

The limitations of the discontinuous system are clearly revealed in the proposition of the complicated continuous system. Japanese Patent JP-A-58 032,634 provides for the installation of an internal screw in a reactor to thicken the polymer in the discharge direction. US 4,958,006 proposes the installation of an extruder, the screw of which is introduced directly into the interior of the fluidized bed reactor. Regardless of the complexity and difficulties in industrial use, the proposed systems are in any case completely inappropriate for introducing the polymer into the next reaction stage.

The state of rapid fluidization is obtained if the speed of the fluidizing gas is greater than the rate of blowing particles out of the apparatus and is characterized in that the pressure gradient along the transfer direction is a monotonous function of the amount of solid injected, with the same amount of flow and fluidization gas density. In the fluidized bed technology of the prior art, the fluidization gas rate is kept below the blow-out rate to avoid entrainment and blow-out. The terms blow rate and state of rapid fluidization are known in the art. Their definitions are presented in D. Geldart, "Gas fluidization technology" p. 155 et seq. J.Wiley & Sons Ltd. 1986.

The present invention relates to a polymerization process for α-olefins of formula CH₂ = CHR, in which R is hydrogen or a hydrocarbon radical having 1 to 12 carbon atoms, in the gas phase, conducted in the first and interconnected second polymerization zone into which at least one of the α-olefins is introduced under reaction conditions, the presence of a catalyst and from which the polymer product is derived, according to the invention is characterized in that the growing polymer particles flow through the first of the polymerization zones under rapid fluidization, they leave the first polymerization zone and enter the second polymerization zone, through which they flow in concentrated form under gravity, leave the second polymerization zone and re-enter the first polymerization zone, whereby the circulation of the polymer between these two polymerization zones has ceased.

The conditions for rapid fluidization are determined by feeding a gas mixture containing at least one of the α-olefins of formula CH₂= CHR to the first polymerization zone and the gas mixture is fed to the first polymerization zone in the area below the point of polymer re-introduction into the first polymerization zone.

Preferably, the gas mixture is fed by using gas dispersing agents.

The polymer and the gas mixture leaving the first polymerization zone are led to the solid / gas separation zone, wherein the polymer leaving the solid / gas separation zone enters the second polymerization zone.

Control of the polymer circulating between the two polymerization zones is carried out by measuring the amount of polymer leaving the second polymerization zone.

The polymer produced is continuously withdrawn from the second polymerization zone.

The catalyst components are fed to the first polymerization zone.

Preferably, the first polymerization zone is fed by the catalyst in polymerized form.

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Preferably, any of these polymerization zones is fed with the catalyst in a dispersed form in the polymer slurry.

Preferably, any of these polymerization zones is fed with the catalyst in a dispersed form in the dry polymer.

The gaseous mixture leaving the solid / gas separation zone is compressed, cooled and sent to the first polymerization zone, and optionally, with the addition of prepared complementary monomers.

Part of the gaseous mixture leaving the solid / gas separation zone is used to transfer the polymer from the second zone to the first polymerization zone.

Part of the gaseous mixture leaving the solid / gas separation zone is compressed and sent to the second polymerization zone near the area where the polymer leaves this second zone.

The gas mixture leaving the solid / gas separation zone is cooled to a temperature below the dew point.

The first polymerization zone is cooled by external coolants.

The pre-prepared supplementary portion of the monomer or monomers is fed to the first polymerization zone in at least partially condensed form.

The speed of the fluidizing gas introduced into the first polymerization zone is between 2 and 15 m / s, preferably between 3 and 8 m / s.

In the presence of a controlled morphology catalyst, the polymer is obtained in the form of spheroidal particles with average dimensions between 0.2 and 5 mm, preferably between 0.5 and 3 mm.

Preferably an operating pressure between 0.5 and 10 MPa, preferably between 1.5 and 6 MPa is used.

Preferably, at least one inert gas is present in the polymerization zones at partial pressures between 5 and 80% of the total gas pressure.

Preferably, nitrogen or an aliphatic hydrocarbon having inert gas is used

2-6 carbon atoms, preferably propane.

Between the first and second polymerization zones an intermediate polymerization zone is placed, working with the fluidized bed.

The polymerization apparatus for α-olefins of the formula CH2 = CHR, in which R is hydrogen or a hydrocarbon radical having 1 to 12 carbon atoms, in the gas phase, according to the invention is characterized in that the first vertical cylindrical reactor is equipped with a catalyst feed line, and the second vertical cylindrical reactor is equipped with a polymer discharge system. The upper zone of the first reactor is connected through a first line to a solid / gas separator, which in turn is connected to the upper part of the second reactor. The bottom of the second reactor is connected through a second line to the bottom of the first reactor. In addition, the solid / gas separator is connected by a recirculation line for the gas mixture to the first reactor in the area at the bottom of the first reactor below the entry point of the second line.

Preferably the first reactor is equipped with gas scattering means located between the second line entry point and the area at the bottom of the first reactor.

Preferably, a first control valve is placed between the second reactor and the second line to control the polymer flow rate.

Preferably the first valve is a mechanical valve.

Preferably the first valve is a non-mechanical valve.

The catalyst feed line is connected via a third line to the first reactor at a point above the gas scattering agent.

The gas recirculation line is equipped with a compressor, cooling system and monomer introduction systems and molecular weight regulator.

The first line leaves the upper area of ​​the first reactor, stepping aside.

The upper area of ​​the first reactor is shaped as a truncated cone with the wider end located at the top.

The gas recirculation line is connected to the second line through the line at a location below the compressor.

Preferably the first reactor is equipped with external cooling means.

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According to the invention of the second polymerization zone, where the polymer flows in concentrated form under the action of gravity, high values ​​of solid density (solid density = kg polymer per m³ reactor occupied by the polymer), approaching the specific density of the polymer. A positive pressure advantage can thus be obtained along the flow direction, so that it becomes possible to reintroduce the polymer into the first polymerization zone without the help of special mechanical means. In this way, "annular" circulation is set up, which can be defined as the balance of pressure between the two polymerization zones and the loss of origin introduced into the system.

The solution according to the invention is described with reference to the drawings, in which Figure 1 is a schematic representation of the method according to the invention, Figure 2 is a schematic representation of the first embodiment of the method according to the invention, and Figure 3 is a schematic representation of the second embodiment of the method according to the invention.

According to Fig. 1, the growing polymer particles flow through the first polymerization zone 1 under rapid fluidization conditions along the direction of arrow 14. In the second polymerization zone 2, the growing polymer particles flow in a concentrated form under gravity along the direction of arrow 14 '. The two polymerization zones 1 and 2 are respectively connected by sections 3 and 5. The material balance is maintained by introducing monomers and catalyst and discharging the polymer.

Generally, the rapid fluidization conditions in the first polymerization zone 1 are determined by introducing a gas mixture containing one or more α-olefins CH2 = CHR (line 10) into zone 1. Preferably, the gas mixture is introduced below the point of polymer reintroduction into zone 1, with using appropriate gas distribution methods, such as a scattering grid, for example.

The gas flow rate to the first polymerization zone is higher than the particle blowing rate under operating conditions and is preferably between 2 and 15 m / s, more preferably between 3 and 8 m / s.

The control of the polymer circulating between the two polymerization zones can be carried out by measuring the amount of polymer leaving the second polymerization zone 2, using methods to control the flow of solids, such as, for example, mechanical valves (slide valve, V-ball valve and so on) or valves non-mechanical (valve L, valve J, check seal and so on). Generally, the polymer and gas mixture leaving the first polymerization zone 1 are transferred to the gas and solid separation zone 4. Separation of solids and gas can be accomplished by conventional separation methods such as, for example, a separator of the inertial or centrifugal type or a combination of both. The centrifugal separator (cyclone) can be of the axial, spiral, helical or tangential type.

From the separation zone 4, the polymer enters the second polymerization zone 2. After leaving the separation zone 4, the gas mixture is compressed, cooled and sent, respectively with the addition of prepared monomers and / or molecular weight regulators, to the first polymerization zone 1. This transfer can be accomplished via a gas mixture recirculation line 6 equipped with means for compressing 7 and cooling 8 and means for introducing monomers and molecular weight regulator 13.

The part of the gaseous mixture leaving the separation zone 4 can be sent after compression to the connecting zone 5 through line 9 to facilitate the transfer of the polymer from the second to the first polymerization zone.

Preferably, various catalyst components are introduced into the first polymerization zone 1 at any point in said polymerization zone 1. However, they can also be introduced at any point in said polymerization zone 2. Any type of catalyst used in olefin polymerization can be used in the process of the invention, because it is not important that it is in a certain physical state, and therefore catalysts can be used in both solid and liquid form, because unlike phase polymerization methods according to the prior art, the method according to the invention does not necessarily require the use of a catalyst of which at least one component

187 The nent must be in granular form, but may be carried out with catalysts in which the various components are in the form of a solution. For example, titanium, chromium, vanadium or zirconium based catalysts can be used, either in supported or unsupported form. Examples of catalysts that can be used are described in US Patent Nos. 4,748,272, 4,302,566, 4,472,520, and 4,218,339. Particularly useful catalysts are those with controlled morphology, described in US Patent Nos. 4,399,054, 5,139,985 and European Patent Descriptions EP-395,083, EP-553,805, EP-553,806 and EP-601,525, and generally catalysts capable of producing a polymer in the form of spheroidal particles having average dimensions between 0.2 and 5 mm, preferably between 0.5 and 3 mm. The process of the invention is particularly suitable for the use of metallocene catalysts, both in solution and on a support. Various catalyst components can be introduced at the same point or at different points in the first polymerization zone.

The catalyst can be introduced without pre-treatment or after pre-polymerization. If other stages of polymerization are located further along the course of the material, it is also possible to introduce polymerization zones with a catalyst dispersed in the polymer suspension, derived from a previously set bulk reactor, or a catalyst dispersed in a dry polymer from a previous gas phase reactor.

The concentration of the polymer in the reaction zones can be monitored by conventional methods known in the art, for example by measuring the pressure difference between two respective points along the axis of the polymerization zones or by measuring the density with radioactive detectors (e.g., y-radiation).

Operating parameters such as, for example, temperature are those normally used for polymerizing olefins in the gas phase, e.g. between 50 and 120 ° C.

The method of the invention has several advantages. The circular loop configuration allows the use of relatively simple reactor shapes. In practice, each reaction zone can be thought of as a cylindrical reactor with high elongation (height / diameter ratio). From a construction point of view, this particular geometry allows the use of high operating pressures that are not economical in conventional fluidized bed reactors. The method of the present invention can thus be carried out at operating pressures between 0.5 and 10 MPa, preferably between 1.5 and 6 MPa. The resulting high gas density favors both the heat exchange of individual particles and the total heat removal of the reaction. It is therefore possible to choose reaction conditions that support the reaction kinetics. Further, the reactor through which the polymer flows under rapid fluidization (first polymerization zone) can operate completely full at a polymer concentration that can reach or exceed 200 kg / m³. By cooperating with the second polymerization zone and taking into account more favorable kinetic conditions that can be established, the method of the invention allows obtaining unit yields (hourly yield per unit volume of reactor) much higher than the level obtained by conventional fluidized bed technology. In this way it is possible to match or even exceed the catalytic efficiency of conventional gas-phase processes, using equipment of much smaller sizes, with a significant saving in installation costs.

In the method of the invention, loading solid particles into the gas recovery line at the exit of the solid / gas separation zone and the possible presence of liquid leaving the cooler to the same line does not limit the efficiency of the first polymerization zone. Even if gas scattering methods are used, such as, for example, a grate, the speed of the transporting gas in the space below the grate is still high and such that it provides droplet entrainment of even large sizes and moist polymer particles, without creating dead spots. Assuming that the transport gas comes into contact with the hot polymer stream coming from the second polymerization zone, the evaporation of any liquid is indeed instantaneous. It is therefore possible to cool the gas mixture leaving the solid / gas separation zone to a temperature below the dew point to liquefy some of the gases. The gas / liquid mixture that forms is then introduced into the first polymerization zone, where it contributes to heat removal without taking into account the problems

187 165 and the limitations of the known state of the art and without the need for complicated devices to avoid them. In addition to this and / or to replace partial condensation of recirculated gases, the process of the invention opens a new path for removing heat of reaction. Proper geometry (high surface to volume ratio) in the polymerization zone along with rapid fluidization gives a significant external surface available for direct heat exchange in this zone (and hence the maximum heat transfer between the cooling liquid and the reaction system). Conveniently, additional or alternative heat exchange surfaces can be installed inside the reactor. The first polymerization zone can preferably be cooled by external cooling methods. The high turbulence associated with the conditions of rapid fluidization and the high gas density ensure in each case a very high heat transfer coefficient. Any condensation on the inner wall is removed continuously by strong radial and axial mixing of the polymer due to rapid fluidization conditions. In addition, this property makes the proposed technology suitable for operation when the second zone is fed directly from the bulk polymerization reactor located earlier. It is also possible to introduce some of the prepared monomers in the form of condensate without any difficulties. As far as heat removal is intensified, the capacities offered by the method according to the invention are much larger than according to the known state of the art and the difficulties associated with earlier technologies are thus overcome. In addition, the volumetric rates of circulating gas are not necessarily dependent on the heat exchange required.

Preferably, one or more inert gases are maintained in the polymerization zones in amounts such that the sum of the inert gas partial pressures is preferably between 5 and 80% of the total gas pressure. The inert gas may be a nitrogen or an aliphatic hydrocarbon with 2-6 carbon atoms, preferably propane . The presence of inert gas presents a number of benefits, it creates the possibility of inhibiting the reaction kinetics, while the overall reaction pressure is maintained at the same time, which is sufficient to keep the head of the circulating compressor at a low level and to ensure that the mass flow rate is sufficient for heat exchange to particles in the bed and through the condenser in a mixture of circulating gases, to remove the heat of reaction, which is not removed by surfaces.

In the process of the invention, the presence of inert gases presents further advantages so far as it creates the possibility of limiting the temperature rise in the second polymerization zone, which operates essentially in an adiabatic system, and also makes it possible to control the extent of the molecular weight distribution of the polymer, in particular during polymerization of ethylene. This is possible because, as previously stated, the polymer flows vertically down through the second polymerization zone in the form of a plastic flow (flow of whipped material) surrounded by limited amounts of entrained gas. As is known, the molecular weight of the polymer during ethylene polymerization is controlled by the hydrogen / ethylene gas phase ratio and to a lesser extent by temperature. In the presence of an inert gas, causing ethylene to be consumed in the reaction, and hydrogen only to a minimal extent, the ethylene / hydrogen ratio decreases along the polymer flow axis in the direction of movement, causing the same molecule to grow at a lower molecular weight polymer. An increase in temperature causes an increase in this effect. It is therefore possible by controlling the gas composition and residence time in both polymerization zones, controlling the distribution of the molecular weight of the polymer in an effective manner and maintaining maximum homogeneity of the product at the same time.

On the contrary, if it is desired to obtain a polymer with a relatively small molecular weight distribution, the above mechanism may be turned off, or it can be avoided by selecting the correct reaction conditions, for example by limiting the amount of inert gas or introducing the appropriate amounts of reaction gases and / or portions of the prepared monomer (s) in a suitable location for the second polymerization zone. Preferably, the gas to be introduced into the second polymerization zone can be taken from the gas mixture leaving the solid / gas separation zone, after which it is subjected to compression. The amount of gas introduced is set in such a range that the relative injection rate12

187 The liquefied gas, taking into account the solids flow rate, is kept below the minimum fluidization rate characteristic of the solid / gas system present in the second said polymerization zone. Under these conditions, the polymer flow down is not significantly disturbed. The operational flexibility of the method according to the invention is therefore total, the production of polymers with different molecular weight distribution can be controlled by the composition of the gases and, if necessary, by simply closing or opening the gas line valve.

Preferably, the polymer can be discharged from zones where the density of the solid is higher, for example from a suitable point in the second polymerization zone, where large amounts of thickened flowing polymer are available to minimize the amount of occluded gas. By switching on the control valve at a suitable point above the polymer exit region from the second polymerization zone, it becomes possible to continuously monitor the receipt of the produced polymer. The amount of gas accompanying the polymer is extremely small and only slightly larger than that obtained by the device with the setting of a series of hoppers at alternating periodic operation '. In this way, all the limitations of the discharge system of the known state of the art have been overcome, both in terms of the amount of occluded gas and the nature of the products being unloaded.

As already stated, the method of the invention can be combined with conventional technologies into sequences of multi-stage processes in which one or more stages using conventional technologies (in bulk or in the gas phase) are present before or after the polymerization sections operating in accordance with the present invention. or with a fluidized or mixed bed). Multi-stage processes using the procedure of the present invention are also possible.

It is therefore possible to combine the method according to the invention with conventional fluid phase gas phase technology by inserting between two polymerization zones as defined according to the present invention a bubble bed polymerization zone, i.e. with a fluidization gas rate higher than the critical and lower than the blow-out rate, while constantly maintaining the recycling loop characteristic of the method of the present invention. For example, one possible embodiment includes a solution in which the second polymerization zone comprises first and second sections. In the first (with respect to the direction of flow of the polymer) from said sections, the fluidization state is maintained by appropriate gas introduction and dispersion. In the second section connected to the first, the polymer flows in concentrated form under the influence of gravity. From the second section, the polymer is reintroduced into the first polymerization zone to form a circulation loop. By appropriately dimensioning the various zones, it becomes possible to achieve an extension of the molecular weight distribution of the polymer while maintaining all the advantages described above. The above example is only one possible embodiment of the solution according to the invention, which in its general definition contains at least a zone of rapid fluidization in combination with the zone where the polymer flows in concentrated form under the influence of gravity.

The process of the present invention is capable of producing a large number of olefin polymers without the disadvantages described above. Examples of polymers that can be made are:

- high density polyethylenes (HDPE polymers with average densities higher than 0.940) including ethylene homopolymers and copolymers of ethylene in α-olefins with 3 to 12 carbon atoms, .....

- low density linear polyethylenes (LLDPE polymers with average densities lower than 0.940) and very low density and ultra low density (VLDPE and ULDPE polymers having average densities lower than 0.920 to 0.880) including copolymers of ethylene with one or more α-olefins 3 to 12 carbon atoms,

- elastomeric terpolymers of ethylene and propylene with a small amount of diene or elastomeric copolymers of ethylene and propylene with an ethylene mer content of between 30 and 70% by weight,

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- isotactic polypropylene and crystalline copolymers of propylene with ethylene and / or other α-olefins with a content of propylene units greater than 85% by weight,

- heterophasic propylene polymers obtained by alternating polymerization of propylene and mixtures of propylene with ethylene and / or other α-olefins,

- atactic polypropylene and amorphous copolymers of propylene and ethylene and / or other α-olefins containing more than 70% by weight of propylene units,

poly-α-olefins, such as, for example, polybutene-1, poly-4-methylpentene-1, polybutadiene and other polydiene rubbers.

A further aspect of the present invention relates to a device for the gas-phase polymerization of α-olefins. The device according to the invention consists of a first vertical cylindrical reactor 20 equipped with a catalyst feed line 34, and a second vertical cylindrical reactor 30 equipped with a polymer discharge system 23. The upper zone of the first reactor 20 is connected through the first line 21 to the solid / gas separator 22, which in turn is connected to the upper part of the second reactor 30, the lower part of the second reactor 30 is connected via the second line 31 to the bottom of the first reactor 20. Separator solid / gas 22 is connected by a recirculation line for the gas mixture 36 to the first reactor 20 in zone 37 at the bottom of said first reactor 20 below the entry point of the second line 31. Preferably, the first reactor 20 is equipped with gas dissipation means 33, e.g. a grate, located between the entry point of the second line 31 and the area 37 at the bottom of this reactor. As an alternative, referring to fig. 3, the gas scattering means in the first reactor 60 can be replaced by a cylindrical line 65 through which the gas flows at a high speed and which is connected to the reactor 60 through a frustoconical section 62 whose angle of inclination to the vertical is preferably less than 45 ° and more preferably between 30 ° and 10 °. Preferably, both the catalyst (via line 66) and the polymer fed from the second reactor 70 via line 77 can be transported through this frustoconical connection.

The first valve 2A for controlling the polymer flow rate is usually inserted between the second reactor 30 and the second line 31. This valve 24 can be of both mechanical and non-mechanical type.

In the case where a gas scattering agent is used, several or all of the catalyst components may advantageously be injected through the third line 32 into said first reactor 20 at a point above the gas distribution center.

The gas recirculation line 36 can be equipped with a compressor 26, a cooling system 27 and a system for introducing, together or separately, monomers 28 and molecular weight regulator 29. Two cooling systems can be used, one before and one after the compressor.

It is recommended that the first line 21 leaves the top of the first reactor from the side. It has been observed that the lateral exit of the solid / gas mixture from the first reactor 20 contributes significantly to maintaining the dynamic stability of the entire reaction system.

The upper part of the first reactor 20 may have a cylindrical shape equal to the diameter of the reactor or preferably it may have a frustoconical shape with a widened end at the top. The first line 21 may be horizontal or have an inclination in the direction of gravity to facilitate the discharge of the polymer (system of lines 71 in Fig. 3). The second line 31 can be inclined downwards and can be connected (at a point immediately after the first valve 24) through the line 25, to the gas recycle line 36 at the point after the compressor 26. In this way the polymer flow is accompanied by a flow of gas under pressure from gas recycling lines, thus avoiding polymer dead zones in the line itself and at the point of entry into the reactor 20.

The system of connections between the lower zones of the reactors can also be the system described in Fig. 3, in which the circulation of the polymer is obtained through the valve L, pneumatic 74, controlled by gas taken from the recirculation line through the line 75. The valve L is connected to the line 77, leading to the first reactor 60, wherein said line 77 is connected via line 76 to recirculation line 81. Through this line, the polymer is re-introduced into the reactor 60 with a suitable gas stream from line 76.

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The first reactor 20 may advantageously have external cooling equipment 35, such as a shell-type heat exchanger.

Two possible embodiments of the invention are shown in Figs. 2 and 3. These embodiments are given only to illustrate the invention and do not limit its scope.

Referring to Fig. 2, 20 shows the first reactor operating under rapid fluidization conditions and 30 shows the second reactor through which the polymer flows in concentrated form under the influence of gravity. References 21 and 31 denote lines connecting the upper and lower parts of these two reactors, and 34 is the catalyst introduction line, 22 is a solid / gas separator, 23 is a polymer discharge system, 36 is a recycling gas line that connects said separator to area 37 at the bottom of the first reactor, 24 is a control valve to control the amount of polymer flow, 33 is a gas scattering device, 32 is a catalyst introduction line, 26 is the compressor and 27 is the recycle gas mixture cooling system, 28 and 29 are the monomer and molecular weight regulator injection systems, and 25 is the line connecting recirculation line 36 with line 31, 35 is the external cooling system of the first reactor 20.

With reference to Figs. 3, 60 shows a first reactor operating under high fluidization conditions and 70 is a second reactor through which the polymer flows in a concentrated form under gravity. References 71 and 77 denote lines connecting the upper and lower parts of these two reactors, 66 is the catalyst introduction line, 72 is a solid / gas separator, 73 is a polymer discharge system, 81 is a recycling gas line that connects said separator 72 to line 65 connected to the base of the first reactor 60 by section in the form of a truncated cone 62, 74 is a control valve L to control the amount of polymer flow, 79 is a compressor and 80 is a cooling system for recycled gas mixture, 63 and 64 are monomer introduction systems and molecular weight regulator, 75 is a line connecting recirculation line 81 with valve L 74, 76 is a line connecting recirculation line 81 with line 77, 78 is a line, which connects recirculation line 81 to the area at the bottom of the second reactor 70, and 61 is the external cooling system of the first reactor 20.

The following examples will further explain the invention without limiting its scope.

General polymerization conditions.

The polymerization was carried out continuously in an installation that contained a pre-section in which various catalyst components, a pre-polymerization section, and a gas-phase polymerization section were pre-mixed, the gas-phase polymerization being carried out in a reactor of the type described in Figure 2.

The solid catalyst component prepared according to the procedure described in Example 3 of EP-A-395083, consisting of triethylaluminum (TEAL) and the silane compound was pre-reacted in hexane at 10 ° C for 10 minutes in the pre-reaction vessel. The activated catalyst was introduced into the pre-polymerization section, where it was subjected to suspension polymerization of propylene, using propane as the dispersing agent. The monomer feed and reaction time were set to obtain the desired yield, in terms of the amount of g polymer per g solid catalyst component.

The prepolymer was continuously introduced into the gas phase polymerization apparatus. The installation, which is shown with reference to Fig. 2, consisted of two cylindrical reactors 20 and 30 connected by pipes 21 and 31. The reactor 20 was equipped with a heat exchanger 35. Rapid fluidization in the reactor 20 was obtained by recycling gas from a gas / solid separator 22 downstream of reactor 20 via gas recycling line 36. No additional gas scattering methods were used, the recycle gases were fed directly to the area 37 at the bottom of the reactor 20, below the feed point 31. The gas recycle line was equipped with a compressor and heat exchanger 27. The reactor for the preparation of the prepolymer slurry was connected to the reactor 20 at the point directly above the pipe entry point 31. The polymer circulation was controlled by a valve L 24 operating by means of a gas stream 25 taken from the recirculation line 36. Pre-prepared monomers were fed into the re-circulation line 36. The produced polymer was continuously withdrawn from the reactor 30 through a pipe 23. The total volume of the apparatus (i.e. reactors 20 and 30 and connecting zones 21 and 31) was 250 1.

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Example 1

Polypropylene was prepared against a catalyst containing dicyclopentyl dimethoxy silane (DCPMS) as a silane compound. At the gas phase polymerization stage, propane was used as the inert gas.

Main working conditions.

Catalyst pretreatment step

<td>- TEAL / solid component</td><td>(Wt.)</td><td> 8</td>

<td>- TEAL / DCPMS</td><td>^(wag.)</td><td> 3</td>

<td>Pre-polymerization step</td><td></td><td></td>

<td>- performance</td><td>(G / g)</td><td> 100</td>

<td>Gas phase polymerization</td><td></td><td></td>

<td>- temperature</td><td>(° C)</td><td> 85</td>

<td>- pressure</td><td> (•10⁵ pa)</td><td> 25</td>

<td>- propylene</td><td>(%moth)</td><td> 91</td>

<td>- propane</td><td>(%moth)</td><td> 8</td>

<td>- hydrogen</td><td>(%moth)</td><td> 1</td>

<td>- performance</td><td>(Kg / hm³)</td><td> 140</td>

<td>Product properties</td><td></td><td></td>

<td>- density by mass</td><td>(Kg / l)</td><td> 0,45</td>

Example 2

Hexene-modified LLDPE was prepared against a catalyst containing cyclohexylmethyl dimethoxy silane (CMMS) as a silane compound. During gas phase polymerization, propane was used as the inert gas. Main reaction conditions. Catalyst pretreatment step

<td>TEAL / Ti</td><td>(Wt.)</td><td> 120</td>

<td>TEAL / CMMS</td><td>(Wt.)</td><td> 20</td>

<td>Pre-polymerization step</td><td></td><td></td>

<td>- performance</td><td>(G / g)</td><td> 400</td>

<td>Gas phase polymerization</td><td></td><td></td>

<td>- temperature</td><td>(° C)</td><td> 75</td>

<td>- pressure</td><td>(1 () 5 Pa)</td><td> 24</td>

<td>- ethylene</td><td>(%moth)</td><td> 15</td>

<td>- hexene-1</td><td>(%moth)</td><td> 1,5</td>

<td>- hydrogen</td><td>(%moth)</td><td> 3</td>

<td>- propane</td><td>(%moth)</td><td> 80,5</td>

<td>- performance</td><td>(Kg / hr-m3)</td><td> 80</td>

<td>Product properties</td><td></td><td></td>

<td>- flow indicator E</td><td>(g / 10 min)</td><td> 1,4</td>

<td>- density</td><td>(G / cm3)</td><td> 0,908</td>

<td colspan="2">The temperatures quoted above were measured</td><td>»At the top of the reactor 30. Dew point of the mixture</td>

gas at operating pressure is 66 ° C. The cooling liquid flowed through the heat exchanger in such a way as to maintain a temperature of 63 ° C on the surface of the reactor 20. Under these conditions, the gas mixture partially condensed on the surface of the reactor, which contributed to the removal of the heat of reaction. There were no clogging problems at work.

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