Process for dehydrogenation of hydrocarbons with multiple catalyst inlets

Abstract

A process for producing one or more olefin compounds can include dehydrogenating a feed stream in a reactor in the presence of a catalyst to form a product stream and a deactivated catalyst. The feed stream may comprise one or more hydrocarbons comprising alkyl moieties. The product stream may include one or more olefin compounds. The method may also include separating the deactivated catalyst into a first portion of deactivated catalyst and a second portion of deactivated catalyst. The method may include passing the second portion of the deactivated catalyst to a regenerator. The method may include treating the second portion of the deactivated catalyst in the regenerator to form a regenerated catalyst. The method may also include passing the first portion of the deactivated catalyst and the regenerated catalyst to the reactor. The first portion of deactivated catalyst may enter the reactor downstream of the regenerated catalyst relative to the flow direction of the feed stream. The first portion of the deactivated catalyst may have a lower temperature than the regenerated catalyst.

Description

[0001] Cross - reference to related applications

[0002] This application claims the benefit of U.S. Provisional Application Serial No. 63 / 428,521, filed on November 29, 2022, the entire disclosure of which is hereby incorporated by reference herein. Technical field

[0003] The embodiments described herein generally relate to chemical processing, and more particularly, to methods and systems for dehydrogenating chemical substances. Background art

[0004] Olefin compounds can be used as base materials to produce many types of goods and materials. For example, propylene can be used to manufacture polypropylene, propylene oxide, and acrylonitrile. Such products can be used in product packaging, chemical manufacturing, textiles, etc. Thus, there is an industrial need for olefin compounds such as ethylene, propylene, butene, and styrene, as well as methods for producing such materials. Summary of the invention

[0005] One method for producing olefin compounds is by dehydrogenating hydrocarbons. In some embodiments, the dehydrogenation reaction can be facilitated by using a solid - particle catalyst in a circulating fluidized - bed (CFB) system. In an embodiment, when the catalyst is used in the dehydrogenation reaction, it may become deactivated. Such deactivated catalyst can be transferred to a regenerator to restore at least a portion of the catalyst's activity, such as by decoking the catalyst or heating the catalyst. Alternatively, some deactivated catalyst can be recycled and reused in the dehydrogenation reaction without regeneration.

[0006] As described herein, it has been found that it may be beneficial to utilize a specific catalyst distribution pattern relative to the recycle reaction catalyst, the regenerated catalyst, and the feed stream entering the reactor. Embodiments described herein include a catalyst distribution pattern in which the recycle reaction catalyst may enter the reactor downstream of the regenerated catalyst relative to the flow direction of the feed stream. Such a catalyst distribution pattern may, for example, improve the catalyst activity in the reactor and increase the yield of the dehydrogenation reaction when compared to a catalyst distribution pattern that does not have the recycle reaction catalyst enter the reactor downstream of the regenerated catalyst relative to the flow direction of the feed stream. Generally, the recycle reaction catalyst is colder and has lower catalytic activity than the regenerated catalyst. Those skilled in the art would expect that premixing the two catalyst streams would help reduce the spatial variations in catalyst activity and temperature in the reactor and would be beneficial to the process. However, contrary to what those skilled in the art would expect, the method described herein unexpectedly provides excellent yields by having the less catalytically active recycle reaction catalyst contact the feed only after the more catalytically active regenerated catalyst has contacted the feed when compared to embodiments in which the recycle reaction catalyst is premixed with the regenerated catalyst.

[0007] According to one or more embodiments described herein, a method for producing one or more olefinic compounds may include dehydrogenating a feed stream in a reactor in the presence of a catalyst to form a product stream and deactivated catalyst. The feed stream may comprise one or more hydrocarbons comprising an alkyl moiety. The product stream may comprise one or more olefinic compounds. The method may further include separating the deactivated catalyst into a first portion of deactivated catalyst and a second portion of deactivated catalyst. The method may include passing the second portion of deactivated catalyst to a regenerator. The method may include treating the second portion of deactivated catalyst in the regenerator to form regenerated catalyst. The method may further include passing the first portion of deactivated catalyst and the regenerated catalyst to the reactor. The first portion of deactivated catalyst may enter the reactor downstream of the regenerated catalyst relative to the flow direction of the feed stream. The first portion of deactivated catalyst may have a lower temperature than the regenerated catalyst.

[0008] It should be understood that both the foregoing general description and the following detailed description describe various embodiments and are intended to provide an overview or framework for understanding the nature and characteristics of the claimed subject matter. Additional features and advantages of the embodiments will be set forth in the detailed description, and in part, will be readily apparent to those of ordinary skill in the art from the description including the drawings and the claims, or may be recognized by practicing the described embodiments. Drawings are included to provide a further understanding of the embodiments and, together with the detailed description, serve to explain the principles and operation of the claimed subject matter. However, the embodiments depicted in the drawings are illustrative and exemplary in nature and are not intended to limit the claimed subject matter. BRIEF DESCRIPTION OF THE DRAWINGS

[0009] Reference will now be made in more detail to various embodiments, some of which are illustrated in the accompanying drawings, in which:

[0010] Figure 1 schematically depicts a reactor system according to one or more embodiments of the present disclosure;

[0011] Figure 2 schematically depicts a reactor according to one or more embodiments of the present disclosure;

[0012] Figure 3A A continuous stirred tank reactor system configuration is schematically depicted;

[0013] Figure 3B schematically depicts a continuous stirred tank reactor system configuration; and

[0014] Figure 3C A continuous stirred tank reactor system configuration is schematically depicted.

[0015] When describing Figure 1 , Figure 2 , Figure 3A , Figure 3B and Figure 3C While the simplified schematic illustration of FIG. 1 does not include the many valves, temperature sensors, electronic controllers, etc. that may be used and are well known to those of ordinary skill in the art, are not included. In addition, the accompanying components that are typically included in such reactor systems, such as, air supply, heat exchangers, buffer tanks, etc., are also not included. However, it should be understood that these components are within the scope of the present disclosure. DETAILED DESCRIPTION

[0016] The specific embodiments of the application will now be described. However, the present disclosure can be implemented in different forms and should not be construed as being limited to the embodiments set forth in the present disclosure. On the contrary, these embodiments are provided so that the present disclosure will be thorough and complete, and will fully convey the scope of the subject matter to those skilled in the art.

[0017] This article will now focus on a process for operating a fluidized dehydrogenation reactor system to produce olefin compounds such as propylene. Figure 1 Embodiments of the present disclosure are described in detail in the context of a reactor system for a reactor system. However, it should be understood that the principles disclosed and taught herein can be applied to other systems utilizing different system components oriented in different ways, or to different reaction schemes utilizing various catalyst compositions. For example, the concepts described can be equally applied to other systems having alternative reactor units and regeneration units, such as those operating under non-fluidized conditions, or those systems including downcomers instead of risers. It should also be understood that not all systems are identical.Figure 1 All parts of the reactor system should be construed as being essential to the claimed subject matter.

[0018] Now referring to Figure 1 , an example reactor system 103 suitable for use with the methods and / or apparatuses described herein is schematically depicted. The reactor system 103 generally includes a plurality of system components, such as a reactor section 206 and a regeneration unit 306. As used herein, a "system component" refers to a part of the reactor system 103, such as a reactor, a separator, a transfer line, and combinations thereof, etc. As used herein in the context of Figure 1 , the reactor section 206 generally refers to the part of the reactor system 103 that conducts the main process reaction (e.g., dehydrogenation) to form a product stream. A feed stream enters the reactor section 206 via a feed inlet 434, is converted into a product stream (containing products and unreacted feed), and exits the reactor section 206. The reactor section 206 includes a reactor 202, which may include an upstream reactor section 254 and a downstream reactor section 230. According to one or more embodiments, as depicted in Figure 1 , the reactor section 206 may include a catalyst separation section 214 for separating the catalyst from the chemical products formed in the reactor 202. The catalyst that exits the reactor 202 and is transferred to the catalyst separation section 214 may be deactivated and may, for example, be colder or less active than the catalyst transferred to the reactor 202 from the regeneration unit 306 via line 424.

[0019] In one or more embodiments, the catalyst may be separated into multiple portions of catalyst and thus exit via lines 422 and 426. As used herein, the "first portion" of the deactivated catalyst is transferred back to the reactor 202 in a recycle stream via line 422 (without entering the burner 355). The "second portion" of the deactivated catalyst is transferred to the regeneration unit 306 via line 426. As used herein, the first portion of the catalyst and the "regenerated catalyst" (i.e., the regenerated form of the second portion of the catalyst transferred to the reactor 202 via line 424) are generally introduced into the reactor 202 separately and at different parts (such as different heights) of the reactor 202. Such a configuration may have a positive impact on catalytic efficiency.

[0020] As used herein, the regeneration unit 306 generally refers to a part of the reactor system 103 where the catalyst is treated in some manner, such as by combustion, to, for example, improve the catalytic activity and / or heat the catalyst. The regeneration unit 306 can include a combustor 355 and a riser 330, a granular solid separation section 316, and can additionally include an oxygen treatment zone 370. In one or more embodiments, the granular solid separation section 214 can be in fluid communication with the combustor 355 (e.g., via a standpipe 426), and the granular solid separation section 316 can be in fluid communication with the upstream reactor section 250 (e.g., via a standpipe 424 and a transport riser 430).

[0021] Generally as described herein, in Figure 1 the illustrated embodiment, the catalyst is circulated between the reactor section 200 and the regeneration unit 300. It should be understood that when referring to catalysts herein, they can refer to solid materials that are catalytically active for the desired reaction, or can equally refer to other granular solids that are mentioned with respect to the Figure 1 system and that are not necessarily catalytically active but affect the reaction. The terms "catalytic activity" and "catalyst activity" refer to the extent to which a catalyst is able to catalyze the reactions taking place in the reactor system 106. In an embodiment, the deactivated catalyst of the second part can be reactivated by catalyst reactivation in the regeneration unit 306. Reactivation (sometimes referred to herein as "regeneration") can remove contaminants such as coke, raise the temperature of the granular solid, or both. The deactivated catalyst can be reactivated by, but not limited to, removing coke by combustion, restoring the catalyst acidity, oxidizing the granular solid, other reactivation processes, or combinations thereof. The regenerated catalyst from the regeneration unit 306 can then be transferred back to the reactor section 202 via line 424.

[0022] Now referring in detail to Figure 1 , the reactor section 206 can include an upstream reactor section 254, a transition section 258, and a downstream reactor section 230, such as a riser. The transition section 258 can connect the upstream reactor section 254 to the downstream reactor section 230. As Figure 1 depicted, the upstream reactor section 254 can be positioned below the downstream reactor section 230. Such a configuration can be referred to as an upflow configuration in the reactor 202. The upstream reactor section 254 can include a vessel, a drum, a barrel, a tank, or other vessel suitable for a given chemical reaction. As Figure 1As depicted, the upstream reactor section 254 can be connected to the downstream reactor section 230 via a transition section 258. The upstream reactor section 254 can generally include a larger cross-sectional area than the downstream reactor section 230. The transition section 258 can taper from the dimensions of the cross-section of the upstream reactor section 254 to the dimensions of the cross-section of the downstream reactor section 230 such that the transition section 258 projects inwardly from the upstream reactor section 254 into the downstream reactor section 230. For example, the transition section 258 can be a frustum.

[0023] As described with respect to Figure 1 As described, the feed stream can enter the reactor 202 via the feed inlet 434 and the product stream can exit the reactor system 103 via the tube 420. According to one or more embodiments, the reactor system 103 can be operated by feeding a chemical feed (e.g., in the feed stream) and a fluidized catalyst into the upstream reactor section 254. The chemical feed contacts the catalyst in the upstream reactor section 254 and each chemical feed flows upwardly into and through the downstream reactor section 230 to produce a chemical product.

[0024] The upstream reactor section 254 can be connected to a transport riser 430 which can provide regenerated catalyst in the feed stream to the reactor section 206 in operation. In one or more embodiments, a first portion of deactivated catalyst can enter the reactor downstream of the regenerated catalyst with respect to the flow direction of the feed stream. As will be described herein, in one or more embodiments, the regenerated catalyst and the first portion of deactivated catalyst can enter the reactor 202 via a granular solid distributor 100. The granular solid distributor 100 can deliver the regenerated catalyst and the first portion of deactivated catalyst separately into the reactor 202. The catalyst entering the upstream reactor section 254 via the transport riser 430 can be delivered to the transport riser 430 via a line 424 and thus arrives from the regeneration unit 306. The first portion of deactivated catalyst can come directly from the catalyst separation section 214 via a line 422 and enter the transport riser 430 where the catalyst enters the upstream reactor section 254. This catalyst can be slightly deactivated but can still be suitable for reaction in the upstream reactor section 254 in some embodiments, particularly when used in combination with the regenerated catalyst. The regenerated catalyst arriving from the regeneration unit 306 and the first portion of deactivated catalyst arriving from the catalyst separation section 214 via the line 422 can remain separated within the transport riser 430 before being delivered into the reactor 202 via the granular solid distributor 100.

[0025] Still referring to Figure 1, in one or more embodiments, based on the shape, size, and other processing conditions (such as temperature and pressure) in the upstream reactor section 254 and the downstream reactor section 230, the upstream reactor section 254 can be operated as a fluidized bed, such as in a fast fluidized, turbulent, or bubbling bed upflow reactor, while the downstream reactor section 230 can be operated in a more plug flow manner, such as in a riser reactor. For example, Figure 1 The reactor 202 can include an upstream reactor section 254 that operates as a fast fluidized, turbulent, or bubbling bed reactor and a downstream reactor section 230 that operates as a dilute-phase riser reactor, such that the average catalyst and gas flow move upward simultaneously. As used herein, the term "average flow" refers to the net flow, i.e., the total upward flow minus the reverse or backward flow, which is typically characteristic of fluidized particles. As described herein, a "fast fluidized" reactor can refer to a reactor that utilizes a fluidization scheme in which the superficial velocity of the gas phase is greater than the choking velocity and can be semi-dense in operation. As described herein, a "turbulent" reactor can refer to a fluidization scheme in which the superficial velocity is less than the choking velocity and is denser than the fast fluidization scheme. As described herein, a "bubbling bed" reactor can refer to a fluidization morphology in which clearly defined bubbles exist in a highly dense bed in two clearly demarcated phases. The "choking velocity" is the minimum velocity required to maintain solids in a dilute-phase mode in a vertical transport line. As described herein, a "dilute-phase riser" can refer to a riser reactor that operates at a transport velocity, where the gas and catalyst have approximately the same velocity in the dilute phase.

[0026] According to an embodiment, the chemical product and the catalyst can be transferred from the downstream reactor section 230 to the separation device 226 in the catalyst separation section 214, where the catalyst is separated from the chemical product, and the chemical product is transported out of the catalyst separation section 214. According to one or more embodiments, after separation from the vapor in the separation device 226, the deactivated catalyst typically moves through the stripping zone 224, where a first portion of the deactivated catalyst is transferred from the stripping zone 224 to the reactor 202 via line 422, and a second portion of the deactivated catalyst is transferred to the catalyst outlet port 222, where the second portion of the deactivated catalyst is conveyed out of the reactor section 206 via line 426 and into the regeneration unit 306.

[0027] Now return to reference Figure 1, according to one or more embodiments, the separation device 226 can be a cyclone separation system, which can include two or more cyclone separation stages. In embodiments where the separation device 226 includes more than one cyclone separation stage, the first separation device into which the fluidized stream enters is referred to as the primary cyclone separation device. The fluidized effluent from the primary cyclone separation device can enter a secondary cyclone separation device for further separation. The primary cyclone separation device can include, for example, a primary cyclone separator and systems commercially available under the names VSS (commercially available from UOP), LD2 (commercially available from Stone and Webster), and RS2 (commercially available from Stone and Webster). The primary cyclone separator is described, for example, in U.S. Patent No. 4,579,716, U.S. Patent No. 5,190,650, and U.S. Patent No. 5,275,641, the entire texts of each of which are incorporated herein by reference. In some separation systems that utilize a primary cyclone separator as the primary cyclone separation device, one or more additional sets of cyclone separators, such as secondary and tertiary cyclone separators, are used to further separate the catalyst from the product gas. It should be understood that any primary cyclone separation device can be used in the embodiments of the present invention.

[0028] As previously described herein, the deactivated catalyst in the catalyst separation section 214 can be separated into a first portion of the deactivated catalyst that is transferred back to the reactor 202 via line 422 and a second portion of the deactivated catalyst that is transferred to the combustor 350 via line 426. The mass flow rate ratio of the first portion of the deactivated catalyst to the second portion of the deactivated catalyst can be from 0.1 to 5. For example, the mass flow rate ratio of the first portion of the deactivated catalyst to the second portion of the deactivated catalyst can be from 0.1 to 4.5, such as from 0.1 to 4, from 0.1 to 3.5, from 0.1 to 3, from 0.1 to 2.5, from 0.1 to 2, from 0.1 to 1.5, from 0.1 to 1, from 0.1 to 0.5, from 0.5 to 5, from 0.5 to 4.5, from 0.5 to 4, from 0.5 to 3.5, from 0.5 to 3, from 0.5 to 2.5, from 0.5 to 2, from 0.5 to 1.5, from 0.5 to 1, from 1 to 5, from 1 to 4.5, from 1 to 4, from 1 to 3.5, from 1 to 3, from 1 to 2.5, from 1 to 2, from 1 to 1.5, from 1.5 to 5, from 1.5 to 4.5, from 1.5 to 4, from 1.5 to 3.5, from 1.5 to 3, from 1.5 to 2.5, from 1.5 to 2, from 2 to 5, from 2 to 4.5, from 2 to 4, from 2 to 3.5, from 2 to 3, from 2.5 to 5, from 2.5 to 4.5, from 2.5 to 4, from 2.5 to 3.5, from 2.5 to 3, from 3 to 5, from 3 to 4.5, from 3 to 4, from 3 to 3.5, from 3.5 to 5, from 3.5 to 4.5, from 3.5 to 4, from 4 to 5, from 4 to 4.5, from 4.5 to 5 or any combination of these ranges. The mass flow rate of the second portion of the deactivated catalyst that is transferred to the combustor 350 can be substantially the same as the mass flow rate of the regenerated catalyst that is transferred to the reactor 202 via line 424, as discussed later herein.

[0029] In one or more embodiments, the temperature of the deactivated catalyst (delivered via line 422) in the first section can be from 580 °C to 800 °C, such as from 580 °C to 775 °C, from 580 °C to 750 °C, from 580 °C to 725 °C, from 580 °C to 700 °C, from 580 °C to 675 °C, from 580 °C to 650 °C, from 580 °C to 625 °C, from 580 °C to 600 °C, from 600 °C to 800 °C, from 600 °C to 775 °C, from 600 °C to 750 °C, from 600 °C to 725 °C, from 600 °C to 700 °C, from 600 °C to 675 °C, from 600 °C to 650 °C, from 600 °C to 625 °C, from 625 °C to 800 °C, from 625 °C to 775 °C, from 625 °C to 750 °C, from 625 °C to 725 °C, from 625 °C to 700 °C, from 625 °C to 675 °C, from 625 °C to 650 °C, from 650 °C to 800 °C, from 650 °C to 775 °C, from 650 °C to 750 °C, from 650 °C to 725 °C, from 650 °C to 700 °C, from 650 °C to 675 °C, from 675 °C to 800 °C, from 675 °C to 775 °C, from 675 °C to 750 °C, from 675 °C to 725 °C, from 675 °C to 700 °C, from 700 °C to 800 °C, from 700 °C to 775 °C, from 700 °C to 750 °C, from 700 °C to 725 °C, from 725 °C to 800 °C, from 725 °C to 775 °C, from 725 °C to 750 °C, from 750 °C to 800 °C, from 750 °C to 775 °C, from 775 °C to 800 °C, or any combination of these ranges. This is typically lower than the temperature of the regenerated catalyst delivered to the reactor 202 via line 424.

[0030] Still referring to Figure 1, The deactivated catalyst of the second part can be transferred from the catalyst separation section 214 to the combustor 355 via pipeline 426. In the combustor 355, the second part of the deactivated catalyst can be processed by, for example, burning any coke on the catalyst using oxygen or using supplementary fuel. For example but not limited to, the catalyst can be decoked and / or the fuel can be burned to heat the catalyst. Then, the catalyst is transferred out of the combustor 355 and through the riser 330 to the riser termination separator 378, where the gas and solid components from the riser 330 are at least partially separated. The vapor and the remaining solids are transported to the secondary separation device 326 in the catalyst separation section 316, where the remaining catalyst is separated from the gas from the catalyst treatment (e.g., the gas emitted by burning the used catalyst or fuel, which is referred to as flue gas herein). The flue gas can be transferred out of the regeneration unit 306 via the outlet pipe 432. The separated catalyst is then transferred through the oxygen treatment zone 370 within the catalyst separation section 316 via pipeline 424 and the transport riser 430 to the upstream reactor section 254, where the catalyst is further used for catalytic reactions. The catalyst treated as described is referred to as "regenerated catalyst" and is transferred back to the reactor 202 from the deactivated catalyst of the first part in a separate stream.

[0031] In one or more embodiments, when the regenerated catalyst is transferred from the regeneration unit 306 to the reactor 202 via pipeline 424, it can have a temperature of 680 °C to 900 °C. For example, the regenerated catalyst can have a temperature of 680 °C to 875 °C, such as 680 °C to 850 °C, 680 °C to 825 °C, 680 °C to 800 °C, 680 °C to 775 °C, 680 °C to 750 °C, 680 °C to 725 °C, 680 °C to 700 °C, 700 °C to 900 °C, 700 °C to 875 °C, 700 °C to 850 °C, 700 °C to 825 °C, 700 °C to 800 °C, 700 °C to 775 °C, 700 °C to 750 °C, 700 °C to 725 °C, 725 °C to 900 °C, 725 °C to 875 °C, 725 °C to 850 °C, 725 °C to 825 °C, 725 °C to 800 °C, 725 °C to 775 °C, 725 °C to 750 °C, 750 °C to 900 °C, 750 °C to 875 °C, 750 °C to 850 °C, 750 °C to 825 °C, 750 °C to 800 °C, 750 °C to 775 °C, 775 °C to 900 °C, 775 °C to 875 °C, 775 °C to 850 °C, 775 °C to 825 °C, 775 °C to 800 °C, 800 °C to 900 °C, 800 °C to 875 °C, 800 °C to 850 °C, 800 °C to 825 °C, 825 °C to 900 °C, 825 °C to 875 °C, 825 °C to 850 °C, 850 °C to 900 °C, 850 °C to 875 °C, 875 °C to 900 °C, or any combination of these ranges.

[0032] Now referring to the regeneration unit 306, as Figure 1 depicted, the burner 355 of the regeneration unit 306 may include one or more lower burner inlet ports 356 and may be in fluid communication with the riser 330. An oxygen-containing gas (such as air) may be delivered to the burner 350 through the pipe 428. The burner 355 may be in fluid communication with the catalyst separation section 214 via a pipeline 426, which may supply a second portion of the deactivated catalyst from the reactor section 206 to the regeneration unit 306 for regeneration. The burner 355 and the riser 330 (collectively referred to as the catalyst combustion reactor 302) may be operated with a fluidization scheme similar to or the same as that disclosed for the upstream reactor section 254 and the downstream reactor section 230 of the reactor section 206. That is, the burner 350 may be operated as a fluidized bed, for example, in a fast fluidization, turbulent, or bubbling bed upflow reactor, while the riser 330 may be operated in a more plug flow manner, for example, in a riser reactor. The geometries described for the upstream reactor section 254 and the downstream reactor section 230 may equally apply to the burner 355 and the riser 330. Additionally, the burner 355 may further include a fuel inlet 354, which may supply a fuel such as a hydrocarbon stream or hydrogen to the burner 355.

[0033] As described in one or more embodiments, after separating the catalyst from the flue gas in the riser termination separator 378 and the secondary separation device 326, the treated catalyst is treated with an oxygen-containing gas in the oxygen treatment zone 370. In some embodiments, the oxygen treatment zone 370 includes a fluid-solid contact device. The fluid-solid contact device may include baffle or grid structures to facilitate the contact between the treated catalyst and the oxygen-containing gas. Examples of the fluid-solid contact device are further described in detail in U.S. Patent No. 9,827,543 and U.S. Patent No. 9,815,040. The fluidization scheme within the oxygen treatment zone 370 may be of the bubbling bed type fluidization.

[0034] As described herein, the regenerated catalyst and the first portion of the deactivated catalyst enter the reactor 202 in separate streams. In one or more embodiments, the first portion of the deactivated catalyst and the regenerated catalyst may constitute at least 95 wt% of the catalyst delivered to the reactor 202. For example, the first portion of the deactivated catalyst and the regenerated catalyst may constitute at least 95 wt%, at least 96 wt%, at least 97 wt%, at least 98 wt%, at least 99 wt%, or even at least 99.9 wt% of the catalyst delivered to the reactor 202.

[0035] As described below with respect to Figure 2As described in a particular embodiment, in one or more embodiments, the deactivated catalyst and the regenerated catalyst of the first portion may enter the reactor 202 through separate distributors. In one or more embodiments, the deactivated catalyst and the regenerated catalyst of the first portion may not enter the reactor 202 through the bottom end of the reactor 202.

[0036] Now referring to Figure 2 , an example reactor 202 that may be used with a Figure 1 reactor system 103 is schematically depicted. The reactor system includes a granular solid distributor 100 adapted to deliver the regenerated catalyst and the first portion of the deactivated catalyst at different heights. Figure 2 An envisioned granular solid distributor is depicted. However, other solid distributors may be suitable, and the embodiments described herein should not be construed as being limited by the design, shape, size, architecture, etc. of the one or more distributors that deliver the granular solids to the burner 350.

[0037] Now referring to Figure 2 , the reactor 202 of the Figure 1 reactor system 103 is schematically depicted. The first portion of the deactivated catalyst 104 and the regenerated catalyst 105 may be delivered to the reactor 202 through the granular solid distributor 100. The granular solid distributor 100 may deliver the first portion of the deactivated catalyst 104 to the reactor 202 downstream of the regenerated catalyst 105 relative to the flow direction of the feed stream, which enters the reactor 202 through the feed inlet 434 and through the feed distribution plate 450. The first portion of the deactivated catalyst 104 may be delivered upward from the inner section of the granular solid distributor 100 to the inner conduit outlet 220. Then the first portion of the deactivated granular solid 104 may contact the first catalyst deflector 240, which may direct the first portion of the deactivated catalyst 104 into the reactor 202. The regenerated catalyst 104 may pass through the outer section of the granular solid distributor to contact the second catalyst deflector 340, which may direct the regenerated catalyst 105 into the reactor 202. As Figure 2As shown, in one or more embodiments, the regenerated catalyst 105 can enter the reactor 202 between the feed stream and the first portion of the deactivated catalyst 104. When the feed stream is transferred from the feed inlet 434 into the reactor 202, it travels in an upward direction from the upstream reactor section 254 to the downstream reactor section 230. As the feed stream travels through the reactor 202, it first contacts the regenerated catalyst 105 that is directed into the reactor 202 by the second catalyst director 340. At least a portion of the feed stream can react with the regenerated catalyst 105 to produce one or more products. The mixed stream that can include one or more products, the regenerated catalyst 105, and the feed stream continues to travel upward through the reactor to contact the first portion of the deactivated catalyst 104 that is directed into the reactor 202 by the first catalyst director 240. The first portion of the deactivated catalyst 104 can contact the mixed stream and can react with the unreacted feed from the feed stream to produce one or more products. The mixture of the regenerated catalyst 105, the first portion of the deactivated catalyst 104, one or more products, and any remaining unreacted feed travels to the upstream reactor section 230, where the mixture can be transferred to Figure 1 the catalyst separation section 214 of the reactor system 103.

[0038] Without being bound by any particular theory and referring to the following examples, it was unexpectedly found that in some embodiments, compared to embodiments that do not include the described catalyst input mode, the catalyst input mode described herein can produce excellent results. Specifically, it would be expected that introducing a recycle catalyst at a lower temperature upstream of a regenerated catalyst at a higher temperature would provide better selectivity and subsequently higher yields. It would also be expected that premixing the two catalyst streams would help reduce the spatial variation of catalyst activity and temperature in the reactor and would be beneficial for better yields. However, as demonstrated in this example, unexpectedly, the opposite is the case.

[0039] In one or more additional embodiments, hydrocarbons such as, for example, methane and ethane may be entrained in the deactivated catalyst of the first portion. In one or more embodiments, hydrocarbons may be entrained therein when the deactivated catalyst of the first portion is transferred back to the reactor 202 via line 422. In the regeneration unit, the entrained hydrocarbons in the deactivated catalyst of the second portion may be combusted, and the regenerated catalyst transferred back to the reactor 202 from the regeneration unit 306 via line 424 may not have or may have substantially no entrained hydrocarbons, such as less than 0.05 mole % of entrained hydrocarbons in the regenerated catalyst. However, there may be an oxygen-containing gas entrained by the regenerated catalyst particles. Without wishing to be bound by theory, it is believed that when the hydrocarbons entrained in the deactivated catalyst of the first portion contact the regenerated catalyst entraining oxygen at high temperatures (such as in the comparative embodiment), coke and steam may form on the regenerated catalyst. This may cause at least a portion of the regenerated catalyst to be partially deactivated. For example, contacting the high-temperature regenerated catalyst with hydrocarbons may partially deactivate the regenerated catalyst before it has a chance to contact the feed stream in the reactor, which may reduce the efficiency of the reactor system 103. It is believed that introducing the deactivated catalyst of the first portion and the regenerated catalyst into the reactor at different locations may prevent or reduce premature deactivation of the regenerated catalyst because the entrained hydrocarbons in the deactivated catalyst of the first portion may not contact the regenerated catalyst before the regenerated catalyst enters the reactor 202. Additionally, since the deactivated catalyst of the first portion may be transferred to the reactor 202 downstream of the regenerated catalyst relative to the flow direction of the feed stream, the regenerated catalyst may contact the feed stream before contacting the deactivated catalyst of the first portion and any entrained hydrocarbons therein.

[0040] As described with respect to Figure 2 above, in one or more embodiments, the reactor 202 may include a feed distribution plate 450, as Figure 1 and Figure 2 shown. In one or more embodiments, where the reactor 202 includes a feed distribution plate 450, the regenerated catalyst may enter the reactor 202 between the feed distribution plate 450 and the deactivated catalyst of the first portion.

[0041] In one or more embodiments, the temperature of the feed distribution plate can be from 25°C to 700°C. For example, the temperature of the feed distribution plate can be from 25°C to 600°C, from 25°C to 500°C, from 25°C to 400°C, from 25°C to 300°C, from 25°C to 200°C, from 25°C to 100°C, from 100°C to 700°C, from 100°C to 600°C, from 100°C to 500°C, from 100°C to 400°C, from 100°C to 300°C, from 100°C to 200°C, from 200°C to 700°C, from 200°C to 600°C, from 200°C to 500°C, from 200°C to 400°C, from 200°C to 300°C, from 300°C to 700°C, from 300°C to 600°C, from 300°C to 500°C, from 300°C to 400°C, from 400°C to 700°C, from 400°C to 600°C, from 400°C to 500°C, from 500°C to 700°C, from 500°C to 600°C, from 600°C to 700°C, or any combination of these ranges.

[0042] In one or more embodiments, the deactivated catalyst can be separated into a third portion of deactivated catalyst in the catalyst separation section 214 ( Figure 1 not depicted). In such embodiments, the regenerated catalyst can be mixed with the third portion of deactivated catalyst to form a mixed catalyst before being introduced into the reactor 202, and then passed into the reactor 202. In such embodiments, the first portion of deactivated catalyst can be passed into the reactor 202 downstream of the mixed catalyst relative to the flow direction of the feed stream. Without wishing to be bound by theory, it is believed that mixing the regenerated catalyst with the third portion of deactivated catalyst to form a mixed catalyst may have a lesser negative impact on the activity of the regenerated catalyst due to the lower amount of hydrocarbons entrained in the third portion of deactivated catalyst.

[0043] In one or more embodiments, the mixed catalyst can have a temperature of from 600°C to 850°C. For example, the mixed catalyst can have the following temperatures: from 600°C to 825°C, from 600°C to 800°C, from 600°C to 775°C, from 600°C to 750°C, from 600°C to 725°C, from 600°C to 700°C, from 600°C to 675°C, from 600°C to 650°C, from 600°C to 625°C, from 625°C to 850°C, from 625°C to 825°C, from 625°C to 800°C, from 625°C to 775°C, from 625°C to 750°C, from 625°C to 725°C, from 625°C to 700°C, from 625°C to 675°C, from 625°C to 650°C, from 650°C to 850°C, from 650°C to 825°C, from 650°C to 800°C, from 650°C to 775°C, from 650°C to 750°C, from 650°C to 725°C, from 650°C to 700°C, from 650°C to 675°C, from 675°C to 850°C, from 675°C to 825°C, from 675°C to 800°C, from 675°C to 775°C, from 675°C to 750°C, from 675°C to 725°C, from 675°C to 700°C, from 700°C to 850°C, from 700°C to 825°C, from 700°C to 800°C, from 700°C to 775°C, from 700°C to 750°C, from 700°C to 725°C, from 725°C to 850°C, from 725°C to 825°C, from 725°C to 800°C, from 725°C to 775°C, from 725°C to 750°C, from 750°C to 850°C, from 750°C to 825°C, from 750°C to 800°C, from 750°C to 775°C, from 775°C to 850°C, from 775°C to 825°C, from 775°C to 800°C, from 800°C to 850°C, from 800°C to 825°C, from 825°C to 850°C, or any combination of these ranges.

[0044] Still referring to Figure 1 , in a non-limiting example, the reactor system 103 described herein can be used to produce olefinic compounds from a hydrocarbon feed stream. As used herein, the term "olefinic compound" refers to a hydrocarbon having one or more carbon-carbon double bonds other than the formal double bonds in aromatic compounds. For example, ethylene and styrene are olefinic compounds, but ethylbenzene is not an olefinic compound because the only double bond present in ethylbenzene is a formal double bond that is part of an aromatic structure. Olefinic compounds can be produced from various hydrocarbon feed streams by utilizing different reaction mechanisms. For example, olefinic compounds can be produced by at least dehydrogenation reactions, cracking reactions, dehydration reactions, and methanol-to-olefins reactions. These reaction types can utilize different feed streams and different particulate solids to produce olefinic compounds. It should be understood that when particulate solids are referred to herein, these particulate solids can equally refer to the catalysts mentioned with respect to Figure 1 and Figure 2 the systems of.

[0045] According to one or more embodiments, the reaction can be a dehydrogenation reaction. According to such embodiments, one or more hydrocarbons can be a hydrocarbon feed stream. In one or more embodiments, one or more hydrocarbons can include an alkyl moiety. As used in the present disclosure, a hydrocarbon includes an "alkyl moiety" if the molecule has at least one carbon-carbon single bond capable of dehydrogenating to form a carbon-carbon double bond. The hydrocarbon feed stream can include one or more of ethylbenzene, ethane, propane, n-butane, and isobutane. In one or more embodiments, the hydrocarbon feed stream can include at least 50 wt%, at least 60 wt%, at least 70 wt%, at least 80 wt%, at least 90 wt%, at least 95 wt%, or even at least 99 wt% of ethylbenzene. In one or more embodiments, the hydrocarbon feed stream can include at least 50 wt%, at least 60 wt%, at least 70 wt%, at least 80 wt%, at least 90 wt%, at least 95 wt%, or even at least 99 wt% of ethane. In additional embodiments, the hydrocarbon feed stream can include at least 50 wt%, at least 60 wt%, at least 70 wt%, at least 80 wt%, at least 90 wt%, at least 95 wt%, or even at least 99 wt% of propane. In additional embodiments, the hydrocarbon feed stream can include at least 50 wt%, at least 60 wt%, at least 70 wt%, at least 80 wt%, at least 90 wt%, at least 95 wt%, or even at least 99 wt% of n-butane. In additional embodiments, the hydrocarbon feed stream can include at least 50 wt%, at least 60 wt%, at least 70 wt%, at least 80 wt%, at least 90 wt%, at least 95 wt%, or even at least 99 wt% of isobutane. In additional embodiments, the hydrocarbon feed stream can include at least 50 wt%, at least 60 wt%, at least 70 wt%, at least 80 wt%, at least 90 wt%, at least 95 wt%, or even at least 99 wt% of the sum of ethylbenzene, ethane, propane, n-butane, and isobutane.

[0046] In one or more embodiments, the dehydrogenation reaction can utilize gallium and / or platinum particulate solids as a catalyst. In such embodiments, the particulate solid can include a gallium and / or platinum catalyst. As described herein, the gallium and / or platinum catalyst includes gallium, platinum, or both. The gallium and / or platinum catalyst can be supported on an alumina or alumina-silica support and can optionally include potassium. Such gallium and / or platinum catalysts are disclosed in U.S. Patent No. 8,669,406, the entire text of which is incorporated herein by reference. However, it should be understood that other suitable catalysts can be utilized to conduct the dehydrogenation reaction.

[0047] In one or more embodiments, the reaction mechanism can be dehydrogenation followed by combustion (in the same chamber). In such embodiments, the dehydrogenation reaction can produce hydrogen as a by-product, and the oxygen carrier material can come into contact with the hydrogen and promote its combustion to form water. Examples of such reaction mechanisms are disclosed in WO 2020 / 046978 and U.S. Patent Publication 2021 / 0292259, which are considered to be possible reaction mechanisms for the systems and methods described herein, and the teachings of these documents are incorporated herein by reference in their entirety.

[0048] In one or more embodiments, the particulate solid can comprise an oxygen carrier material and a dehydrogenation catalyst material. In some embodiments, the particulate solid can consist essentially of the oxygen carrier material. As used herein, "consisting essentially of" means a material having less than 1 wt% of unenumerated materials (i.e., consisting essentially of A means that A comprises at least 99 wt% of the composition). In some embodiments, the particulate solid can be free of the dehydrogenation catalyst material. In some embodiments, the oxygen carrier material and the dehydrogenation catalyst material can be separate particles of the particulate solid. In some embodiments, the oxygen carrier material and the dehydrogenation catalyst can be included in the same particle of the particulate solid.

[0049] In embodiments where the particulate solid comprises a dehydrogenation catalyst, the dehydrogenation of one or more hydrocarbons can be effected at least in part by catalytic dehydrogenation. Catalytic dehydrogenation is the dehydrogenation of a hydrocarbon facilitated by the use of a dehydrogenation catalyst. In embodiments where the particulate solid is free of the dehydrogenation catalyst material, the dehydrogenation reaction can be a non-catalytic thermal dehydrogenation reaction. Non-catalytic thermal dehydrogenation refers to the dehydrogenation of a hydrocarbon that occurs due to high temperature, pressure, or a combination thereof without the use of a dehydrogenation catalyst.

[0050] In some embodiments, the particulate solid can comprise a "dual-use material" that can serve as both a dehydrogenation catalyst and an oxygen carrier material. It should be understood that in at least the embodiments described herein where the oxygen carrier material and the dehydrogenation catalyst are used in the same reaction vessel (such as Figure 1 the reaction vessel), such dual-use materials can be used in place of or in combination with the oxygen carrier material of the particulate solid and the dehydrogenation catalyst of the particulate solid.

[0051] In one or more embodiments, the particulate solid may be capable of fluidizing. In some embodiments, the particulate solid may exhibit characteristics that are industrially referred to as "Geldart A" or "Geldart B" characteristics. According to D. Geldart, Gas Fluidization Technology, John Wiley & Sons (New York, 1986), pages 34 - 37; and D. Geldart, "Types of Gas Fluidization", Powder Technol. 7 (1973) 285 - 292, which are incorporated herein by reference in their entirety, the particles can be classified as "Group A" or "Group B".

[0052] Group A is understood by those skilled in the art to represent aeratable powders, having a fluidization with a range of no bubbles; high bed expansion; slow and linear rate of de - aeration; bubble characteristics, which may include the advantage of splitting / re - coalescing bubbles, having a maximum bubble size and a large wake; a high level of solid mixing and gas back - mixing, assuming U - Umf is equal (U is the velocity of the carrier gas and Umf is the minimum fluidization velocity, usually but not necessarily measured in meters per second (m / s), i.e., there is an excessive gas velocity); axisymmetric slug characteristics; and no spouting except in very shallow beds. The listed characteristics tend to improve with decreasing mean particle size, assuming cfp is equal; or with an increasing proportion of <45 micrometers (μm); or with increasing pressure, temperature, viscosity, and density of the gas. Generally, the particles may exhibit a small mean particle size and / or a low particle density (<1.4 grams per cubic centimeter, g / cm 3 ); are easily fluidized, where they fluidize smoothly at low gas velocities; and may exhibit controlled bubbling with small bubbles at higher gas velocities.

[0053] Group B is understood by those skilled in the art to represent "sand - like" powders, which start to bubble at Umf; which exhibit moderate bed expansion; rapid de - aeration; no limitation on bubble size; a moderate level of solid mixing and gas back - mixing, assuming U - umf is equal; both axisymmetric and non - axisymmetric slugs; and spouting only in shallow beds. These characteristics tend to improve with decreasing mean particle size, but the particle size distribution and a certain uncertainty in the gas, pressure, temperature, viscosity, or density do not seem to play much of a role in improving these characteristics. Generally, when the density (pp) is 1.4 < pp < 4 g / cm 3 , the particle size (cfp) of most particles is 40 μm < cfp < 500 μm, and preferably, when the density (pp) is 4 g / cm 3When, the particle size of most particles is 60 μm <cfp <500 μm, and when the density (pp) is 1 g / cm 3 When, the particle size of most particles is 250 μm <cfp <100 μm.

[0054] In one or more embodiments, the olefin compound may be present in a "product stream" sometimes referred to as an "olefin-containing effluent". Such streams leave Figure 1 the reactor system and may subsequently be processed. In one or more embodiments, the olefin compound may include one or more of ethylene, propylene, butene, or styrene. The term butene includes any butene isomer, such as α-butene, cis-β-butene, trans-β-butene, and isobutene. In some embodiments, the olefin-containing effluent may contain at least 20 wt%, at least 30 wt%, at least 40 wt%, at least 50 wt%, or even at least 60 wt% ethylene. In additional embodiments, the olefin-containing effluent may contain at least 20 wt%, at least 30 wt%, at least 40 wt%, at least 50 wt%, or even at least 60 wt% propylene. In additional embodiments, the olefin-containing effluent may contain at least 20 wt%, at least 30 wt%, at least 40 wt%, at least 50 wt%, or even at least 60 wt% butene. In additional embodiments, the olefin-containing effluent may contain at least 20 wt%, at least 30 wt%, at least 40 wt%, at least 50 wt%, or even at least 60 wt% styrene. In additional embodiments, the olefin-containing effluent may contain at least 20 wt%, at least 30 wt%, at least 40 wt%, at least 50 wt%, or even at least 60 wt% of the sum of one or more of ethylene, propylene, butene, and styrene. The olefin-containing effluent may also contain unreacted components of the feed stream and other reaction products not considered light olefins. The olefin compound may be separated from the unreacted components in a subsequent separation step.

[0055] Example

[0056] Example 1 - Catalyst Deactivation from CH4 and Steam

[0057] In Example 1, the effects of CH4 exposure and steam exposure on catalyst activity were observed. A 0.5 g supported gallium catalyst loaded with a platinum promoter was tested in a fixed-bed apparatus. A laboratory-simulated reaction-combustion-reactivation cycle was carried out in the fixed-bed apparatus. In each cycle, a dehydrogenation reaction was first carried out at 625 °C for 60 seconds, where the weight hourly space velocity "WHSV" of propane was 10 h -1, and the feed composition is 90% propane / 10% nitrogen; then the catalyst is treated with a simulated combustion stream at 750 °C for 3 minutes; finally, the catalyst is reactivated in air at 750 °C for 15 minutes. The dehydrogenation performance is collected at 15 seconds of operation. To test the effect of CH4 exposure on the catalyst activity, after reactivation in air, the catalyst is treated with 100% CH4 at a flow rate of 10 standard cubic centimeters per minute (sccm) at 750 °C for 2 minutes. Then the cycle is run again, and the dehydrogenation performance of the CH4-treated catalyst is collected at 15 seconds of operation. Then, ten laboratory-simulated reaction-combustion-reactivation cycles are run without CH4 treatment to restore the catalyst activity to the baseline. The catalyst is treated with 100% CH4 at a lower temperature of 625 °C for 2 minutes. Then the cycle is run again, and the dehydrogenation performance of the catalyst treated with low-temperature CH4 is collected at 15 seconds of operation.

[0058] Then, ten laboratory-simulated reaction-combustion-reactivation cycles are run without CH4 treatment to restore the catalyst activity to the baseline. Then the catalyst is treated with steam at a flow rate of 24 sccm at 625 °C for 2 minutes, followed by stripping with helium for 5 minutes, and then the catalyst is run in the laboratory-simulated reaction-combustion-reactivation cycle. The dehydrogenation performance of the steam-treated catalyst is collected at 15 seconds of operation. The dehydrogenation performance of the catalyst under various treatment conditions is recorded in Table 1.

[0059] Table 1

[0060]

[0061] As indicated in Table 1, exposing the catalyst to steam or CH4 at high temperature (750 °C) before using it in the dehydrogenation reaction has a negative impact on the propane conversion and propane selectivity performance of the catalyst, as well as on the intrinsic rate of the catalyst. As shown by sample B1, when compared to a catalyst that has not been exposed to CH4 before being used in the dehydrogenation reaction (i.e., sample A), exposure to CH4 results in a loss of approximately 40% in propane conversion performance, a 2.4% loss in propane selectivity, and a 68% loss in activity. Similarly, sample C shows that when compared to a catalyst that has not been exposed to steam before being used in the dehydrogenation reaction (i.e., sample A), exposing the catalyst to steam before use in the dehydrogenation reaction results in a loss of approximately 60% in propane conversion performance, an 8% loss in propane selectivity, and an 86% loss in activity. This indicates that the premixing of a high-temperature regenerated catalyst carrying oxygen with a low-temperature recycled deactivated catalyst carrying a stripping hydrocarbon gas (such as methane) may lead to the unexpected deactivation of the regenerated catalyst. In contrast, exposing the catalyst to CH4 at low temperature (625 °C) before using it in the dehydrogenation reaction has only a minor impact on the propane conversion and propane selectivity performance of the catalyst, as well as on the intrinsic rate of the catalyst.

[0062] Example 2 - Influence of Catalyst Dispenser Configuration

[0063] To simulate the catalyst mixing and dehydrogenation reaction in a typical fluidized catalytic dehydrogenation reactor, as shown in Figure 3, a reactor model consisting of three continuously stirred tank reactors (CSTRs) in series is used, where in each CSTR, the catalyst and reactants are well mixed. In Figure 3, the three CSTRs are shown as part of a single reactor system 600. The catalyst and / or reactants flow from the first CSTR 610 to the second CSTR 620 and then to the third CSTR 630. For the case of premixing where there is regenerated catalyst 640 and part 1 deactivated catalyst 606 upstream of the reactor, it is represented by a mixer (mixing tank) 640. The kinetic models of Lobera et al. (2008) and Sundaram and Froment (1977) are used to solve the catalytic dehydrogenation and thermal reactions. The mass and energy conservation equations for the mixer and each CSTR are solved to determine the composition and temperature of the effluent gas. Fifty percent of the catalyst from the third CSTR 630 is recycled as recycled catalyst 606 to be used again in the reactor simulation without regeneration. Due to the net amount of propane reacting in each CSTR, the catalyst will continuously deactivate as it moves from one CSTR to another. The same total amount of catalyst in each CSTR is used when simulating each of the reactor configurations shown below. The reactor model uses the same regeneration process for each configuration, resulting in the same catalyst activity of the regenerated catalyst 604 with adsorbed oxygen. The temperature of the regenerated catalyst is 750 °C. The model also uses CH4 as the stripping gas in all configurations, which causes some CH4 to be adsorbed by the recycled catalyst. The temperature of the recycled catalyst is predicted to be the effluent temperature of the third CSTR.

[0064] The above reactor model is used to evaluate three reactor configurations. In Configuration A, as Figure 3A shown, the regenerated catalyst 604 carrying adsorbed oxygen is mixed with the recycled catalyst 606 from the third CSTR 630 in a mixing tank 640 upstream of the first CSTR 610. Then the mixed catalyst 642 enters the first CSTR 610, where propane 602 is fed into the reactor 600. Based on the experimental data presented in Example 1, it is estimated that the adsorbed CH4 in the recycled catalyst 606 reacts with the residual oxygen in the regenerated catalyst 604 at the mixing tank 640 temperature of 685 °C, resulting in 40% activity. Configuration A approximates the reactor conditions that would occur in the case where the recycled catalyst 606 and the regenerated catalyst 604 are premixed in the mixing tank 640 before entering the reactor 600.

[0065] In Configuration B, as Figure 3BAs shown, the recycled catalyst 606 from the third CSTR 630 is fed into the first CSTR 610, propane 602 is also fed into the first CSTR 610 separately, and the regenerated catalyst 604 is fed into the second CSTR 620. Thus, the recycled catalyst 606 is transferred from the third CSTR 630 to the first CSTR 610 and then to the second CSTR 620 before being mixed with the regenerated catalyst 604 fed into the second CSTR 620. Due to the endothermic dehydrogenation reaction, the process temperature decreases after the series of CSTRs. The mixing of the regenerated catalyst 604 carrying adsorbed oxygen and the recycled catalyst 606 with adsorbed CH4 takes place in the second CSTR 620, where the temperature is expected to be 623.5 °C. As shown in Example 1, the effect of CH4 on the catalyst activity is significantly reduced at lower temperatures and is then negligible for Configuration B.

[0066] In Configuration 1, as Figure 3C shown, the regenerated catalyst 604 is fed into the first CSTR 610, propane 602 is also fed into the first CSTR 610 separately, and the recycled catalyst 606 from the third CSTR 630 is fed into the second CSTR 620 separately. Thus, the regenerated catalyst 604 is transferred from the first CSTR 610 to the second CSTR 620 before being mixed with the recycled catalyst 606 in the second CSTR 620. It is estimated that the adsorbed oxygen in the regenerated catalyst 604 is consumed in the first CSTR 610 before being mixed with the recycled catalyst 606 having adsorbed CH4. Thus, the deactivation of CH4 in the catalyst dehydrogenation activity is negligible. Configuration 1 represents the reactor conditions that would occur when the regenerated catalyst 604 is fed into the reactor 600 at a position between the propane feed 602 and the recycled catalyst 604. The results of the runs from Configurations A, B, and 1 are recorded in Table 2.

[0067] Table 2

[0068]

[0069]

[0070] As indicated in Table 2, Configuration 1 based on the present invention has the highest total propane conversion (36.3%) and propylene yield (32.7%) among the three configurations. The next closest propane conversion is Configuration B, which has a propane conversion and propylene yield that are 2.4% and 1.3% lower than those of Configuration 1, respectively, indicating that Configuration 1 has significantly better propane conversion compared to the other test configurations.

[0071] In a first aspect of the present disclosure, one or more olefin compounds can be produced by a method that includes dehydrogenating a feed stream in a reactor in the presence of a catalyst to form a product stream and a deactivated catalyst. The feed stream includes one or more hydrocarbons containing an alkyl moiety, and the product stream includes one or more olefin compounds. The method further includes separating the deactivated catalyst into a first portion of the deactivated catalyst and a second portion of the deactivated catalyst. The method further includes passing the second portion of the deactivated catalyst to a regenerator and treating the second portion of the deactivated catalyst in the regenerator to form a regenerated catalyst. The method further includes passing the first portion of the deactivated catalyst and the regenerated catalyst to the reactor. The first portion of the deactivated catalyst enters the reactor downstream of the regenerated catalyst with respect to the flow direction of the feed stream. The first portion of the deactivated catalyst has a lower temperature than the regenerated catalyst.

[0072] A second aspect of the present disclosure includes any of the foregoing aspects or combinations of aspects, wherein the first portion of the deactivated catalyst and the regenerated catalyst are passed to the reactor through a granular solid distributor that passes the first portion of the deactivated catalyst and the regenerated catalyst separately into the reactor.

[0073] A third aspect of the present disclosure includes any of the foregoing aspects or combinations of aspects, wherein the granular solid distributor extends into the reactor through the bottom end of the reactor.

[0074] A fourth aspect of the present disclosure includes any of the foregoing aspects or combinations of aspects, wherein the one or more hydrocarbons include propane and the one or more olefin compounds include propylene.

[0075] A fifth aspect of the present disclosure includes any of the foregoing aspects or combinations of aspects, wherein the catalyst includes one or more of gallium or platinum.

[0076] A sixth aspect of the present disclosure includes any of the foregoing aspects or combinations of aspects, wherein the reactor operates as a fast fluidized, turbulent, or bubbling fluidized bed reactor.

[0077] A seventh aspect of the present disclosure includes any of the foregoing aspects or combinations of aspects, wherein the temperature of the first portion of the deactivated catalyst passed to the reactor is from 580 °C to 800 °C.

[0078] An eighth aspect of the present disclosure includes any of the foregoing aspects or combinations of aspects, wherein the temperature of the regenerated catalyst passed to the reactor is from 680 °C to 900 °C.

[0079] A ninth aspect of the present disclosure includes any of the foregoing aspects or combinations of aspects, wherein hydrocarbons are entrained in the first portion of the deactivated catalyst and the second portion of the deactivated catalyst.

[0080] The tenth aspect of the present disclosure includes any of the foregoing aspects or combinations of aspects, wherein the one or more hydrocarbons include propane, and the one or more olefin compounds include propylene. The catalyst includes one or more of gallium and platinum. The temperature of the deactivated catalyst transferred to the first part of the reactor is 580 °C to 800 °C. The temperature of the regenerated catalyst transferred to the reactor is 680 °C to 900 °C. The deactivated catalyst of the first part and the regenerated catalyst are transferred to the reactor through a granular solid distributor, and the granular solid distributor transfers the deactivated catalyst of the first part and the regenerated catalyst separately into the reactor. The granular solid distributor extends into the reactor through the bottom of the reactor. The reactor operates as a fast fluidized, turbulent or bubbling fluidized bed reactor.

[0081] The eleventh aspect of the present disclosure includes any of the foregoing aspects or combinations of aspects, wherein the reactor includes a feed distribution plate, the regenerated catalyst enters the reactor between the feed distribution plate and the deactivated catalyst of the first part, and the temperature of the feed distribution plate is 25 °C to 700 °C.

[0082] The twelfth aspect of the present disclosure includes any of the foregoing aspects or combinations of aspects, wherein the deactivated catalyst of the first part and the regenerated catalyst constitute at least 95% by weight of the catalyst transferred to the reactor.

[0083] The thirteenth aspect of the present disclosure includes any of the foregoing aspects or combinations of aspects, wherein the mass flow rate ratio of the deactivated catalyst of the first part to the regenerated catalyst is 0.1 to 5.

[0084] The fourteenth aspect of the present disclosure includes any of the foregoing aspects or combinations of aspects, wherein in addition to the deactivated catalyst of the first part and the deactivated catalyst of the second part, the deactivated catalyst is separated into a third part. Before being transferred to the reactor, the deactivated catalyst of the third part is combined with the regenerated catalyst to form a mixed catalyst.

[0085] The fifteenth aspect of the present disclosure includes any of the foregoing aspects or combinations of aspects, wherein the mixed catalyst has a temperature of 600 °C to 850 °C.

[0086] It will be apparent to those skilled in the art that various modifications and variations can be made to the technology disclosed in this invention without departing from the spirit and scope of the technology. Since those skilled in the art can make modifications, combinations, sub - combinations, and variations of the disclosed embodiments that incorporate the spirit and substance of the technology disclosed in this invention, this technology should be construed to include all things within the scope of the appended claims and their equivalents. Additionally, although some aspects of this disclosure may be identified herein as preferred or particularly advantageous, this disclosure is not limited to these aspects.

[0087] It should be noted that the various details described in this disclosure should not be taken to imply that these details relate to elements that are fundamental components of each of the embodiments described in this disclosure, even in cases where a particular element is shown in each of the drawings appended to this specification. Unless expressly stated otherwise, none of the features disclosed and described herein should be construed as "essential". The contemplated embodiments of this technology include those that include some or all of the features of the appended claims.

[0088] For the purposes of describing and defining this disclosure, it should be noted that the term "about" is used in this disclosure to denote the inherent uncertainty that can be attributed to any quantitative comparison, value, measurement, or other representation. The term "about" is also utilized in this disclosure to denote the degree to which a quantitative representation can vary from a stated reference without resulting in a change in the basic function of the subject matter of concern.

[0089] In relevant cases, where a composition is described as "comprising" one or more elements, embodiments of the composition "consisting of" or "consisting essentially of" those one or more elements are contemplated herein.

[0090] It should be understood that in some embodiments, the compositional ranges of chemical components in a stream or reactor should be understood to contain mixtures of isomers of that component. For example, a specified compositional range of butene can include mixtures of the various isomers of butene. It should be understood that the examples provide compositional ranges for various streams, and the total amount of isomers of a particular chemical composition can constitute a range.

[0091] It should be noted that one or more of the following claims and detailed description utilize the terms "where" or "wherein" as transitional phrases. For the purposes of defining this technology, it should be noted that this term is introduced in the claims as an open - ended transitional phrase that is used to introduce a recitation of a series of characteristics of a structure and should be interpreted in a manner similar to the more commonly used open - ended preamble term "comprising".

[0092] It should be understood that any two quantitative values assigned to a characteristic can form a range of the characteristic, and all combinations of ranges formed by all the said quantitative values of a given characteristic are contemplated in the present disclosure. In cases where multiple ranges of quantitative values are provided, these ranges can be combined to form a wider range, which is contemplated in the embodiments described herein.

[0093] As understood in the context of the terms used herein, the term "transfer" can include directly transferring a substance between two parts of the disclosed system, and in some cases, means indirectly transferring a substance between two parts of the disclosed system. For example, indirect transfer can include steps in which the substance transfer passes through an intermediate operating unit, valve, sensor, etc.

Claims

1. A method for producing one or more olefin compounds, the method comprising: In a reactor, a feed stream is dehydrogenated in the presence of a catalyst to form a product stream and a deactivated catalyst, wherein the feed stream comprises one or more hydrocarbons containing an alkyl moiety, and the product stream comprises one or more olefin compounds; The deactivated catalyst is separated into a first portion of the deactivated catalyst and a second portion of the deactivated catalyst; The second portion of the deactivated catalyst is transferred to a regenerator; The second portion of the deactivated catalyst is treated in the regenerator to form a regenerated catalyst; And The first portion of the deactivated catalyst and the regenerated catalyst are transferred to the reactor, wherein the first portion of the deactivated catalyst enters the reactor downstream of the regenerated catalyst with respect to the flow direction of the feed stream, and wherein the first portion of the deactivated catalyst has a lower temperature than the regenerated catalyst.

2. The method according to claim 1, wherein the deactivated catalyst of the first part and the regenerated catalyst are transferred to the reactor through a granular solid distributor, and the granular solid distributor transfers the deactivated catalyst of the first part and the regenerated catalyst to the reactor separately.

3. The method according to claim 2, wherein the granular solid distributor extends into the reactor through the bottom end of the reactor.

4. The method according to any one of the preceding claims, wherein the one or more hydrocarbons comprise propane, and the one or more olefin compounds comprise propylene.

5. The method according to any one of the preceding claims, wherein the catalyst comprises one or more of gallium and platinum.

6. The method according to any one of the preceding claims, wherein the reactor is operated as a fast fluidized, turbulent or bubbling fluidized bed reactor.

7. The method according to any one of the preceding claims, wherein the temperature of the first part of the deactivated catalyst transferred to the reactor is 580 °C to 800 °C.

8. The method according to any one of the preceding claims, wherein the temperature of the regenerated catalyst transferred to the reactor is 680 °C to 900 °C.

9. The method according to any one of the preceding claims, wherein hydrocarbons are entrained in the first part of the deactivated catalyst and the second part of the deactivated catalyst.

10. The method according to claim 1, wherein: The one or more hydrocarbons include propane, and the one or more olefin compounds include propylene; The catalyst comprises one or more of gallium and platinum; The temperature of the first portion of the deactivated catalyst transferred to the reactor is 580 °C to 800 °C; The temperature of the regenerated catalyst transferred to the reactor is 680 °C to 900 °C; and Wherein the first portion of the deactivated catalyst and the regenerated catalyst are transferred to the reactor by a granular solid distributor that transfers the first portion of the deactivated catalyst and the regenerated catalyst separately into the reactor; The granular solid distributor extends into the reactor through the bottom of the reactor; and The reactor operates as a fast fluidized, turbulent or bubbling fluidized bed reactor.

11. The method according to any one of the preceding claims, wherein the reactor comprises a feed distribution plate, the regenerated catalyst enters the reactor between the feed distribution plate and the deactivated catalyst of the first part, and the temperature of the feed distribution plate is from 25 °C to 700 °C.

12. The method according to any one of the preceding claims, wherein the deactivated catalyst of the first part and the regenerated catalyst constitute at least 95% by weight of the catalyst transferred to the reactor.

13. The method according to any one of the preceding claims, wherein the mass flow rate ratio of the deactivated catalyst of the first part to the regenerated catalyst is from 0.1 to 5.

14. The method according to any one of the preceding claims, wherein in addition to the deactivated catalyst of the first part and the deactivated catalyst of the second part, the deactivated catalyst is separated into a third part, and wherein before being transferred to the reactor, the deactivated catalyst of the third part is combined with the regenerated catalyst to form a mixed catalyst.

15. The method according to claim 14, wherein the mixed catalyst has a temperature of from 600 °C to 850 °C.

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