Apparatus for preparing polyolefin products and methodology for using the same

Apparatus for preparing polyolefin products and methodology for their use

apparatus for preparing polyolefin products and methodology for using the same

Full description

Background of the invention

This application is considered part of the application filed with the Saudi office No. 04250101 dated 03/14/1425 AH.

The present invention relates to a liquid phase olefin polymerization process, to the preparation of polyolefin products, and to an apparatus useful in the preparation of polyolefin products. In particular, the present invention relates to apparatuses and equipment used to prepare a variety of polyolefin products using a liquid phase polymerization process and to the methodology used in the operation of such devices and equipment. Specifically, the present invention relates to a device and methodology that enhances the operation and control of polyolefin reactors.

The currently pending U.S. patent application, serial number 515790/09 filed on February 29, 2000, and entitled “Process for Producing High Vinylidene Polyisobutylene” (hereinafter referred to as Application No. 790) relates to liquid-phase polymerization processes used to prepare Low molecular weight, highly reactive polyisobutylene. As disclosed in Application No. 790, a catalyst composition is introduced, which may desirablely comprise a composition of BF3 and methanol,

The feedstock containing isobutylene is transported to the reaction zone, where they are intimately mixed with a residual reaction mixture to obtain an intimately intermixed reaction admixture in the reaction zone. The homogeneous reaction mixture is maintained in its state of homogeneous mixing and at a relatively constant temperature not less than about 0°C while it is in the reaction zone. Thus, the isobutylene present in the reaction mixture is polymerized to form polyisobutylene (PIB) with a degree of High terminal unsaturation. A crude product stream containing residual catalyst composition, unreacted isobutylene and polyisobutylene is then withdrawn from the reaction zone. Both the introduction of the feedstock and the withdrawal of the product stream from the reaction zone are controlled so that the residence time of the isobutylene undergoing the polymerization process in the reaction zone does not exceed about 4 minutes, and thus the product stream contains a highly reactive polyisobutylene product. It is preferable for the reaction area to be the tube side of the heat exchanger consisting of a shell-and-tube exchanger, where the coolant is circulated in the shell side. A recirculation loop may preferably be used to circulate the reaction mixture through the reaction zone represented by the tube side at a linear velocity sufficient to stabilize and maintain the homogeneous state in the mixture and to remove the heat generated by the exothermic polymerization reaction.

U.S. Patent No. 6,525,149 issued on February 25, 2003, entitled “Process for preparing polyolefin products” (hereinafter referred to as Patent No. 149), relates to a new liquid-phase polymerization process used to prepare a polyolefin product that has the properties Pre-selected. The process, according to US Patent No. 149, includes the steps of supplying a liquid feedstock containing an olefinic component and a catalytic composition that may include a stable complex of BF3 and a complexing agent. The feedstock may comprise one or more olefins, including branched olefins such as isobutylene, linear alpha olefins containing 3 to 15 carbon atoms (C3-C15), and compounds other than C4-C15 alpha olefins. Reactive C4-C15 reactive non-alpha olefins. The feedstock and catalyst composition may be desirablely added to a residual reaction mixture that is recycled in a reaction zone of a cyclic reactor provided at the tubular side of a shell-and-tube heat exchanger at a recirculation rate sufficient to effect homogeneous mixing of the remaining reaction mixture, the added feedstock and the catalyst composition. The heat generated by the polymerization reaction is removed from the recycled homogeneous reaction mixture at a rate calculated to provide it with an essentially constant temperature during its recycling in the reaction zone. Conditions in the reactor are suitable to cause olefinic components introduced into the feedstock to undergo a polymerization process to form the desired polyolefin product in the presence of the catalyst composition. A crude product stream containing the desired polyolefin product, unreacted olefins, and remaining catalyst composition is withdrawn from the reaction zone. The introduction of the feedstock into the reaction zone and the withdrawal of the product stream from the reaction zone are controlled so that the residence time of the olefinic components undergoing the polymerization process in the reaction zone is appropriate to produce the desired polyolefin product.

US Patent Publication No. A1 0040587-2003 published on February 27, 2003 and entitled “Mid-Range Vinylidene Content Polyisobutylene Polymer Product And Process For Producing The Same” describes Bulletin No. 587 follows a polymeric product of PIB with medium-range vinylidene content and a process for making it. As disclosed in Bulletin No. 587, at least approximately 90% of the PIB molecules in the product contain isomers in the alpha position or beta position. The alpha isomer (vinylidene) may occupy a percentage of the product ranging from 20% to 70%, and the content of internal double bonds carrying four tetra-substitued internal double bonds is very low, preferably less than about 5% and typically less than It ranges from about 1-2%. Polymeric products of PIB with desirable intermediate-range vinylidene content are prepared by a liquid-phase polymerization process conducted in a loop reactor similar to the reactors described in Application No. 790 and Patent No. 587 at a temperature of approximately 15.56°C (centigrade degree). (60 F (Fahrenheit degree)) or higher using a catalytic compound consisting of BF3/methanol and a contact time of no more than about 4 minutes.

Application No. 790, Publication No. 587 and Patent No. 149 are each assigned to the assignee of the current application, and the contents of the relevant specifications are specifically mentioned in this statement for reference.

When performing the reactions described above, highly specialized equipment may often be used to enhance the operation and control of the polymerization reactors. In each case, for example, the crude product leaving the reactor may be contaminated with residual catalyst, which should preferably be rapidly quenched or quenched to avoid further polymerization of low molecular weight monomers and oligomers without appropriate quenching and/or isomerization resulting from displacement Location of the remaining double bond. The catalyst structure may be subject to contamination with residual materials recycled with the reaction mixture during the polymerization reaction. Moreover, as is the case in any industrial activity, there is a constant search for methodology and/or equipment used for the purpose of increasing capacity and throughput.

An important objective of the present invention is to satisfy the needs described above. In this regard, and in one of the important aspects of the invention, the invention provides a device for the polymerization of olefins that includes several reactors. In accordance with the concepts and principles of the invention, each of these reactors may desirablely have a structure defining a reaction zone, an olefin polymerization reaction mixture inlet connection and an olefin polymerization reaction mixture outlet connection. It is desirable that these connections be in fluid communication with the reaction zone. All reactors are prepared and arranged to facilitate an exothermic olefins polymerization reaction in the reaction zone.

In accordance further with the concepts and principles of the invention, each of the reactors may further comprise a recirculation system including a pump arranged and configured to circulate the reaction mixture in the reaction zone independently of the introduction of an olefin-containing feedstock into the reactor.

The device according to the invention also desirablely includes an olefin containing feedstock distribution assembly containing an inlet for an olefin-containing feedstock and several outlets for an olefin-containing feedstock. The arrangement of the distribution structure is such that each of the feedstock outlets is in fluid contact with the reaction zone of a given reactor. The device, according to the invention, may also include, as desired, a product collection assembly including several inlets for raw polyolefin products and an outlet for raw polyolefin products, and the arrangement of the assembly assembly is such that each inlet for raw polyolefin products is in contact. Fluid with the reaction zone of a given reactor.

In summary, the device according to the invention may include two or more reactors, for example three, four, five, six or more reactors.

In another important aspect of the invention, it provides a method for polymerizing olefins. According to the invention, the method includes powering several reactors, each of which defines an internal reaction zone. The method also includes supplying a feedstock containing olefin, dividing this feedstock into several separate feedstock streams (2, 3, 4, 5, 6 or more), and adding each feedstock stream to a circulating reaction mixture. mixture in the reaction zone of one of the reactors in question, and performing an exothermic olefins polymerization reaction in each reaction zone.

The method according to this aspect of the invention also includes the steps of circulating the reaction mixture separately in each reactor independently of the addition of the feedstock stream in question to the reaction mixture, removing the feedstock polyolefin product stream in question from the reaction mixture circulating in each of the reactors, and mixing the feedstock streams in question. Crude polyolefin products to form a single crude product stream.

In another aspect, the invention provides a reaction device used for the polymerization of olefins comprising at least one reactor defining a reaction zone and including an inlet connection for the olefin polymerization reaction mixture and an exit connection for the olefins polymerization reaction mixture. It is desirable that these connections be in fluid contact with the reaction zone. The reactor is prepared and arranged to facilitate an exothermic olefins polymerization reaction on the reaction mixture in the reaction zone in the presence of a catalytic composition comprising a catalyst and a catalyst modifier. According to this aspect of the invention, the reaction apparatus further includes a feedstock inlet, a feedstock outlet, and a recirculation system containing a pump arranged and prepared for the purpose of circulating the reaction mixture in the area independently of the feedstock being introduced into the reaction mixture through said feedstock inlet. The reaction apparatus according to this aspect also includes an inlet for a catalyst composition that is in fluid contact with the region, which facilitates the addition of the catalyst composition to the olefin polymerization reaction mixture, and an inlet for at least one catalytic modifier that is in fluid contact with the region, which facilitates the addition of the catalytic modifier to the reaction mixture. Olefin polymerizes at a rate that does not depend on the rate of addition of the catalytic composition.

Another important feature of the invention is to provide a method used to operate an olefins polymerization reactor. This method includes the steps of supplying a reactor for the polymerization of olefins that contains a reaction zone, recycling the olefins polymerization reaction mixture in the zone, adding feedstock containing olefin to the aforementioned reaction mixture, where the aforementioned polymerization reaction mixture is recycled at a flow rate of no. Depends on the rate at which feedstock is added to the recycled reaction mixture for olefin polymerization, adding a catalytic composition including a catalyst and a catalytic modifier to the recycled reaction mixture for olefin polymerization, subjecting the polymerization reaction mixture to conditions An exothermic zone olefin polymerization reaction in the presence of the catalyst composition, and the addition of a catalytic modifier to the recycled reaction mixture to polymerize olefins at a rate that does not depend on the rate of addition of the catalyst composition.

According to the concepts and principles of the invention, both the aforementioned system and the aforementioned methodology may be used in conjunction with a system and/or methodology that includes and/or includes a single reactor vessel only or that includes and/or includes several reactor vessels arranged in parallel as Description above. In this regard, it should be noted that according to the invention, the invention further provides an apparatus and/or method comprising and/or including a multi-reactor system as described above in combination with the described system for adding the catalyst modifier to the recycled reaction mixture at a rate independent of Catalyst composition addition rate.

An additional aspect of the invention is to provide a catalytic removal and washing system for a raw polyolefin product. According to this aspect of the invention, the catalyst removal and washing system includes an upstream settler vessel defining an internal settlement chamber configured and arranged to receive a mixture of raw polyolefin product and aqueous wash media and allow the product and medium to be separated under the influence of gravity. . The system also includes a line for entering a raw product for olefin polymerization that contains a catalyst that is in fluid contact with the compartment of the previous settling vessel, a channel for entering the inlet conduit catalyst killing agent that is in fluid contact with the compartment of the previous settling vessel, and a first passage for entering The first make-up water inlet passagway is in fluid contact with the compartment of the previous settling vessel.

In addition to the above, it is desirable for the catalyst removal and washing system to include a downstream settler system containing at least one aftersettling vessel defining an internal settling chamber prepared and arranged to receive a mixture of partially washed crude polyolefin product. polyolefin product and water washing medium and allowing the product and medium to be separated under the influence of gravity, a partially washed polyolefin product line, an overhead line that connects the pre-sedimentation vessel compartment with the post-sedimentation system, line The washed crude olefin polymerization product outlet line is in fluid contact with the post-precipitation system, and a second passage is for the entry of make-up water, which is in fluid contact with the post-precipitation system. Finally, the system includes a first drain line that connects the compartment of the previous settling vessel with an inlet connection to the waste water receiving system, and a second drain line that connects the subsequent settling system with the inlet connection of the wastewater receiving system. Therefore, the used wash water from the preceding and subsequent parts of the system can be drained separately from the system.

According to the aspect of the foregoing invention, the subsequent sedimentation system may include one, two, three, or more separate sedimentation vessels.

Another important feature of the invention is to provide a method for treating a catalytically formed crude polyolefin product containing residual catalyst to avoid further reaction in the product and to remove residual catalyst from it. According to this aspect of the invention, the method includes homogeneously mixing a raw polyolefin product containing the residual catalyst with a first aqueous media containing a catalytic quenching agent so as to form a first intimately admixed two phase mixture capable of being separated by gravity. , gravity separable mixture, introducing the first two-phase mixture into the first settlement zone and allowing it to settle in the first zone under the influence of gravity in the presence of an upper phase of a partially washed polyolefin raw product and a first lower aqueous phase containing Dissolved catalyst salts, withdrawing the first lower aqueous phase from the first sedimentation area, recycling a first part of it and adding it to the first two-phase mixture to form part of the first aqueous medium, directing a second part of the lower aqueous phase The first to the drain for disposal or extraction, adding a first amount of make-up water to the first two-phase mixture to form part of the first aqueous medium, withdrawing the partially washed raw polyolefin product phase from the first deposition area and mixing it thoroughly with a second aqueous medium to form a binary mixture. The second homogeneous phase can be separated by gravity. Add the second biphasic mixture to a second sedimentation zone and allow it to settle in the second zone under the influence of gravity to obtain an upper phase of the washed raw product. Completely made of polyolefin fully washed crude polyolefin product phase and a second lower aqueous phase, withdrawing the second lower aqueous phase from the second sedimentation area and recycling a first part of it and adding it to the second two-phase mixture to form part of the second aqueous medium, directing a second part from the second lower aqueous phase To a drain for disposal or extraction, remove the completely washed raw polyolefin product phase from the second settling zone, and add a separate quantity of make-up water to the second two-phase mixture To form part of the second aqueous medium.

It is worth noting that the catalyst removal and washing system and/or method described and/or described above is and/or is suitable for use in conjunction with any system containing only a single reactor as well as a system containing multiple reactors as described above. Accordingly, an important aspect of the invention is to provide a system and/or method comprising and/or including a multi-reactor system, a catalytic removal and washing system and/or the method as described. In addition, such combined system may also include the described system for adjusting the amount of catalyst rate in the recycled reaction mixture.

Figure 1 represents a schematic illustration of a reactor system including a multi-pass shell and tube heat exchanger and a recirculation system that is useful in connection with the invention;

Figure 2: A process flow diagram illustrating an apparatus embodying the concepts and principles of the invention and using two reactors of the type shown in Figure 1 and arranged for parallel operation; And

Figure 3: A process flow diagram showing a system for receiving a polyolefin crude product from the devices shown in Figures 1 and 2, for example, and processing it to wash the crude product and remove the remaining catalyst.

Several reactors of potential value and use when performing liquid-phase polymerization of polyolefins are known to those skilled in the technology to which the invention relates. However, for the purposes of a preferred embodiment of the present invention, the reactor may desirablely comprise a two-pass shell-and-tube heat exchanger as shown in Figure 1, where designated 10. Reactor 10, for example, may comprise three hundred and eighty-eight (388) tubes each with a diameter of 9.525 millimeter (0.375 inch) and a wall thickness of 0.889 mm (0.035 inch), thus each providing an inner tube diameter of 7.75 mm (0.305 in). The reactor can be 3.7 meters (12 feet) long and can contain internal baffling and internal partitions to provide two passages with 194 tubes per passage. The lanes are designated as 50 and 51 in Figure 1, and the 194 tubes per lane are represented by individual tube segments 52 and 53, respectively. Such a composition is well known in heat exchanger and reactor technology and it is not believed that any further clarification would be necessary.

Upon operation, a feedstock (e.g., isobutylene, 1-butene, 2-butene) containing olefin enters the reaction system through pump 14 and pipe 15. It is desirable that the downstream end be The pipe 15 is in a location that directs the feedstock to the suction line 20 of the recirculation pump 25. A catalyst composition can be injected into the circulation reaction system via pump 29 and pipe 30 at a location subsequent to the pump 25 and adjacent to the first pass, as shown in Figure 1. It is desirable that the catalyst composition be a mixture of methanol/BF3 with a molar ratio of methanol to BF3 of about 1:1.9 or less, and it is preferable that it be a mixture of methanol/BF3 with a molar ratio of methanol to BF3 of about 1. :1.7 or less. Desirably the molar ratio of methanol to BF3 may be as low as about 1:1.1 or less in some applications.

The circulation pump 25 pushes the reaction mixture through line 35, control valve 40 and line 45 towards the bottom head 11 of reactor 10. A flow meter 46 can be placed in line 45 as shown. Suitable temperature indicators (TI) and pressure indicators (PI) can be provided to monitor the system. The reaction mixture travels upward through tubes 52 of passage 50 and downward through tubes 53 of passage 51. The circulating reaction mixture leaves reactor 10 through suction line 20. Thus the reaction system is of a type sometimes referred to as a loop reactor. With this system, which is only preferable because there are many other arrangements that may be obvious to those familiar with the technique, the reaction mixture flow rate in the reactor can be adjusted and utilized independently of the feedstock introduction and product removal rates to achieve thorough intermixing of the catalyst and material composition. reactant and appropriate temperature control.

As previously explained, lanes 50 and 51 may each desirablely comprise one hundred and ninety-four (194) separate pipes. However, for the purpose of clarification, only part of a single pipe in each passage is shown schematically in Figure 11. These pipes are designated with reference numbers 52 and 53, respectively. Although only a portion of each of the represented tubes 52 and 53 has been shown, it must be realized by those skilled in the technique that each of these tubes extends across the entire distance between the top head 12 and the bottom head 11 and is in a state of communication. Fluid with the contents of heads 11 and 12.

It is worth noting in this statement that the reaction mixture should preferably be circulated through the tubes 52, 53 of the reactor at a flow rate sufficient to produce turbulent flow, in order to achieve homogeneous mixing of the catalyst composition and reactants and with an adequate heat transfer coefficient to provide adequate cooling. In this regard, the flow rate, reaction mixture properties, reaction conditions and reactor configuration should be appropriate to obtain a Re Reynolds number in the range from about 2000 to about 3000 and a heat transfer coefficient (U) in the range from about 24.41 to About 73.22 calories. min.cm2.m (50 to 150 Btu/min ft2 F) in pipes 52, 53 of the reactor. These media can generally be obtained when the linear flow rate of a typical reaction mixture through a tube with an internal diameter of 0.84 cm (0.331 in) is within the range of approximately 182.9 to 274.3 cm/s (6 to 9 ft/s). .

The outlet line 55 of the product may preferably be in fluid contact with the suction line 20 of the pump. However, as will be readily apparent to those versed in the technique, the exit line can be located almost anywhere in the system because the conditions in the reactor, at least theoretically, and as will be described below, can desirably approach those of a continuous stirred tank reactor (CSTR) continuous stirred tank reactor where both temperature and composition remain constant so that the composition of the product stream leaving the reactor is identical to the composition of the reaction mixture recycled in the reactor. Likewise, feedstock introduction line 15 can be located almost anywhere in the system, although, in practice, it is desirable to connect line 15 to the recycling system as far in advance as possible of line 55 to ensure that monomers entering through line 15 have a greater opportunity To polymerize before meeting line 55.

Desirably, a coolant can be circulated in the jacket side of the reactor at a rate that ensures the heat of reaction is removed and a desired temperature is maintained in the reactor.

A catalyst compound-forming agent may desirablely be added to the reaction mixture circulating through the pump 18 and line 16 located in the overhead header 12. This is a particularly valuable feature when the desired product is highly reactive polyisobutylene (HR PIB) and the catalyst composition includes BF3 catalyst and methanol composite forming agent. Compounds of BF3 with methanol come in two different forms, for example, mono-complex (composed of 1 mole of BF3 to 1 mole of methanol) and di-complex (composed of 1 mole of BF3 to 2 moles of methanol). ), depending on the availability of methanol. The monomeric complex is the true catalytic species, while the dimeric complex does not have any catalytic properties, especially in the absence of the monomeric complex. When referring to fractional complexes, they represent the actual average of the unilateral and binary complexes. In this regard, it has been shown that a catalyst composition of 0.59 to 0.62 moles of BF3 per mole of methanol is particularly valuable in the production of HR PIB. However, when this composition is introduced into the system, variations and contaminants in the hydrocarbon feed stream can often lead to suboptimal reactor tuning. This is thought to be, at least in part, a result of the tendency of many contaminants to effectively increase the apparent ratio of methanol to catalyst in the composition. Moreover, it is not always possible to determine the exact contamination level in advance for some feedstocks.

However, in accordance with the concepts and principles of the invention, it has been discovered that these problems can be solved and that optimal results can be achieved simply by introducing a catalyst composition, which for some purposes is desired to have a methanol lean composition, for example a catalyst composition containing a concentration exceeding the desired optimum concentration of the monomeric compound. , to reactor 10 via line 30, and independently add relatively pure methanol via a line that may want to be far from line 30, as in line 16. Desirably, a pump 18 may be provided to force the methanol through the pipe 16. Alternatively, essentially the same effect may be achieved by adding a separate stream of methanol directly to the catalyst assembly stream in line 30 through a line (not shown) and introducing both methanol The additive and catalyst are combined into the system. In either case, additional methanol is available to adjust the catalyst composition and so the desired ratio of methanol to BF3 can be achieved and maintained in reactor 10.

Further in accordance with the concepts and principles of the invention, the amount of methanol added must be sufficient enough to obtain and maintain a preferred ratio of BF3 per mole of methanol in the circulating reaction mixture. For some applications, for example where highly reactive polyisobutylene is the desired product, the composition of the catalyst added through line 30 may desirablely include a molar ratio of BF3 and methanol ranging from about 0.59:1 to about 0.62:1, and ideally may It is about 1:0.61. Alternatively, for other applications, for example where the desired product is polyisobutylene and the vinylidene content is not very important, the composition of the catalyst added through line 30 may ideally include a molar ratio of BF3 and methanol of about 1:1.

The product exiting the system via line 55 must be quickly quenched with a substance capable of quenching the activity of the catalyst, such as, for example, ammonium hydroxide, thus immediately stopping the ongoing exothermic polymerization reactions. Therefore, any unwanted rise in temperature caused by lack of cooling means (and the production of low-molecular-weight polymers associated with high temperatures) or rearrangement of polymer molecules can be reduced. The polyolefin products according to the invention can then be directed to a work up system, including a washing system as described below, where the catalyst salts can be removed and a purification and separation system (not shown in the drawing) where the A polyolefin product for monomers, dimers, unreacted oligomers and other unwanted contaminants such as diluents, etc. These last-mentioned materials can be recycled or converted for other uses using a well-known methodology.

In the recycling system described, the rate of feedstock addition to the reaction mixture and the rate of product removal do not depend on the recycling rate. As those familiar with the technology will realize, the number of passages through the reactor, its size, and its shape are simply matters of choice. For a single-reactor system as shown in Figure 1, the feedstock flow rates and product withdrawal rates can preferably be chosen such that the residence time of new monomers entering the reactor with the feedstock is 4 minutes or less, desirable 3 minutes or less, and preferably 2 minutes Or less, better yet one minute or less, ideally less than one minute. Regarding what was mentioned previously, the residence time is defined as the total reactor system size divided by the volumetric flow rate of the feedstock entering the system through pipe 15.

The recirculation flow rate, i.e. the flow rate of the reaction mixture in the system induced by the recirculation pump 25, is adjusted as described above, to achieve suitable turbulence and/or heat transfer characteristics. This recycling flow rate often depends on the system itself and other desired conditions of the process. For the systems described above, in general the ratio between the recirculation flow rate to the inlet feedstock flow rate (recycling ratio) should be maintained in the range from about 1:20 to about 1:50 and desirablely in the range from about 1:25. to about 1:40, and ideally in the range from about 1:28 to 1:35. In particular, in addition to creating turbulence and providing an adequate heat transfer coefficient, the recirculation flow rate of the reaction mixture must be sufficient to keep the concentrations of ingredients in it essentially constant and/or to reduce the temperature gradients in the circulating reaction mixture to a minimum. Thus, essentially isothermal conditions are established and maintained in the reactor.

As noted above, recycling ratios generally should be in the range from about 1:20 to about 1:50. Higher recycling ratios increase the degree of mixing and bring the reactor closer to isothermal operation, leading to narrower polymer distributions. But higher recycling rates also result in higher power consumption. Low recirculation ratios reduce the amount of mixing in the reactor and, as a result, significant discrepancy in the temperature profiles. When the recycling rate approaches zero, the design equations for the reactor reduce to those of a plug flow reactor model. On the other hand, when the recycling ratio approaches infinity, the modeling equations reduce to those of CSTR. When CSTR conditions are achieved, both temperature and composition values ​​remain constant and the composition of the product stream leaving the reactor matches the composition of the reaction mixture circulating in the reactor. Needless to say, after equilibrium is achieved, once the feedstock enters the system, an equal amount of product is forced out of the loop reactor. Thus, under CSTR conditions, the point at which the product stream is withdrawn does not depend on the reactor geometry.

The feedstock entering the system via line 15 may be any stream containing olefin. When polyisobutylene is the preferred product, the feedstock may be, for example, an isobutylene concentrate, a dehydro effluent stream, or a raff-1 stream. These feed raw materials are described respectively below in Tables 1, 2 and 3.

C3s

zero

I-butane

6.41

n-butane (n-butane refers to normal)

1.68

1-Butene 1-butene

1.3

I-butene

89.19

trans-2-butene

0.83

cis-2-butene

0.38

1,3-Butadiene

0.21

C3s

0.38

I-butane

43.07

n-butane

1.29

1-Butene 1-butene

0.81

I-butene

52.58

trans-2-butene

0.98

cis-2-butene

0.69

1,3-Butadiene

0.20

C3s

0.57

I-butane

4.42

n-butane

16.15

1-Butene 1-butene

37.22

I-butene

30.01

trans-2-butene

8.38

cis-2-butene

2.27

1,3-Butadiene

0.37

methyl tert butyl ether (MTBE)

0.61

On the other hand, streams suitable for the production of polyolefins generally contain feedstocks such as those described in Tables 4 and 5.

I-butane

2.19

n-butane

61.50

1-Butene 1-butene

0.64

trans-2-butene

28.18

cis-2-butene

7.49

1-decene 1-decene

94.00

C10 C10 isomers

6.00

Referring to Figure 2, and also in accordance with the concepts and principles of the invention, it was unexpectedly discovered that an operating system comprising several reactors arranged for operation in parallel provides greater operational flexibility than a system with a single, larger reactor of a size ensuring the same overall production rate. In fact, the principle of a multi-reactor system according to the invention ensures lower operating risk, greater process flexibility, lower feed rates (higher conversion rates), improved reactor design, and increased production capacity per unit time. Furthermore, the principle of a multi-reactor system according to the invention allows, for example, a 1:20 scale-up of pilot plant operation when the system includes two reactors, as well as a 1:40 scale-up using a larger reactor. . This principle greatly reduces the uncertainties of experimental industrial unit scale-up data.

A multi-reactor system embodying the concepts and principles of the invention is shown in Figure 2, where it is generally designated by reference number 200. System 200 includes reactors 202a and 202b, connected as shown in the drawing for operation in parallel on both the reaction side and the cooling fluid side. In addition, each of the reactors 202a, 202b, has its own desirable recycling system, 204a, 204b. Ideally, reactors 202a and 202b would be identical. However, according to the clear concepts of the invention, the identity of the two reactors is not a critical feature of the invention.

Ideally, any of the reactors 202a, 202b would be essentially identical to the reactor 100 shown in Figure 1. This means that each of the reactors 202a, 202b would be a two-pass reactor, where each pass includes On one hundred and ninety-four 0.952 cm (3/8 in) diameter tubes as described above. Other equipment shown in Figure 2 is designated and is essentially the same as the corresponding equipment shown in Figure 1 with the same reference numbers followed by either “A” or “B”, as the case may be. Accordingly, both reactors 202a and 202b include a feedstock inlet line (15a, 15b), a recirculation pump (25a, 25b), a suction line (20a, 20b) for the recirculation pump, a product outlet line (55a, 55b), Catalyst assembly inlet line (30a, 30b) and methanol inlet line (16a, 16b). In Figure 2, a common feedstock entry line for the multi-reactor system 200 is designated with reference number 215, and a common product exit line for the multi-reactor system 200 is designated with reference number 255.

The benefits of the multi-reactor system according to the invention are due to conversion rates and dispersive polymerization of the polymer. The multi-reactor system also facilitates a reduction in the amount of off-spec material generated during early unit operation because achieving equilibrium and developing the necessary operating parameters for a given product is faster.

The optimal inlet feedstock flow rate for each reactor in the multi-reactor system 200 according to the invention is approximately 56.781 to 64.35 l/min (15 to 17 gal/min) with appropriate refrigeration capacity and back-end processing capabilities. This means that using the multi-reactor system 200 according to the invention, it is possible to achieve conversion rates (70-75%) higher at this flow rate than can be achieved at higher flow rates such as greater than 75.71 l/min (> 20 gal/min). This is the result of increased residence times in the range of about 120 to 135 seconds. Higher conversion rates result in improvements (decreases) in dispersion polymerization, as a dispersion polymerization value of about 1.7 can be achieved by using the multi-reactor system 200 according to the invention to produce a PIB product having a number average molecular weight (MN) of about 950 A dispersion polymerization value of about 2.2 can be achieved by using the multi-reactor system 200 according to the invention to produce a PIB product having an MN of about 2300. The best dispersion polymerization values ​​that could be achieved when using a single reactor to produce products of similar molecular weight were 1.9 and 2.3, respectively.

For the two-reactor system described above, feedstock flow rates and product withdrawal rates can preferably be chosen so that the residence time of the reaction mixture in each reactor is, for example, about 4 minutes or less, 3 minutes or less, and can be chosen such that Residence time is ideally about 120 to about 135 seconds, perhaps even less than about 2 minutes, and perhaps as low as 1 minute or less.

The multi-reactor system 200 according to the invention also facilitates the use of smaller reactors that have improved pressure drop characteristics leading to more efficient energy consumption. This may be due at least in part to the fact that larger reactors may require the use of longer tubes with the same linear flow rates for recirculation.

Example

Tests were conducted to determine the improvements in operational characteristics that could be achieved through the use of a multi-reactor system, in this case two similar reactors operated in parallel. According to the testing protocol, the tests were conducted in three phases. In these phases, all modes of operation other than those specifically indicated are maintained constant. In the first phase, a single reactor was operated in such a way as to produce highly reactive polyisobutylene with a terminal double bond content of more than 70% and an MN value of approximately 1600. The feedstock was a concentrated stream of isobutylene and the recycling rate was maintained At a level that ensures a homogeneous mixture between the composition of the catalyst and the reactants and a suitable heat transfer coefficient to provide adequate cooling. The single reactor was initially operated at a feedstock inlet rate of approximately 102.2 l/min (27 gal/min). Then, this rate was increased to 121.1 l/min (32 gal/min). In the second phase, two reactors were operated in parallel. These two parallel reactors were essentially identical to the reactor used during the first phase. During this phase, the rate of feedstock entering each reactor was 56.78 l/min (15 gal/min). Again, the recirculation rate was maintained at a level to ensure a homogeneous mixing of catalyst composition and reactants and a suitable heat transfer coefficient to provide adequate cooling. In the third phase, the operating process was similar to that in the second phase. As an initial step in this third phase, the conversion rate was increased while the feed stream inlet rate was maintained at 56.78 l/min (15 gal/min), and then the supply of chilled water to the tube side of the two reactors was reduced to increase the conversion rate. Then, the feed stream entry rate into each reactor was increased to 64.35 L/min (17 gal/min). The results of these tests are summarized below in Table 6.

Table 6

Test Phase A

Test duration

Reactor

Hydrocarbon feed stream flow rate L/min (gal/min)

Reactor temperature m (F)

Percentage of thermal equilibrium conversion

Reactor production rate kg/min (lbs/min)

Cooling data

Flow rate l/min (gal/min)

Supply temperature m(F)

Quantity required kg (tons)

22 hour

a

120 (31.7)

21.11 (70)

40.6 (89.6)

2244.8 (593)

-1.4 (29.4)

164000 (164)

2

45 hour

a

B

56.8 (15)

56.8 (15)

113.6 (30)

18 (64.5)

18 (64.5)

20.8 (45.9)

20.8 (45.9)

41.6 (91.8)

972.9 (257)

953.9 (252)

1926.8 (509)

3.78 (38.8)

3.78 (38.8)

86000 (86)

87000 (87)

173000 (173)

103

22 hour

And 40 minutes

a

B

56.8 (15)

56.8 (15)

113.6 (30)

17.8 (64)

17.8 (64)

22.6 (49.9)

22.6 (49.9)

45.2 99.8

923.6 (244)

942.6 (249)

1866.2 (493)

-1.2 (29.8)

-1.2 (29.8)

95000 (95)

95000 (95)

190000 (190)

203

3 hours

And 10 minutes

a

B

64.3 (17)

64.3 (17)

128.6 (34)

17.8 (64)

17.8 (64)

23.9 (52.6)

23.9 (52.6)

47.8 (105.2)

923.6 (244)

942.6 (249)

1866.2 (493)

-0.39 (31.3)

-0.39 (31.3)

99000 (99)

99000 (99)

198000 (198)

Phase 1 - Single reactor, increasing feed rate to maximize reactor production rate

Phase 2 - Two reactors with feed rates at the baseline estimated at 56.78 liters/minute (15 gallons/minute), and at 3.78 C (38.8 F) for suddenly cooled water.

Phase 103 - Two reactors with a feed rate of 56.78 liters/minute (15 gallons/minute), suddenly reducing the temperature of the cooled water to increase the conversion rate to the maximum.

Phase 203 - Two reactors, increasing feed rate to maximize the reactor production rate

Table 6 represents the duration of each test phase, the flow rate of the feed stream, the reaction temperature, the heat balance conversion, the reactor make rate, and the refrigeration system data. The thermal equilibrium shift was estimated based on the flow rate of the feed stream, the temperature of the reaction, the flow rate of the suddenly cooled water, and the temperature increase through the reactor. The flow rate of the flash-cooled water and the increase in temperature determine the amount of heat generated by the reaction, and both the heat of reaction and the flow rate of the feed stream are used to calculate the percentage of the feed stream transferred to the PIB and oligomers. The reactor production rate, in kg/min (lb/min), is calculated from the feed rate and thermal equilibrium shift.

During single reactor operation, the reactor production rate was maximized at 39.9 kg/min (88 lb/min) at a feed rate of 120 L/min (31.7 gal/min). When the feed rate was increased to 121.1 L/min (32 gal/min), the reactor production rate began to decrease, so there was no further increase in the feed rate. The highest reactor production rate was achieved during the operation of the two reactors, with a feed rate for each of the two reactors reaching 64.35 liters/minute (17 gallons/minute). While when each of the two reactors was fed at a rate of 56.78 l/min (15 gal/min), the production rate for each reactor increased from 20.8 kg/min (45.9 lb/min) (i.e. a total production rate for the two reactors of 41.6 kg/min (91.8 lbs/min) to 22.6 kg/min (49.9 lbs/min) for one reactor (i.e. a total production rate for the two reactors of 45.2 kg/min (99.8 lbs/min)) by suddenly reducing the temperature of the cooled water supply from 3.3°C (38°F). ) to 1.1 m (30 F). The reactor production rate was further increased to 23.9 kg/min (52.6 lb/min) per reactor (i.e. a total production rate for the two reactors of 47.8 kg/min (105.2 lb/min)) by increasing the feed rates for each reactor from 56.78 L/min (15 gal/min) to 64.35 l/min (17 gal/min).

During Phase 1 of the test program, the feed rate was maintained at 115.5 kg/min (30.5 gal/min) for approximately 8 hours. A direct comparison can be made between the operation of one reactor and the operation of two reactors by comparing the transformation rate during this period with the transformation rate obtained during phase 3.1. During Phase 1, the feed rate was slightly higher at 115.5 L/min versus 113.6 L/min (30.5 versus 30 gal/min), but the flash-cooled water supply temperature was slightly lower at 2.2 C versus 1.1 C during Phase 3.1. (28 F versus 30 F during phase 3.1). In the operation of two reactors, the thermal equilibrium shift rate reached 73% compared to 64% for the operation of one reactor, although the reactor temperatures were lower by about 2.78°C (5°F) when operating two reactors (17.8°C versus 20.56°C (64°F versus 69°F). )). When operating two reactors, there is a doubling of the residence time and surface area for heat removal compared to the case of single reactor operation. The additional residence time explains why the reactor temperature must be lowered, and the additional surface area explains why the conversion rate is higher even at a lower reaction temperature.

Considering the above, it is readily apparent that by using two reactors in parallel, the transformation rates increase and the dispersion polymerization values ​​decrease relative to those when using a single reactor. This result was achieved because the principle of a multi-reactor system facilitates feeding at a lower rate per reactor, accompanied by an increase in residence time.

As mentioned above, the product leaving the polymerization reaction system via line 55 (Figure 1) or 255 (Figure 2) must be immediately quenched with a substance capable of quenching the activity of the catalyst, such as, for example, ammonium hydroxide. Any possible unwanted decrease in molecular weight or rearrangement of the polymer molecules can thus be reduced to a minimum. The polyolefin products according to the invention can then be directed to a work up treatment system, including a washing system as described below where the catalyst salts can be removed.

In Fig. 3, a washing system embodying the concepts and principles of another aspect of the invention is generally designated by reference number 300. As shown, the system 300 includes a pre-sedimentation vessel 302 and a post-sedimentation system 304 comprising, in the preferred embodiment of this aspect of the invention shown in Fig. 3, On two subsequent sedimentation vessels 306, 308. It is noted here that, alternatively, the subsequent sedimentation system 304 can also contain only one sedimentation vessel or three or more sedimentation vessels, depending on the nature of the product and the nature of the remaining catalysts to be removed from it.

The system 300 further includes an inlet line 310 connecting line 55 or line 255, as the case may be, and a suction line 311 for a pump 312 that pumps the raw product and its admixtures to the settling vessel 302 via line 314. A worker is introduced to line 310 via Pump 316 and line 318 for quenching any remaining catalyst in the crude product entering system 300 via line 310. NH4OH in aqueous solution is a particularly good agent for quenching any remaining BF3/methanol complex in the polyethylene product. However, the invention does not limit any method to using NH4OH. Rather, the actual nature of the catalyst quenching factor will depend entirely on the nature of the catalyst itself and/or the nature of the product in the product stream.

Wash water is added to the raw product in line 316 through line 320 and mixed with it. The mixture consisting of the raw product containing the remaining catalyst composition, catalyst quenching agent and wash water is fed to pump 312 via suction line 311. Desirably, pump 312 can be a centrifugal pump where the rotation of the impellers ensures that the water, the catalyst salts resulting from the reaction between the catalyst quenching agent and the catalyst and the raw polyolefin product are mixed uniformly so that complete washing is achieved. . In addition, pump 312 may be equipped with a recirculation line 322, including a flow controlling device 324, to return a portion of the mixture to the suction pump for additional mixing.

A mixture of the hydrocarbon product, damper catalyst salts, and water is added via line 314 to an internal settling chamber of the settling vessel 302, where the hydrocarbon phase is separated from the aqueous phase under the influence of gravitational forces in a manner known by itself. Desirably, with respect to the latter, the reaction between the quenching agent and the catalyst forms a water-soluble salt with most of this salt present in the aqueous phase.

The upper phase of the partially washed raw polyolefin product is removed from the vessel 302 via an overhead line 326 and the aqueous phase exits the vessel 302 via line 328. A portion of the removed aqueous phase is recycled to the wash water inlet line 320 via a return line 330 and a flow control. flow controller 332. Another portion of the removed aqueous phase is discharged from the system via a drain line 334 and a level controller 336 that adjusts the level of the aqueous phase in the vessel 302. Drainage line 334 is connected to a system (not shown in the drawing) to either extract or dispose of spent and contaminated wash water.

Make-up wash water for vessel 302, which may preferably be demineralized water, is added to the drain water recycled in line 320 via line 321. In this regard, it should be noted that the relevant amounts of make-up wash water must be adjusted, The catalyst quenching agent entering the system and the purged aqueous phase to ensure that the amount of quenching agent entering the system is always present in excess of the amount of catalyst remaining in the raw product.

The partially washed crude polyolefin product phase in line 326 is fed to the suction line 336 of the pump 338 with additional wash water drained through line 340. Desirably the pump 338 can be a centrifugal pump like the pump 312 to ensure mixing The aqueous phase is combined with the hydrocarbon phase before the mixture is discharged through line 342 to vessel 306. The pump 338 may be equipped with a recirculation line 344 and a flow controlling device 346 to return a portion of the mixture to the suction line 336 of the pump for additional mixing. The two phases of the mixture in the vessel 306 are allowed to separate under the influence of gravity to form an upper hydrocarbon phase 306 and a lower aqueous phase.

The largely completely washed crude polyolefin product phase is removed from vessel 306 via another overhead line 348, and the lower precipitated aqueous phase exits vessel 306 via line 350. A portion of the removed aqueous phase is recycled to the wash water inlet line 340 via line Return 352 Another portion of the aqueous phase removed from the system is drained through a drain line 354 and a level adjuster 356 that adjusts the level of the aqueous phase in the vessel 306. The drain line 354 is connected to the drain line 334.

The largely completely washed crude polyolefin product phase is mixed in line 348 with additional wash water introduced through line 358. The mixture consisting of the crude polyolefin product phase and additional wash water is fed to the settling vessel 308 via line 359 where separation is permitted. Develop the mixture again under the influence of gravity. The lower aqueous phase is removed from the vessel 308 by the action of a pump 362 via a bottom line 360, and a portion of it is recycled to the line 348 via a flow control device 364, a return line 366, and a line 358. Another portion of the aqueous phase leaving vessel 308 is recycled through line 367 and flow control 369 and is introduced to line 340 to be used as make-up wash water in vessel 306.

The completely washed crude polyolefin product is removed from vessel 308 via an overhead line 374 and fed into a subsequent purification system (not shown) to remove diluents, unreacted monomers, and unwanted volatiles such as dimers, trimers, and oligomers. , and others.

New make-up wash water is introduced to vessel 308, which again may be desirable demineralized water, via line 368. Make-up water is introduced into the system via pump 370 and line 372. In this regard, it should be noted that line 372 is connected to Line 321 for supplying new make-up water to the preceding vessel 302 and with line 368 for separate and independent supply of new make-up water to the successor sedimentation system 304. Accordingly, another make-up water can be introduced into the post-sedimentation system 304 without unnecessary dilution of the catalyst quenching agent (NH4OH) required in the pre-sedimentation vessel 302. This result is achieved because the wash system of the previous vessel is operated entirely and independently of the wash system of the after-sedimentation system 304. It is also worth noting that the concentration of the catalyst quenching agent in the aqueous phase of the precipitation vessel 302 must always be in excess of the amount of catalyst remaining. Furthermore, the concentration of catalyst salts in the aqueous phase must always be low enough to avoid precipitation. Accordingly, the amount of new make-up water introduced into the previous settling vessel 302 must be carefully controlled, while the amount of new make-up water introduced into the subsequent settling system must be generous and limited only when the largest possible amount of contaminants needs to be removed of the final product. Therefore, new make-up water with a small flow is used in the previous settling vessel to minimize the use of catalyst quenching agent, while new make-up water with a very large flow is used in the subsequent settling vessel to provide better washing.