WO2005030811A1

WO2005030811A1 - Process

Info

Publication number: WO2005030811A1
Application number: PCT/GB2004/003872
Authority: WO
Inventors: Erling Rytter; Arne GRISLINGÅS
Original assignee: Statoil Asa; Smaggasgale, Gillian, Helen
Priority date: 2003-09-26
Filing date: 2004-09-09
Publication date: 2005-04-07
Also published as: GB0322648D0

Abstract

A process for olefin polymerization comprising the steps of (a) providing a feed comprising olefin and the corresponding alkane in an amount of from about 1 mol% to about 50 mol% to a polymerization reaction zone, (b) operating the reaction zone under reaction conditions such that at least some of the olefin is polymerised to the desired polyolefin, and (c) removing at least some of the unreacted olefin and corresponding alkane from the reaction zone, purging at least some of the alkane and recycling any remaining alkane and the unreacted olefin to the reaction zone.

Description

PROCESS

The present invention relates to a process for the polymerization of olefms in the presence of an alkane. More particularly, the present invention relates to an improved process for the polymerisation of propylene from a feed comprising propane and propylene.

Conventionally olefins, particularly propylene, for use in the production of polyolefins are produced as co-product from steam crackers or are extracted as a byproduct from catalytic crackers at oil refineries. However, as the demand for olefins, in particular propylene, has increased, new means for producing the starting materials from which the olefin can be obtained have been considered. In the case of propylene, such means include dehydrogenation of propane, olefin conversion technologies and the conversion of methanol to propylene. This latter method is particularly useful as the starting product from which the methanol is produced may be natural gas (methane).

The use of natural gas as a starting material means that the production of propylene is no longer dependent on the supply and price of crude oil, condensates, naphtha and natural gas liquids. Further, since natural gas is available in many parts of the world, often with the volume of supply being substantially greater than local demand, it is a cost-effective starting material. The methane may be converted to methanol at the site of extraction. The methanol can then readily be transported as a liquid to the site at which further reactions are to be carried out or may be directly converted to propylene at the site of methanol production. Processes for the production of propylene from methanol are well known in the art.

Generally, reactions for the production of polyolefins such as polypropylene require very pure olefin, which is often referred to as "polymer grade olefin". Polymer grade polypropylene comprises at least approximately 99.5 mol% propylene with the remainder comprising essentially propane. To produce this grade of olefin trace amounts of mercaptans, dienes, and other reactive compounds which are usually present in the olefin must be removed.

In addition, the presence of inerts such as the alkane corresponding to the olefin and other hydrocarbons should be rninimised. Although these inerts do not interfere with the polymerisation reaction, it is necessary to remove them as any recycle of unreacted olefin to the reactor will carry with it the inerts which will result in buildup of the inerts in the reactor. If the reactor is to operate effectively, this build-up of inerts will have to be removed by means of a purge stream. However, the use of the purge stream will mean that desirable feedstock, i.e. olefin, is also purged from the reactor thereby reducing the efficiency of the reaction.

In conventional processes, a splitter is conventionally included prior to the polyolefin reactor to separate out the inerts such as alkanes. A conventional process for the production of, for example, polypropylene can be described with reference to Figure 1 which illustrates a flowsheet for a process from a natural gas starting material to the polypropylene product. It will be understood that the figure is schematic only and that in use, an integrated process may be used or the "lines" may be replaced by transport means.

In Figure 1, natural gas is fed in line 1 to a gas separation unit 2 and a light fraction is passed in line 3 to the synthesis gas production unit 4. Synthesis gas (known as syngas) which, comprises a mixture of carbon monoxide and hydrogen with some carbon dioxide, unconverted methane and steam, is passed in line 5 to a methanol production plant. It will be understood that the syngas production unit may be integrated with, the subsequent methanol production facility.

The methanol, or the dehydrated product dimethyl ether, is passed in line 7 to the methanol to propylene plant where it is converted to a mixture of several products which will include a C₃-fraction comprising propylene and some propane. As the polypropylene plant 14 requires polymer grade polypropylene, which normally comprises at least 99.5 mol% propylene the rest essentially being propane, the C₃- fraction is passed in line 9 to a C₃-splitter 10 where the majority of the propane is separated from the polypropylene. Separated propane, from the bottom of the C₃- splitter 10 may be used for energy purposes such as for fuel gas.

If the feed to the C₃-splitter comes from an alternative source of propane/propylene mixture such as from a dehydrogenation unit, the propane may be recycled to the dehydrogenation unit.

Under normal operations, the propane stream will contain small amounts of unseparated propylene, typically in the range of from 0.5 to 3 mol%. Even these small amounts represent a loss of feed to the polypropylene plant and hence a reduction in the cost-efficiency.

The polymer grade propylene is removed from the top of the C₃-splitter 10 and passed in line 12 to the polypropylene plant 14 where the propylene is converted to polypropylene which is removed in line 15 and unconverted propylene is removed and recycled in line 17 to the polypropylene plant 14. Depending on the polypropylene technology utilised, the recycle may constitute a small or a major part of the total feed to the polypropylene reactor. A purge 16 is required to avoid buildup in the polypropylene reactor system of inert components. This build-up can be significant and may result in a propane content in the reactor of from about 10 mol% and about 30 mol%.

Since the separation of, for example, propane and propylene is difficult because of low boiling points at atmospheric pressure (-47.0°C and -42.1°C respectively), the separation in the C₃-splitter 10 must be carried out at high pressure and with high reflux and thus the C₃-splitter will generally be the largest separation tower on the production site and will require substantial amounts of energy to run. Therefore the presence of the splitter will add significantly to the costs of the process both in terms of capital and running costs. A C₃-splitter for use in the prior art process is typically operated at about 10 to about 15 bar, for e ample at 12 bar, to give 0.5% propane in the top fraction at 43°C and 1% propylene in the bottom fraction. Depending on the feed composition, this separation will require a large number of trays in the column ranging from in the region of 50 to 100 if the feed contains 97 to 98% propylene, to more than 200 if the feed contains about equal amounts of the two components. In addition, the reflux ratio, i.e. the ratio between the reflux stream and the feed stream to the column, typically is around 9-10. This can be compared with typical reflux ratios for most distillation columns being around 2-3. Since a large reflux ratio requires a large sized distillation apparatus and an increased energy requirement for the boiler of the column, the cost of producing the required purity of propylene is high, which results in a corresponding high cost for the production of polypropylene.

Alternative processes for the production of polyolefins have been suggested. One example of this, is the Borstar™ technology which combines a loop reactor in which supercritical propane is used as the polymerisation medium and a gas-phase reactor in which the second stage of the reaction takes place.

In the Borstar process, the amount of propane in the loop reactor may be in the range of from about 10 to about 35 mol%. Depending on the recycle ratio around the polypropylene reactor, the inert, i.e. propane, content, under actual operation may vary substantially from process to process, and even with polymerization grade propylene as a feed may reach above 30% for a typical per path conversion of 50%. It will be understood that the presence of a high level of inerts in the reactor reduces the rate of reaction for the production of polypropylene. As it is usually undesirable to increase the reaction temperature to compensate for the reduction in rate of reaction since it will usually adversely affect the selectivity of the reaction to the desired product, and as it is desirable to maintain the rate of reaction, the pressure may be increased. Additionally, or alternatively, the rate may be increased by increasing the catalyst loading in the reactor. It will be acknowledged that in any reaction of olefins it is desirable to minimise the capital investment, maximise the energy efficiency of the process and where possible take advantage of economies of scale and therefore it is always desirable to adjust process schemes to minimise the cost of the reaction.

One suggested means for reducing the energy consumption in a reaction to produce polypropylene from propylene is described in EP 0 887 359. Here, a process is suggested in which a propylene feed comprising from 0.1% to 20% by weight propane is fed directly to a polymerization reactor, at least a part of the unreacted propylene discharged from the reactor is distilled to remove propane contained in the unreacted propylene and the purified propylene is then recycled to the reactor. The distillation column used for the post-reactor separation is said to be of a smaller size than that used in conventional pre-reactor separations. No purge is included in the system. A similar arrangement is described in EP 0 729 982 and again no purge is utilised. These proposals have not found favour commercially.

An alternative proposal for addressing the problems associated with conventional processes is described in US 5470925. In the process described, the gas mixture containing propylene and propane is separated from the polymer product and is then subjected to adsorption in a bed of absorbent which selectively adsorbs the alkene to adsorb the propylene. The propane is discarded and the propylene is desorbed from the adsorbent and recycled to the polymerization zone.

Whilst these processes go some way to addressing the problems of the conventional processes, there is still a need for alternative processes which will allow for the cost- efficient preparation of a polyolefin, particularly polypropylene, from an olefin feed containing at least some of the alkane corresponding to the olefin and which are capable of being utilised commercially.

Thus according to the present invention there is provided a process for olefin polymerization comprising the steps of: (a) providing a feed comprising olefin and the corresponding alkane in an amount of from about 1 mol% to about 50 mol% to a polymerization reaction zone; (b) operating the reaction zone under reaction conditions such that at least some of the olefin is polymerised to the desired polyolefin; and (c) removing at least some of the unreacted olefin and corresponding alkane from the reaction zone, purging at least some of the alkane and recycling any remaining alkane and the unreacted olefin to the reaction zone.

Thus, surprisingly, the process of the present invention provides a process for the production of a polyolefin such as polypropylene from a feedstock which is rich in olefin but which does not have the required polymer grade specification of 99.5% olefin content but in which the costly large splitter of the prior art is eliminated. Whilst it is acknowledged that the process may result in the loss of some olefin from the feed with the purged alkane, this possible small increase in the amount of lost olefin is outweighed by the savings in investment costs.

The feed comprising an olefin and the corresponding alkane may be provided from any suitable source. In one preferred arrangement the feed is prepared by an oxygenate to olefin process. The olefin is preferably propylene and in this case, the corresponding alkane will be propane. In a preferred arrangement, the propylene is prepared from methanol. In a most preferred arrangement, the methanol to propylene plant is integrated with the plant for the production of polypropylene in accordance with the process of the present invention. In this anangement, the propylene feed may have a propane content of 2 to 3 mol%.

Thus in its simplest arrangement the present invention relates to a process in which the separation of the alkane from the olefin is not required prior to the polymerization reaction and no separation of the propane other than via a purge stream is required. Whilst this may result in the loss of some propylene from the system, economically this is not significant when compared with the costs of constructing, commissioning and operating the large splitter located prior to the polymerisation reactor of the prior art system.

In one arrangement of the present invention, means for partially separating the alkane from the olefin may be included prior to the polymerisation reaction zone. Any suitable separation or adsorption means may be used. For example, a membrane separation means may be utilised or a splitter of smaller dimensions than that conventionally used may be suitable to obtain the required 95 to 99.5% purity of the olefin stream. For example, if the feed contains 97 to 98% that is upconcentrated to 99.5% olefin, in the region of about 60 to about 80 trays may be used for a feed having approximately equal amounts of olefin and alkane that is upconcentrated to 95% olefin, a column having from about 80 to about 120 preferably about 100 trays may be used. This arrangement may be used alone or may be combined with the use of a separation means in the recycle loop. Where the optional splitter is present an additional purge may be taken from the bottom of the splitter. In a system in which the alcohol to olefin plant is combined with the polyolefin plant, the alkane purged from process may be recycled to the olefin production plant.

The process of the present invention may be combined with a process for olefin production from alcohol.

In one alternative arrangement, a means for separating the alkane from unreacted olefin may be included in the recycle loop. In this arrangement, a stream comprising product, unreacted olefin and alkane will be removed from the polymerisation reactor, the product will be recovered and a the unreacted olefin and alkane will be passed via a separation means and recycled to the polymerisation reaction in a recycle stream. A purge stream will be removed.

In an alternative arrangement, a portion of the unreacted olefin and alkane will be recycled to the polymerisation reactor in a first recycle stream. The remainder of the unreacted olefin and alkane will be passed to a separation means. Some of the unreacted olefin and alkane will form a second recycle stream which may be recycled directly to the polymerisation reactor or may be combined with the first recycle stream such that the two are added to the reactor together. Some alkane will be recovered from the separation means and a purge gas will also be removed.

Whilst the processes of the present invention are particularly suitable for use in a homopolymerisation reaction, it may also be applied to copolymerization reactions. For example, propylene may be copolymerised with α-olefms having 2 to 18 carbon atoms, cyclo-olefins having 3 to 18 carbon atoms; vinylidene aromatic monomers such as styrene, vinyl monomers, conjugated dienes, non-conjugated polyenes, acetylenes and aldehydes.

The polymerization may be carried out as a liquid phase reaction or as a gas phase reaction. Any suitable polymerisation catalyst may be used. Any suitable polymerisation process may be used and the polymerisation conditions will vary depending on the desired properties of the polymer, whether there is any co- polymerisable monomer present and the type of polymerisation. However, the polymerization pressure will generally be from atmospheric to about 100 kg/cm² and the temperature may be in the region of 20°C to 200°C. Although the present discussion has been made with reference to a single reactor and reaction, it will be understood that the polymerization may be earned out in two or more stages.

A hydrocarbon solvent which is inert to the polymerization reaction may be present. Any suitable solvent may be used. One suitable solvent is hexane. Hydrogen may be added to the polymerisation reaction as a molecular weight modifier.

In a particularly preferred arrangement, the procedure for polymerisation employs the Borstar™ technology. The present invention will now be described by way of example, in connection with the production of polypropylene from a feed comprising propylene and propane with reference to the accompanying drawings, in which

Figure 1 illustrates a combined process for producing polypropylene from natural gas utilising conventional technology to reduce the propane content of the feed;

Figure 2 illustrates a combined process for producing polypropylene from natural gas starting material in accordance with the present invention;

Figure 3 is a graph illustrating the yield of polypropylene for different propylene/propane feed ratios;

Figure 4 illustrates one alternative arrangement for the polypropylene plant;

Figure 5 illustrates an alternative arrangement for the polypropylene plant;

Figure 6 illustrates a still further alternative arrangement for the polypropylene plant; and

Figure 7 illustrates a modified arrangement to that set out in Figure 4.

It will be understood by those skilled in the art that the drawings are diagrammatic only and that further items of equipment such as reflux drums, pumps, vacuum pumps, temperature senors, pressure sensors, pressure relief valves, control valves, flow controllers, level controllers, holding tanks, storage tanks and the like which may be required in a commercial plant. The provision of such ancillary items of equipment forms no part of the present invention and is in accordance with conventional chemical engineering practice.

Whilst the process of the present invention will be described commencing from natural gas starting material, it will be understood that the propylene feed to the polypropylene plant may be provided from any suitable source. Further, it will be understood that where the feed is produced from methanol, the methanol to propylene plant may be combined with the polypropylene plant.

As illustrated in Figure 2, natural gas is fed in line 18 to a gas separation unit 19 and a light fraction is passed in line 20 to a synthesis gas (syngas) production unit 21. The syngas, which comprises a mixture of carbon monoxide, and hydrogen with some carbon dioxide, unconverted methane and steam is passed in line 22 to a methanol production plant 23. The syngas production unit may be integrated with the methanol production facility.

The methanol, or the dehydrated product dimethyl ether, is passed in line 24 to a methanol to propylene plant 25 where it is converted to a mixture of several products which will include a C₃ -fraction comprising propylene and some propane. This C₃- fraction is passed in line 26 to the polypropylene plant 27 where the propylene is converted to polypropylene which is removed in line 28. Unreacted feed is recycled to the reactor 27 in line 29 with a portion of the unreacted feed being removed as a purge stream 30. Thus the C₃-splitter of the prior art has been removed. The purge, including the propane, may be separated by existing separation means if available in the petrochemical complex, exported as a low grade propylene or it may be used for other purposes such as for fuel gas in the syngas production or it may be transferred to a cracker unit. It will be understood that the flow sheet can be modified in any manner. Further, units of product separation, clean-up and by-product conversion units may be present. The polypropylene reactor can be of any suitable format and may utilise the Borstar technology comprising a slurry loop-reactor followed by one or more gas phase fluidized bed reactors. If in the process of Figure 2, the polypropylene reactor is completely back-mixed of the CSTR type with a constant ratio between propylene and propane of 1 to 6, with conditions of catalyst activity, heat removal capacity and reactor volume to produce the given yield of polypropylene. For these conditions, with a given feed of 100 mol propylene and 3 mol propane form the methanol to propane unit, the amount of propylene lost in the purge will be 0.5 mol% and the yield of polypropylene will be 99.5%. More generally if the feed is composed of 100 mol propylene and X mol propane and the propylene .-propane in the reactor R, then the loss of propylene through the purge is X*R mole and the yield of polypropylene is 100 - X*R mol.

The yield of polypropylene is illustrated in Figure 3. It can be seen that if the propane concentration in the propylene feed (X) is moderate, then fairly large concentrations of the monomer in the reactor are acceptable. Thus, irrespective of the propylene/propane ratios in the polypropylene reactor, where the make-up gas comprises 2 to 4 mol propane for 100 mol propylene, the reduction in yield is minimal.

An alternate arrangement is illustrated in Figure 4. Here a simplified C₃-splitter 31 is located before the polypropylene plant. Propane removed from the bottom of the splitter in line 32 will generally be recycled to earlier in the system. The splitter may be fed by the methanol to propylene plant or by purge from the polypropylene reactor. If the splitter is a distillation column, it will have a reduced number of trays than that required in the conventional system. There will typically be about 80 trays if the feed is concentrated from 97 wt % to 99.5 wt % such as, for example, concentration of a methanol to propeylene type feed to petrochemical grade. Alternatively, the feed can be a 50/50 mixed feed that as upgraded to 97 wt % methanol to propylene quality, or a reduction in column size of from about 200 to about 125 trays. A further alternative arrangement is illustrated in Figure 5 in which a simplified C₃ splitter 32 is located in the recycle loop. Here propane is removed from the bottom of the splitter in line 34 will generally be recycled to earlier in the system. A purge will be removed from the steam feeding the C₃ splitter.

A still further alternative anangement is illustrated in Figure 6 in which only a portion of the recycle stream is passed through a simplified C₃ splitter 35. The product stream from the polymerisation reactor 27 is separated and the polypropylene is recovered in line 28. A portion of the remainder is recycled in line 36 to the reactor and the rest is passed to the splitter 35. Propane is removed as bottoms in line 37 and the lights are recycled to the polymerisation reactor in line 38. A purge is removed in line 39.

Figure 7 illustrates a modification of the anangement illustrated in Figure 4. A portion of the product containing stream is recycled to the polypropylene plant in line 29 and a portion to the simplified C₃ splitter in line 33.

The present invention will be described further with reference to the following examples.

Example 1 (with particular reference to Figure 2).

A full scale methanol to propylene plant of 500 000 tonnes/year C₃ that contains only 3% propane, translates to about 60.6 tonnes/hour propylene and 1.9 tonnes/hour propane. Assuming that the polypropylene unit can only tolerate 10 wt% propane in the reactor in order to reduce catalyst costs. In this anangement the purge stream according to Figure 2 will contain 1.9 tonnes/hour of propane and 17.1 tonnes/hour propylene. This means that 500 000 tonnes/year C₃ produced by the methanol to propylene plant will yield 350 000 tonnes/hour polypropylene and 150 000 tonnes/year refinery grade propylene having a propylene content of 90 wt%.

Example 2 ( with particular reference to Figure 6). If the propylene is upgraded to the original 97 wt% methonal to propylene grade of propylene this requires a column having only a feed stream of 18.8 tonnes/hour compared to the full 61.5 tonnes/hour methanol to propylene C₃ product stream such that a much slimmer column may be used. It is noted that it is easier to concentrate from 90 to 97 wt% than from 97 to 99.5 wt%. A column height approximately 30 wt% shorter than that required in the Figure 1 configuration can be achieved which can equate to a reduction of from 80 to 54 trays. This vastly simplified C₃ splitter is able to boost the polypropylene production to 450 000 tonnes/year polypropylene or further if appropriate recycle of the purge system is applied.

Example 3 (with particular reference to Figure 7).

100 tonnes/hour of propylene is fed to the C₃ splitter. The propylene content in this feed varies widely. It may be the final grade propylene with a content of about 65 to 80 wt% propylene. It may be the product from the dehydrogenation unit containing 30 to 50 wt% propylene or it may be a mixture of these two feed streams. In one alternative it may be the C₃ mixture from a naphta steam cracker. It will also be noted that a minor amount can be recycle from the polypropylene unit. If for simplicity we will assume a 1:1 ratio between propylene and propane, the column produces over the top of the column typically 49.75 tonnes/hour propylene and 0.25 tonnes/hour propane. The bottom stream typically consists of 49.75 tonnes/hour propane and 0.5 tonnes/hour propylene. The latter stream can be recycled to the olefin unit which may be dehydrogenation or cracker or used internally as fuel in the petrochemical complex or exported.

If this petrochemical complex is expanded with a full scale methanol to propylene count of 500000 tonnes/year that contains only 3% propane. This translates to about 60.6 tonnes/hour propylene and 1.9 tonnes/hour propane.

In contrast if one assumes a polypropylene plant with a 2:1 ratio between propylene and propane the total stream that needs to be recycled from the polypropylene unit to the C₃ splitter is only 5.7 tonnes/hour which equates to a 5.7% increase in the original feed. This slight increase is often encountered by design overcapacity of the column, optimization of the column or simply by reducing the original feed to the column whereby a slightly more than than 0.5 wt% propane in the product. Thus one has eliminated an additional C₃ splitter for the expansion even though the expansion has more than doubled the polypropylene capacity.

The splitter of Figures 4, 5, 6 or 7 may be replaced by a membrane separator or an adsorber.

Claims

1. A process for olefin polymerization comprising the steps of: (a) providing a feed comprising olefin and the conesponding alkane in an amount of. from about 1 mol% to about 50 mol% to a polymerization reaction zone; (b) operating the reaction zone under reaction conditions such that at least some of the olefin is polymerised to the desired polyolefin; and (c) removing at least some of the unreacted olefin and conesponding alkane from the reaction zone, purging at least some of the alkane and recycling any remaining alkane and the unreacted olefin to the reaction zone.

2. A process according to Claim 1 wherein the feed is prepared by an oxygenate to olefin process.

3. A process according to Claim 1 or Claim 2 wherein the olefin is propylene and the conesponding alkane is propane.

4. A process according to Claim 3 wherein the propylene is prepared from methanol.

5. A process according to Claim 4 wherein the methanol to propylene plant is integrated with the plant for the production of polypropylene in accordance with Claim 1.

6. A process according to any one of Claims 3 to 5 wherein the olefin feed has a propane content of 2 to 5 mol%.

7. A process according to any one of Claims 1 to 6 wherein means for partially separating the alkane from the olefin are included prior to the polymerisation reaction zone.

8. A process according to Claim 7 wherein the alkane is separated from the olefin prior to the reaction zone to give a feed with less than 99 mol% propylene.

9. A process according to Claim 7 wherein the alkane is separated from the olefin prior to the reaction zone to give a feed with 95 to 98 mol% propylene.

10. A process according to any one of Claims 1 to 9 wherein a separation means is included in the recycle loop.

11. A process according to Claim 10 wherein the alkane is partially separated from the olefin of the purge or recycle stream prior to being recycled to the reaction zone.

12. A process according to Claim 10 wherein a portion of the recycle loop stream is passed through the separation means.

13. A process according to any one of Claims 7 to 12 wherein the means for partially separating the alkane from the olefin is selected from adsorption means, membrane separation means or a simplified splitter.

14. A process according to Claim 13 wherein the means for partially separating the alkane is a simplified splitter having less than 80 trays.

15. A process according to Claim 13 wherein the means for partially separating the alkane is a simplified splitter having less than 65 trays.

16. A process according to any one of Claims 1 to 15 wherein the reactor of step (b) employs the Borstar™ technology.

17. A process according to any one of Claims 1 to 16 wherein the purge gases fed to a conventional splitter of a separate petrochemical plane.