CN117999252A

CN117999252A - Method and system for producing light olefins in high yields and selectivities

Info

Publication number: CN117999252A
Application number: CN202280059827.2A
Authority: CN
Inventors: S·S·马杜斯卡尔; K·H·库克莱尔; 鲍筱颖
Original assignee: ExxonMobil Chemical Patents Inc
Current assignee: ExxonMobil Chemical Patents Inc
Priority date: 2022-04-08
Filing date: 2022-10-10
Publication date: 2024-05-07
Also published as: KR20240024911A

Abstract

A process for forming an olefin, the process comprising: introducing a hydrocarbon feedstream to a reactor comprising a dehydrogenation catalyst; reacting the hydrocarbon feedstream with a dehydrogenation catalyst in a reactor to form a high temperature dehydrogenation product comprising at least a portion of the dehydrogenation catalyst; separating at least a portion of the dehydrogenation catalyst from the high temperature dehydrogenation product in a primary separation device and a secondary separation device downstream of and in fluid communication with the primary separation device; combining the high temperature dehydrogenation product with a quench stream after the high temperature dehydrogenation product exits the secondary separation device to cool the high temperature dehydrogenation product and form an intermediate temperature dehydrogenation product, wherein the quench stream comprises hydrocarbons; and cooling the intermediate temperature dehydrogenation product to form a cooled dehydrogenation product.

Description

Method and system for producing light olefins in high yields and selectivities

Cross Reference to Related Applications

The present application claims priority and benefit from U.S. provisional application No. 63/222,733, having a filing date of 2021, 7, 16, and U.S. provisional application No. 63/329,006, having a filing date of 2022, 4, 8, the disclosures of which are all incorporated herein by reference.

Technical Field

Exemplary embodiments of the present technology advancement relate to chemical treatment systems, and more particularly, to dehydrogenation chemical treatment systems.

Background

Light olefins can be used as a base material to produce many types of products and materials. For example, ethylene may be used to make polyethylene, vinyl chloride or ethylene oxide. Such products may be used in product packaging, construction, textiles, and the like. Thus, there is a need in the industry for light olefins such as ethylene, propylene, and butenes. Light olefins may be produced by different reaction processes depending on a given chemical feed, which may be a product from a crude oil refining operation or a renewable product from a biorefinery or natural gas component stream. Many light olefins may be produced by catalytic processes, such as catalytic dehydrogenation, wherein the feed stream is contacted with a fluidized catalyst that facilitates conversion of the feed stream to light olefins. In such systems, reaction selectivity to light olefins can be important to overall process efficiency.

Patent publication WO 2020/263599, which is incorporated herein by reference in its entirety, describes a process for forming light olefins, the process comprising introducing a hydrocarbon feed stream into a reactor, reacting the hydrocarbon feed stream with a dehydrogenation catalyst in the reactor to form a high temperature dehydrogenation product, separating at least a portion of the dehydrogenation catalyst from the high temperature dehydrogenation product in a primary separation device, combining the high temperature dehydrogenation product with a quench stream to cool the high temperature dehydrogenation product and form an intermediate temperature dehydrogenation product, and cooling the intermediate temperature dehydrogenation product to form a cooled dehydrogenation product.

Patent publication WO 2018/236630, which is incorporated herein by reference in its entirety, describes a method of treating a chemical stream, which method comprises contacting a feed stream with a catalyst in an upstream reactor section of a reactor having an upstream reactor section and a downstream reactor section, passing an intermediate product stream into the downstream reactor section, and introducing a riser quench fluid into the downstream reactor section, the upstream reactor section or the transition section and into contact with the intermediate product stream and the catalyst to slow or stop the reaction. The process includes separating at least a portion of the catalyst from the product stream, passing the product stream to a product treatment section, cooling the product stream and separating a portion of the riser quench fluid from the product stream. The riser quench fluid separated from the product stream can be recycled back to the downstream reactor section, the upstream reactor section, or the transition section as riser quench fluid.

Disclosure of Invention

A system for forming olefins from a hydrocarbon feed stream, the system comprising: a reactor configured to receive a hydrocarbon feedstream and a dehydrogenation catalyst under reaction conditions, wherein the reaction conditions produce a high temperature dehydrogenation product comprising at least a portion of the dehydrogenation catalyst; a plurality of separation devices configured to separate at least a portion of the dehydrogenation catalyst from the high temperature dehydrogenation product, the plurality of separation devices comprising a primary separation device and a secondary separation device downstream of and in fluid communication with the primary separation device; a quench system configured to combine the high temperature dehydrogenation product with a quench stream to cool the high temperature dehydrogenation product and form an intermediate temperature dehydrogenation product, wherein the quench stream comprises hydrocarbons; and a heat transfer system configured to cool the intermediate temperature dehydrogenation product to form a cooled dehydrogenation product.

Brief description of the drawings

FIG. 1 is a graph illustrating the yield loss of olefins per second due to secondary thermal cracking at different temperatures.

FIG. 2 is a graph illustrating the selectivity loss of olefins per second due to secondary thermal cracking at different temperatures.

FIG. 3 illustrates an exemplary embodiment of a generalized flow diagram of a reactor system with a recycle quench stream.

Fig. 4 illustrates a more detailed example of an exemplary reactor embodying the advances in the technology.

Fig. 5 is a graph illustrating PDH performance of a typical PDH catalyst associated with advances in the art.

FIG. 6 is a graph illustrating the yield of C ₃H₆ of product A versus reaction zone temperature.

FIG. 7 is a graph illustrating C ₃H₆ selectivity to product A versus reaction zone temperature.

Figure 8 shows that the catalyst composition (catalyst 8) maintains its performance for 204 cycles.

Detailed description of the preferred embodiments

Definition of the definition

As used herein, and unless otherwise specified, the term "hydrocarbon" means a class of compounds containing carbon-bonded hydrogen and encompasses mixtures of (i) saturated hydrocarbon compounds, (ii) unsaturated hydrocarbon compounds, and (iii) hydrocarbon compounds (saturated and/or unsaturated), including mixtures of hydrocarbon compounds having different n values.

As used herein, "unformed" means the formation of a product in the reactor that is not greater than 0.1 mole% carbon selectivity (as compared to the propylene feed), where (moles of carbon in the product)/(moles of carbon converted in the feed) =mole% carbon selectivity of the product. For example, if 2 moles of ethylene are produced when 3 moles of propane are converted, the carbon mole% selectivity of ethylene is 44.4%.

An "olefin" or "olefin" is a linear, branched or cyclic compound of carbon and hydrogen having at least one double bond.

"Distribution chamber" means the region of the reactor that facilitates fluid communication between the piping or tubing that carries the hot product stream from the reactor to the outlet stream. The distribution chamber may also act as a conduit for collecting gas from the sets of cyclones before it is discharged from the reactor. As described below, the reactor may have a plurality of distribution chambers (e.g., a first distribution chamber and a second distribution chamber as illustrated in fig. 4) and the term distribution chamber will refer to either the first or second distribution chamber unless otherwise indicated.

For the purposes of this disclosure, the nomenclature of the elements is according to the version of the periodic table of the elements (according to the new notation) as described in Hawley' sCondensed Chemical Dictionary, 16 th edition, john Wiley & Sons, inc., (2016), appendix V. For example, the group 2 element includes Mg, the group 8 element includes Fe, the group 9 element includes Co, the group 10 element includes Ni and the group 13 element includes Al. The term "metalloid" as used herein refers to the following elements: B. si, ge, as, sb, te and At. In the present disclosure, when a given element is indicated as being present, it may be present in elemental state or as any compound thereof, unless otherwise specified or clear from context.

The term "mixed metal oxide" refers to a composition comprising oxygen atoms and at least two different metal atoms mixed on an atomic scale. For example, a "mixed Mg/Al metal oxide" has O, mg and Al atoms mixed on an atomic scale and is substantially the same or identical to a composition obtained by calcining Mg/Al hydrotalcite having the general chemical formulaWherein A is a counter anion having a negative charge n, x ranges from > 0 to <1, and m is ≡0. The material consisting of nm-sized MgO particles and nm-sized Al ₂O₃ particles mixed together is not a mixed metal oxide, because Mg and Al atoms are not mixed on an atomic scale but instead on a nm scale.

The term "selectivity" refers to the productivity (based on moles of carbon) of a given compound in a catalytic reaction. By way of example, the phrase "alkane conversion reaction has 100% selectivity to alkene" means that 100% of the alkane (on a carbon mole basis) converted in the reaction is converted to alkene. The term "conversion" when used in connection with a given reactant means the amount of reactant consumed in the reaction. For example, when the reactant is specified to be propane, 100% conversion means 100% of the propane is consumed in the reaction. In another example, when the reactant is specified to be propane, if 1 mole of propane is converted to 1 mole of methane and 1 mole of ethylene, then the methane selectivity is 33.3% and the ethylene selectivity is 66.7%. Yield (on a carbon mole basis) is conversion times selectivity.

For the sake of brevity, only certain ranges are explicitly disclosed herein. However, a range starting from any lower limit may be combined with any upper limit to thereby describe a range not explicitly described, and a range starting from any lower limit may be combined with any other lower limit to thereby describe a range not explicitly described, and a range starting from any upper limit may be combined with any other upper limit in the same manner to thereby describe a range not explicitly described. In addition, each point or individual value between its endpoints is included within the range even though not explicitly recited. Thus, each point or individual value may serve as its own lower or upper limit, combined with any other point or individual value or any other lower or upper limit, thereby recitation of ranges not explicitly recited.

Although terms such as hot, high temperature, and intermediate temperature are used herein to describe the reactor and the various streams, these terms of extent are used to distinguish the streams based on relative temperature characteristics. Those of ordinary skill in the art will understand that no specific temperature is implied and will understand the relative temperatures of the streams with respect to reactor temperature and the thermodynamically determined heat flow.

Exemplary embodiments

Light olefins (e.g., propylene) may be produced by catalytic processes such as dehydrogenation, wherein a feed stream (e.g., propane) is contacted with a fluidized dehydrogenation catalyst at an elevated temperature. The feed stream may be derived from crude oil refining (e.g., propane). The resulting product mixture may contain a mixture of highly reactive molecules and catalyst particles. Exemplary embodiments of the present technology advancement describe reactor configurations that can separate catalyst particles from a hot product stream (e.g., about reactor temperature) and then cool the hot product stream using a direct liquid quench to avoid selective losses of secondary heat and catalytic reactions.

The selectivity of the light olefins from the dehydrogenation reactor is adversely affected by the secondary reaction. Secondary reactions include continuous thermal and catalytic reactions of reactants through side reactions and/or excessive reactions of products. To avoid this, the catalyst particles should be separated from the hot product stream and the resulting hot product stream needs to be cooled to an intermediate temperature low enough to slow or stop the thermal reaction.

Yield and selectivity losses after propylene production through a Propane Dehydrogenation (PDH) reactor can occur due to both thermal and catalytic reactions. These secondary reactions can be minimized by efficient separation of catalyst particles from the high temperature product stream and immediate quenching of the high temperature product stream after catalyst separation. The preferred sequence is to use primary and secondary cyclones to approach complete catalyst separation (i.e., about >98% separation) followed by hydrocarbon quench. This can be contrasted with the conventional approach of intermediate quench and partial catalyst separation as described in WO 2020/263599.

A dehydrogenation reactor configured in accordance with the present technological advancements may include some or all of the following features: (a) Removing catalyst particles from the product stream immediately prior to cooling after the riser by using primary and secondary cyclones; (b) Cooling the high temperature product stream after catalyst separation using a direct liquid/gas quench; (c) Using a quench fluid, which is a gas or saturated or subcooled liquid, which gasifies/flashes upon contact with the high temperature product stream fluid, which provides for efficient contact for heat transfer; (d) Separating the product stream in a downstream separation device in which at least a portion of the byproduct stream (e.g., benzene) from the separator is recycled as a quench fluid; (e) The quench fluid may be either derived from the product stream or it may be an external fluid (e.g., a heavy aromatic solvent) that is not formed in the reactor, or it may be water; (f) the reaction temperature ranges from 600 ℃ to 700 ℃; (g) A Pt catalyst supported on a mixed magnesium aluminum oxide was used; (h) Integrated with the steam cracker heat and/or recovery process; (i) And removing fines from the product stream after the initial quench.

Reactor systems embodying advances in the technology may provide some or all of the following advantages: (a) The primary and secondary cyclones prior to quenching can achieve very high catalyst separation efficiency (> 98%) from the high temperature product stream, which minimizes secondary catalytic reactions that can occur even after quenching; (b) The separated catalyst particles may be recycled to the regenerator at 800 ℃, wherein the higher separation efficiency prior to quenching avoids significant heat loss from the catalyst particles; and (c) injecting quench fluid into the distribution chamber of the reactor reduces reactor design complexity relative to conventional systems of WO 2020/263599 in which multiple quench lines would be required per separation group.

Experiments were performed to investigate the effect of secondary heat and catalytic reactions in a laboratory fixed bed reactor on propylene (C ₃H₆ yield). Experiments were performed in a tubular reactor without catalyst charge to measure yield loss due to secondary thermal cracking. FIG. 1 is a graph illustrating propylene yield loss due to secondary thermal cracking per post-reactor residence time at two different pressures at reactor outlet temperature. For example, a post-reactor residence time of 5 seconds at 26psia at 640 ℃ will result in a 2.5 carbon mole% propylene yield loss. FIG. 2 is a graph illustrating propylene selectivity loss due to secondary thermal cracking per post-reactor residence time at two different pressures at reactor outlet temperature. This loss in yield and selectivity can be attributed to the secondary thermal cracking reaction of unconverted propane to non-propylene products and the thermal cracking reaction of propylene itself. This loss of yield and selectivity can have a significant impact on the overall economics of propylene production.

FIG. 3 illustrates an exemplary embodiment of a generalized flow diagram of a reactor system with a recycle quench stream. It should be noted that many of the method and apparatus details of one or more embodiments of the system of fig. 3 are described herein with reference to fig. 4.

Still referring to fig. 3, the reactor system 102 may include a reactor 202, a separation device 220, and a distribution chamber section 204. The reactor system is connected to a catalyst treatment section 300. Typically, a primary reaction, such as a dehydrogenation reaction, occurs in the reactor 202, wherein a reactant stream 280 (sometimes referred to herein as a hydrocarbon feed stream) from outside the described system is combined with regenerated catalyst from stream 282 of the catalyst treatment section and passed into the reactor 202. After the reaction, the catalyst, unreacted chemicals, and product chemicals are transferred to separation device 220 via stream 284. It is to be understood that the "product stream" can include reaction products and unreacted components from the reactant stream 280. Reactant stream 280 may comprise one or more of propane, n-butane, isobutane, ethane, or ethylbenzene.

Stream 284 is sometimes referred to herein as a high temperature dehydrogenation product. Typically, stream 284 has a temperature that is approximately equal to the temperature at the outlet of reactor 202. The temperature may depend on the reaction and the catalyst system used. In one or more embodiments, for propane dehydrogenation, stream 284 has a temperature in the range of 600 ℃ to 700 ℃ and a pressure in the range of 1.0bara to 4.0 bara.

The separation device 220 may include a primary separation device 402 and a secondary separation device 404. In further embodiments, the primary separation device 402 and/or the secondary separation device 404 may be, without limitation, a cyclone, a filter, or other suitable device for separating solids, such as catalyst, from a gas. After separating at least a portion (typically a majority) of the catalyst of stream 284 from the vapor phase reactants and product chemicals in primary separation device 402, catalyst passing to catalyst treatment portion 300 via stream 294 and product and reactant gases (sometimes also referred to as high temperature dehydrogenation products) can pass out of primary separation device 402 via stream 286. The temperature of stream 286 can be about equal to the temperature of stream 284.

After the high temperature dehydrogenation product 286 exits the primary separation device 402, the product 286 is passed to a secondary separation device 404. Catalyst may be passed to catalyst treatment section 300 via stream 298. There is no quench fluid applied to the output of the primary separation device 402. The primary and secondary separation devices may be operated at temperatures in the range 600 ℃ to 700 ℃ and pressures in the range 1bara to 4 bara.

After the high temperature dehydrogenation product exits the secondary separation device via stream 288, the high temperature dehydrogenation product (or reactor effluent) is combined with quench stream 296 in distribution chamber section 204 of reactor system 102 to cool the high temperature dehydrogenation product and form stream 289 (sometimes referred to as an intermediate temperature dehydrogenation product). The intermediate temperature dehydrogenation product preferably has a temperature of not more than 620 ℃. The quench stream can be a gas or liquid stream that is part of stream 288 separated in a downstream separation system, as explained in detail herein. The quench stream can include one or more of ethylene, propylene, butene isomers, or benzene (e.g., reaction products). The quench stream can include liquid hydrocarbons and the intermediate temperature product can be entirely in the vapor phase.

For example, the quench fluid 296 may include steam, liquid water, liquid hydrocarbons (e.g., quench oil, fuel oil, or other hydrocarbons), or a combination of these. The liquid hydrocarbon may include hydrocarbons having greater than or equal to 6 carbon atoms, for example 6 carbon atoms to 25 carbon atoms, or 6 carbon atoms to 20 carbon atoms. In some embodiments, the quench fluid 296 may include water. Alternatively, in other embodiments, the quench fluid 296 may include liquid hydrocarbons. In still other embodiments, the quench fluid 296 may include one or more of benzene, toluene, pyrolysis gas, or a combination of these. The quench stream may include aromatic compounds that are not formed in the high temperature dehydrogenation product. The quench (stream) may comprise kerosene, light coker gas oil, coker (coker) distillate (CSD), hydrotreated distillate, or fresh untreated straight-run feedstock, e.g., straight-run gas oil, heavy straight-run naphtha, light straight-run naphtha, but preferably comprises a heavy aromatic solvent such as light catalytic cycle oil (LCCO or LCO), heavy catalytic cycle oil (HCCO or HCO), or Heavy Catalytic Naphtha (HCN), aromatic 100 (a 100) solvent, aromatic 150 (a 150) solvent, aromatic 200 (a 200) solvent, or any combination thereof.

Quench fluid 296 may reduce the temperature of stream 289 to about 620 ℃ to 580 ℃, wherein stream 289 experiences a pressure change of less than or equal to 0.1 bar due to gasification.

The temperature of the quench fluid 296 may be greater than or equal to 100 ℃, greater than or equal to 200 ℃, greater than or equal to 300 ℃, greater than or equal to 400 ℃, or greater than or equal to 550 ℃ below the temperature of the intermediate stream and catalyst in stream 288. The temperature of the quench stream and the nozzle location of injection of the quench stream can be optimized as desired to substantially slow down or stop the secondary thermal and catalytic reactions discussed above.

Stream 289 is then treated to cool its contents, for example via heat exchanger 406. In further embodiments, cooling may include adding another liquid quench system 408 or other known means of cooling the stream. Such other quench systems 408 may use any of the quench fluids described above. In general, the heat transfer system may include one or both of a heat exchanger 406 and other quench fluid 408 (or any other device that optionally includes a cooling stream). The product stream of heat exchanger 406 is stream 290, which may be referred to as cooled dehydrogenation product. The temperature of stream 290 can be about the same as described with respect to quench stream 296. It should be appreciated that stream 290 and/or stream 296 can be subjected to a pressure increase such that they flow in a desired direction.

To form quench stream 296, at least a portion of stream 290 is recycled back into the system via separation device 407 (which may be any suitable type of separation device or device for transferring a portion of the stream). It should be noted that the chemical content of stream 296 may be similar or identical to those of stream 286 (i.e., no additional reaction occurs from after those reactions in reactor 202, except for some residual thermal cracking). By contacting stream 288 with quench stream 296 such that quench stream 288 can cool the contents of stream 286 to a temperature that substantially reduces the reaction rate of thermal cracking and other secondary reactions. Stream 286 can be at a temperature at which thermal cracking occurs and such thermal cracking can reduce the selectivity of the desired reaction product.

Heat exchanger 406 may be configured to heat feed stream 280 by transferring heat from product stream 289.

According to one or more embodiments, the catalyst particles separated by separation device 220 and/or the catalyst particles recovered from stream 289 (not shown) may be treated by one or more steps prior to passing to catalyst treatment section 300.

Fig. 4 illustrates an exemplary reactor embodying the advances in the technology. The reactor design illustrated in fig. 4 minimizes the secondary reactions discussed above. In this configuration, the propane feed is contacted with the dehydrogenation catalyst at elevated temperature (> 600 ℃) in the riser section of the reactor. The dehydrogenation catalyst is separated from the high temperature product stream in primary and secondary separators. The resulting high temperature product stream comprises dehydrogenation product (propylene), byproducts (methane, benzene, etc.) and unconverted propane. This stream output from the secondary separator is immediately quenched with a liquid quench stream. The liquid quench stream may include byproducts of the dehydrogenation reaction (benzene) or other external fluids (water or a 150). The quench liquid is vaporized upon contact with the high temperature stream in the distribution chamber section of the reactor. Direct heat transfer results in reduced temperature and secondary reaction rates of the product stream.

Although fig. 4 is described with respect to a propylene/propane embodiment, the present advances in technology are applicable to the production of other olefins. Furthermore, while in some cases similar numbers to fig. 3 are used in fig. 4, it should be understood that the embodiment of fig. 3 may use a variety of reactor types, and fig. 4 is merely an example of one such type.

Fig. 4 illustrates an embodiment of a portion of the reactor system 102 of fig. 3, sometimes referred to as a reactor vessel 410, wherein the separation device comprises a primary separation device 420. The primary separation device 420 is housed within a housing 430 and has a body 421, an inlet 422, an outlet 424, and a solids discharge dipleg 426. The fluidized solids stream enters the primary separation device 420 at inlet 422. In the primary separation device 420, a substantial portion of entrained solids, such as catalyst particles, are separated from the fluidized solids stream. The separated solids leave the primary separation device through a discharge dipleg 426, leaving a primary separation device effluent containing solids and fluids, such as gaseous products, that were not removed by the primary separation device 420. The primary separation device effluent passes vertically upward and out of the primary separation device 420 through outlet 424 and through primary separation device outlet pipe 442 and then through jumper tubes (not shown) into the secondary separation device 440. The secondary separation device 440 further comprises a body 441, an outlet 444 and a solids discharge dipleg 446. The secondary separation device 440 further separates solids from the primary separation device effluent. Solids separated in secondary separation device 440 exit downwardly through dipleg 446.

As depicted in fig. 4, quench stream 296 may enter an upper portion of secondary separation device outlet conduit 444 that is located above secondary separation device 440. It is believed that such an arrangement may be desirable to properly mix the quench stream with the effluent of the secondary separation device 440 and reduce the residence time between the outlet of the secondary separation device 440 and the inlet or nozzle location of the quench stream 296. In some embodiments, the residence time of the high temperature dehydrogenation product in each of the primary separation device and the secondary separation device is less than 1 second.

The secondary separation device outlet 444 is fluidly connected to the second distribution chamber 450. The second distribution chamber allows the secondary separation device effluent to pass out of vessel 410 through outlet 470. As shown in fig. 4, the second dispensing chamber is contained within a larger, higher volume first dispensing chamber 460. The primary separation device 420 may be supported by a first distribution chamber 460.

The exact location of the inlet or nozzle of quench stream 296 may vary depending on the actual reactor used, but it may be disposed at a location within either first distribution chamber 460 or second distribution chamber 450 that minimizes secondary catalytic cracking within the high temperature dehydrogenation product and/or minimizes yield loss of olefins (i.e., propylene) within the high temperature dehydrogenation product (note that the conduit, pipe or duct of 296 in the second distribution chamber is shown in phantom because that portion would be behind the metal plates defining the second distribution chamber 450). Preferably, quench stream 296 can be configured to minimize the required piping so that all secondary separation devices can be quenched through a single inlet or nozzle; and thus quenched in the region of the second distribution chamber 450 to provide a simplified reactor design. Thus, while fig. 4 illustrates two possible locations for the inlet or nozzle of quench stream 296, one or both may be used in practice. However, it is important to note that quench stream is provided to the output region (i.e., output 424) of primary separator 420.

Also shown in fig. 4, the housing 430 also accommodates a riser 435. The unseparated stream of fluidized solid particles enters the housing through riser 435. The riser 435 is in fluid connection with the inlet 422 of the primary separation device 420, i.e. allows the passage of fluidized solid particles, such that an unseparated stream of fluidized solid particles can pass from the riser 435 to the primary separation device 420. It should be appreciated that while fig. 4 schematically illustrates only one primary separation device and one secondary separation device, additional primary and secondary separation devices may be placed around the circumference of the riser. For example, the second distribution chamber 450 may be connected to another secondary separation device (not shown) which in turn feeds through the primary separation device 420 or through another primary separation device (not shown).

Experiments to demonstrate catalyst stability

Those of ordinary skill in the art will appreciate that any catalyst suitable for their purpose may be used in the reactor embodying the advances in the present technology. In some embodiments, the catalyst may include 0.001 wt%, 0.002 wt%, 0.003 wt%, 0.004 wt%, 0.005 wt%, 0.006 wt%, 0.007 wt%, 0.008 wt%, 0.009 wt%, 0.01 wt%, 0.015 wt%, 0.02 wt%, 0.025 wt%, 0.03 wt%, 0.035 wt%, 0.04 wt%, 0.045 wt%, 0.05 wt%, 0.055 wt%, 0.06 wt%, 0.065 wt%, 0.07 wt%, 0.075 wt%, 0.08 wt%, 0.085 wt%, 0.09 wt%, 0.095 wt%, 0.1 wt%, 0.2 wt%, 0.3 wt%, 0.4 wt%, 0.5 wt%, 0.6 wt%, 0.7 wt%, 0.8 wt%, 0.9 wt%, or 1 wt% to 2 wt%, 3 wt%, 4 wt%, 5 wt%, or 6 wt% of the group 10 inorganic carrier disposed on the inorganic carrier. In some embodiments, the catalyst may include 5.5 wt.% or less, 4.5 wt.% or less, 3.5 wt.% or less, 2.5 wt.% or less, 1.5 wt.% or less, 1 wt.% or less, 0.9 wt.% or less, 0.8 wt.% or less, 0.7 wt.% or less, 0.6 wt.% or less, 0.5 wt.% or less, 0.4 wt.% or less, 0.3 wt.% or less, 0.2 wt.% or less, 0.15 wt.% or less, 0.1 wt.% or less, 0.09 wt.% or less, 0.08 wt.% or less, 0.07 wt.% or less, 0.06 wt.% or less, 0.05 wt.% or less, 0.04 wt.% or less, 0.03 wt.% or less, 0.02 wt.% or less, 0.01 wt.% or 0.009 wt.%, or less, 0.008 wt.% or less, 0.006 wt.% or less, 0.003 wt.% or less, 0.002 wt.% or less, based on the inorganic carrier, or no-group carrier, or no-load. In some embodiments, the catalyst may include >0.001, >0.003, >0.005, >0.007, >0.009, >0.01, >0.02, >0.04, >0.06, >0.08, >0.1, >0.13, >0.15, >0.17, >0.2, >0.23, >0.25, >0.27, >0.3, > 0.5, <1, <2, <3, <4, <5, <6, > by weight of the group 10 element disposed on the inorganic support, based on the weight of the inorganic support.

In some embodiments, the group 10 element may be or may include Ni, pd, pt, combinations thereof, or mixtures thereof. In at least one embodiment, the group 10 element may be or may include Pt. If two or more group 10 elements are provided on the inorganic support, the catalyst may include 0.001 wt%, 0.002 wt%, 0.003 wt%, 0.004 wt%, 0.005 wt%, 0.006 wt%, 0.007 wt%, 0.008 wt%, 0.009 wt%, 0.01 wt%, 0.015 wt%, 0.02 wt%, 0.025 wt%, 0.03 wt%, 0.035 wt%, 0.04 wt%, 0.045 wt%, 0.05 wt%, 0.055 wt%, 0.06 wt%, 0.065 wt%, 0.07 wt%, 0.075 wt%, 0.08 wt%, 0.085 wt%, 0.09 wt%, 0.095 wt%, 0.1 wt%, 0.2 wt%, 0.3 wt%, 0.4 wt%, 0.5 wt%, 0.6 wt%, 0.7 wt%, 0.8wt%, 0.9 wt%, or 1 wt% to 2 wt%, 3 wt%, 4wt%, 5 wt%, or 6 wt%, based on the total amount of two or more inorganic supports. In some embodiments, the active component of the regenerated catalyst that may be capable of effecting dehydrogenation of the hydrocarbon feed stream may include a group 10 element.

The inorganic support may be or include, but is not limited to, one or more group 2 elements, combinations thereof, or mixtures thereof. In some embodiments, the group 2 element may be present in its elemental form. In other embodiments, the group 2 element may be present in the form of a compound. For example, the group 2 element may be present as an oxide, phosphate, halide, halite, sulfate, sulfide, borate, nitride, carbide, aluminate, aluminosilicate, silicate, carbonate, metaphosphate, selenide, tungstate, molybdate, chromite, chromate, dichromate, or silicide. In some embodiments, a mixture of any two or more compounds comprising a group 2 element may exist in different forms. For example, the first compound may be an oxide and the second compound may be an aluminate, wherein the first compound and the second compound include the same or different group 2 elements relative to each other.

The inorganic carrier may include 0.5 wt% or more, 1 wt% or more, 2 wt% or more, 3 wt% or more, 4 wt% or more, 5 wt% or more, 6 wt% or more, 7 wt% or more, 8 wt% or more, 9 wt% or more, 10 wt% or more, 11 wt% or more, 12 wt% or more, 13 wt% or more, 14 wt% or more, 15 wt% or more, 16 wt% or more, 17 wt% or more, 18 wt% or more, 19 wt% or more, 20 wt% or more, 21 wt% or more, 22 wt% or more, 23 wt% or more, 24 wt% or more, 25 wt% or more, 26 wt% or more, 27 wt% or more, 28 wt% or more, 29 wt% or more, 30 wt% or more, 35 wt% or more, 40 wt% or more, 45 wt% or more, 50 wt% or more, 55 wt% or more, 60 wt% or more, or less, 85 wt% or less, or more, and 75 wt% or more of the carrier elements. In some embodiments, the inorganic support may include a group 2 element in a range from 0.5 wt%, 1 wt%, 2 wt%, 2.5 wt%, 3 wt%, 5 wt%, 7 wt%, 10 wt%, 11 wt%, 13 wt%, 15 wt%, 17 wt%, 19 wt%, 21 wt%, 23 wt%, or 25 wt% to 30 wt%, 35 wt%, 40 wt%, 45 wt%, 50 wt%, 55 wt%, 60 wt%, 65 wt%, 70 wt%, 75 wt%, 80 wt%, 85 wt%, 90 wt%, or 92.34 wt%, based on the weight of the inorganic support. In some embodiments, the molar ratio of group 2 element to group 10 element may be in the range from 0.24, 0.5, 1, 10, 50, 100, 300, 450, 600, 800, 1,000, 1,200, 1,500, 1,700, or 2,000 to 3,000、3,500、4,000、4,500、5,000、5,500、6,000、6,500、7,000、7,500、8,000、8,500、9,000、9,500、10,000、15,000、20,000、25,000、30,000、35,000、40,000、45,000、50,000、55,000、60,000、65,000、70,000、75,000、80,000、85,000、90,000、95,000、100,000、200,000、300,000、400,000、500,000、600,000、700,000、800,000、 or 900,000.

In some embodiments, the inorganic support may include a group 2 element and Al and may be in the form of a mixed group 2 element/Al metal oxide having O, mg and Al atoms mixed on an atomic scale. In some embodiments, the inorganic support may be or may include a group 2 element and Al in the form of an oxide or one or more oxides of a group 2 element and Al ₂O₃ that may be mixed on the nm scale. In some embodiments, the inorganic support may be or may include oxides of group 2 elements such as MgO and Al ₂O₃ mixed on the nm scale.

In some embodiments, the inorganic support may be or may include a first amount of a group 2 element and Al in the form of a mixed group 2 element/Al metal oxide and a second amount of a group 2 element in the form of an oxide of the group 2 element. In such embodiments, the mixed group 2 element/Al metal oxide and the oxide of the group 2 element may be mixed on the nm scale and the group 2 element and Al in the mixed group 2 element/Al metal oxide may be mixed on the atomic scale.

In other embodiments, the inorganic support may be or may include a first amount of a group 2 element and a first amount of Al in the form of a mixed group 2 element/Al metal oxide, a second amount of a group 2 element in the form of an oxide of a group 2 element, and a second amount of Al in the form of Al ₂O₃. In such embodiments, the mixed group 2 element/Al metal oxide, the oxide of the group 2 element, and Al ₂O₃ may be mixed on the nm scale and the group 2 element and Al in the mixed group 2 element/Al metal oxide may be mixed on the atomic scale.

In some embodiments, when the inorganic support includes a group 2 element and Al, the weight ratio of the group 2 element to Al in the inorganic support can be in the range from 0.001, 0.005, 0.01, 0.05, 0.1, 0.15, 0.2, 0.3, 0.5, 0.7, or 1 to 3, 6, 12.5, 25, 50, 75, 100, 200, 300, 400, 500, 600, 700, 800, 900, or 1,000. In some embodiments, when the inorganic support comprises Al, the inorganic support may comprise Al in a range from 0.5 wt%, 1 wt%, 1.5 wt%, 2 wt%, 2.1 wt%, 2.3 wt%, 2.5 wt%, 2.7 wt%, 3 wt%, 4 wt%, 5 wt%, 6 wt%, 7 wt%, 8 wt%, 9 wt%, 10 wt%, or 11 wt% to 15 wt%, 20 wt%, 25 wt%, 30 wt%, 40 wt%, 45 wt%, or 50 wt%, based on the weight of the inorganic support.

In some embodiments, the catalyst may be prepared by depositing Pt and Sn onto a support containing at least one alkaline earth metal oxide or at least one mixed metal oxide containing at least one alkaline earth metal. The composition of the catalyst may be as follows: pt 0.001 wt% to 6 wt%, based on the weight of the support; 0 to 10 wt% Sn, based on the weight of the carrier; the support comprises at least one alkaline earth metal oxide. The total alkaline earth metal content may be at least 0.5 wt%, based on the weight of the support. The catalyst may have a Geldart A or B type classified according to the Geldart of the particles.

Other catalysts may be used and may be at least partially deactivated at some point in the embodiments of the method and/or system. The catalyst may include a group 10 element, an inorganic support, and a contaminant, wherein the group 10 element has a concentration in the range of 0.001 wt% to 6 wt%, based on the weight of the inorganic support. The group 10 element may be or may include Pt, and wherein the inorganic support comprises at least 0.5 wt% of the group 2 element, based on the weight of the inorganic support. The group 2 element may include Mg, and at least a portion of the group 2 element may be in the form of MgO or a mixed metal oxide including Mg. The catalyst may further comprise up to 10 wt% of a promoter, based on the weight of the inorganic support, and wherein the promoter may comprise one or more of the following elements: sn, ag, cu, combinations thereof, or mixtures thereof. The catalyst may further comprise up to 5 wt% of an alkali metal element disposed on the inorganic support, wherein the alkali metal element may comprise at least one of: li, na, K, rb, and Cs. Additional examples of catalysts useful in the advancement of the present technology can be found in application Ser. No. 63/195966 filed on 6/2 of 2021, which is incorporated herein by reference in its entirety.

Catalyst 1 hydrotalcite containing calcination by spray dryingAn aqueous slurry of MG70/170 (Sasol) and Aluminum Chlorohydrate (ACH) was used to prepare the Mg/Al oxide support. The final support after calcination contained 80 wt.% calcined hydrotalcite and 20 wt.% alumina derived from ACH. An aqueous solution of tin (IV) chloride pentahydrate (Acros Organics) and chloroplatinic acid hexahydrate (BioXtra) was impregnated onto the support. The impregnated material was kept at room temperature in a closed vessel for 45 hours, then dried at 110 ℃ for 6 hours and calcined at 800 ℃ for 12 hours, all in air. The final product nominally contains 0.3 wt% Pt and 1.5 wt% Sn.

Fixed bed experiments were conducted at about 100 kPa-absolute unless otherwise specified. Gas Chromatography (GC) was used to measure the composition of the reactor effluent. The concentration of each component in the reactor effluent was then used to calculate the C ₃H₆ yield and selectivity. The C ₃H₆ yield and selectivity as reported in these examples were calculated on a carbon molar basis.

In each example, an amount of catalyst "M _cat" was mixed with an appropriate amount of quartz diluent and loaded into a quartz reactor. The amount of diluent is determined such that the catalyst bed (catalyst + diluent) overlaps the isothermal zone of the quartz reactor and such that the catalyst bed is largely isothermal during operation. The dead volume of the reactor was filled with quartz chips/rods.

The concentration of each component in the reactor effluent was used to calculate the C ₃H₆ yield and selectivity. the yield and selectivity of C ₃H₆ at the beginning of t _rxn and at the end of t _rxn are denoted Y _ini、Y_end、S_ini and S _end, respectively, and are reported as a percentage in the data table below.

Example 1:1. the system was purged with inert gas. 2. Oxygen-containing gas (O _{Air flow}) is passed through the branch of the reaction zone at a flow rate (F _oxi) while inert gas (inert) is passed through the reaction zone. The reaction zone is heated to an oxidation temperature T _oxi. 3. The oxygen-containing gas is then passed through the reaction zone for a period of time (t _oxi) to oxidize the catalyst. After T _oxi, the temperature within the reaction zone is changed from T _oxi to the reduction temperature (T _red) while maintaining the flow of oxygen-containing gas. 4. The system was purged with inert gas. 5. The H ₂ -containing gas (H gas) is passed through the branch of the reaction zone at a flow rate (F _red) for a period of time while the inert gas is passed through the reaction zone. This is then followed by flowing the H ₂ -containing gas through the reaction zone at T _red for a period of time (T _red). 6. The system was purged with inert gas. During this process, the temperature of the reaction zone was changed from T _red to a reaction temperature of 655 ℃.7. A Hydrocarbon (HC) containing feed comprising 81% by volume C ₃H₈, 9% by volume inert gas (Ar or Kr) and 10% by volume steam is passed through the branch of the reaction zone at a flow rate (F _rxn) for a period of time while inert gas is passed through the reaction zone. The hydrocarbonaceous feed was then passed through the reaction zone at 655 ℃ for 10 minutes. GC sampling of the reaction effluent was started as soon as the feed was switched from branch to reaction zone.

The above method steps are cyclically repeated until stable performance is obtained. Table 1 shows that catalyst 1 is active/selective for propane dehydrogenation and can be effectively regenerated for 60+ cycles by using such a two-step oxidation scheme. Fig. 5 shows that catalyst 1 is stable for propane dehydrogenation over 60 cycles.

Example 2: the simulated propane dehydrogenation product was fed to a reaction zone filled with inert quartz chips to investigate its thermal cracking at different temperatures and pressures. The molar composition of the simulated propane dehydrogenation product is shown in table 2.

The simulated propane dehydrogenation products correspond to propylene yields and selectivities of 63.7% and 95.5%, respectively.

The components of the simulated propane dehydrogenation product are further cracked as they pass through the reactor. The cracking rate was calculated at differential conversion (DIFFERENTIAL CONVERSION) and expressed as propylene yield per second and selectivity loss. FIG. 1 is a graph illustrating propylene yield loss per second at various temperatures (x-axis is temperature in degrees Celsius and y-axis is%). Fig. 2 is a graph illustrating the selectivity loss of propylene per second at different temperatures (x-axis is temperature in degrees celsius and y-axis is%).

The above experiments were performed using-5 to 10mol% steam co-fed with the simulated propane dehydrogenation product. Small amounts of steam were found to have little effect on the cracking rate. The minimal impact on cracking rate can be attributed to the lower hydrocarbon partial pressure due to the presence of steam.

Example 3: the following experiments were conducted to understand the reaction of propane dehydrogenation product with entrained catalyst not separated by the cyclone as the propane dehydrogenation product with entrained catalyst at various temperatures flows through equipment downstream of the cyclone.

Feed a (consisting of simulated propane dehydrogenation product and an appropriate amount of steam) was fed to a reaction zone filled with inert quartz chips and various amounts of catalyst 2 to investigate the hydrogenation, thermal cracking and catalytic cracking of feed a at different temperatures. Catalyst 2 was prepared according to the following procedure: leave 2.3gMG 70/170 (Sasol), which is a MgO-Al ₂O₃ mixed metal oxide obtained by calcining hydrotalcite. The mixed metal oxide contains 70 wt% MgO and 30 wt% Al ₂O₃. BET surface area was 170m ²/g according to Sasol. Tin (IV) chloride pentahydrate (0.103 g) (Acros Organics), chloroplatinic acid hexahydrate (0.0184 g) (BioXtra) and deionized water (2.2 mL) were mixed in a vial to prepare a solution. Impregnating/>, with the solutionMG 70/170 vector. The impregnated material was dried at 110 ℃ for 6 hours and calcined at 800 ℃ for 12 hours, all in air. The final product nominally contains 0.3 wt% Pt and 1.5 wt% Sn.

The molar composition of feed a is shown in table 3.

Feed a corresponds to a simulated propane dehydrogenation product with a propylene yield and selectivity of 63.7% and 95.5%, respectively.

The total pressure of feed A was 100kPa absolute. The amount of catalyst 2 in the reaction zone and the flow rate of feed a were varied such that the weight hourly space velocity (WHSV, calculated by dividing the mass flow rate of the C-containing component in feed a by the weight of catalyst 2 in the reaction zone) was varied from 121h ^-1、350h^-1 and infinity (no catalyst).

The flow rate of feed a was also varied such that the thermal residence time of feed a in the reaction zone (calculated by dividing the void volume of the reaction zone by the volumetric flow rate of feed a at the reaction temperature/pressure) was 2 seconds. The reaction zone is largely isothermal. The temperature outside the reaction zone drops rapidly.

While passing through the reaction zone, feed a may be further cracked (thermal or catalytic), dehydrogenated, or hydrogenated to form product a. The propylene yield and selectivity (63.7% and 95.5%) of feed a may change after the reaction. For example, if hydrogenation of propylene occurs in the reaction zone, the propylene yield of product a decreases.

FIG. 6 (x-axis temperature in degrees Celsius; y-axis propylene yield (%)) shows a plot of propylene yield of product A versus reaction zone temperature when the thermal residence time is maintained at 2 seconds. WHSV varies from 121h ^-1 (orange, line (b)), 350h ^-1 (blue, line (a)) and endless (no catalyst) (gray, line (c)). When no catalyst is present, the propylene yield loss at T >600 ℃ is entirely due to thermal cracking of propylene. When a catalyst is present, substantial losses of propylene are seen at T <600 ℃, mainly due to the hydrogenation of propylene. The extent of the hydrogenation reaction decreases rapidly as the WHSV increases, i.e., as the amount of catalyst in the propane dehydrogenation product that is not separated by the cyclone decreases. The propylene yield loss due to hydrogenation is greatest at T-500 ℃.

FIG. 7 shows a plot of propylene selectivity for product A versus reaction zone temperature when the thermal residence time is maintained at 2 s. WHSV varies from 121h ^-1 (orange, line (b)), 350h ^-1 (blue, line (a)) and endless (no catalyst) (gray, line (c)). In the absence of catalyst, the loss of propylene selectivity at T >600 ℃ is entirely due to thermal cracking of propylene. Interestingly, propylene selectivity was higher when WHSV was 121h ^-1. This may be due to the inhibition of thermal cracking by catalyst 2.

Figures 6 and 7 show that hydrogenation of propylene is avoided during quenching. The following strategy may be employed. (1) The amount of catalyst entrained in the propane dehydrogenation product is reduced prior to quenching the product. (2) The propane dehydrogenation product was rapidly quenched to <320 ℃ because the hydrogenation rate was slow when T <320 ℃. (3) Quenching is used which reversibly deactivates the entrained catalyst.

Catalyst compositions 3-16 were prepared according to the following procedure. Calcination in air at 550℃for each catalyst compositionMG 80/150 (3 g) (Sasol), which is a mixed Mg/Al metal oxide containing 80% MgO by weight and 20% Al ₂O₃ by weight and having a surface area of 150m ²/g, was allowed to form a support for 3 hours. A solution containing the appropriate amount of tin (IV) chloride pentahydrate (Acros Organics) (when used to prepare the catalyst composition) and/or chloroplatinic acid (SIGMA ALDRICH) (when used to prepare the catalyst composition) and 1.8ml of deionized water was prepared in a vial. For each catalyst composition, the calcined/>, was impregnated with the corresponding solutionMG 80/150 support (2.3 g). The impregnated material was equilibrated in a closed vessel at Room Temperature (RT) for 24 hours, dried at 110 ℃ for 6 hours and calcined at 800 ℃ for 12 hours. Table 1 shows nominal Pt and Sn content of each catalyst composition based on the weight of the support.

Example 4-fixed bed experiments using catalysts 3-16 were performed at about 100 kPa-absolute. Gas Chromatography (GC) was used to measure the composition of the reactor effluent. The concentration of each component in the reactor effluent was then used to calculate the C ₃H₆ yield and selectivity. The C ₃H₆ yield and selectivity as reported in these examples were calculated on a carbon molar basis.

In each example, 0.3g of the catalyst composition was mixed with an appropriate amount of quartz diluent and loaded into a quartz reactor. The amount of diluent is determined such that the catalyst bed (catalyst + diluent) overlaps the isothermal zone of the quartz reactor and such that the catalyst bed is largely isothermal during operation. The dead volume of the reactor was filled with quartz chips/rods.

The yield and selectivity of C ₃H₆ at the beginning of t _rxn and at the end of t _rxn are expressed as Y _ini、Y_end、S_ini and S _end, respectively, and are reported as a percentage for catalysts 3-10 in tables 5 and 6 below.

The method steps of the catalyst 3-10 are as follows: 1. the system was purged with inert gas. 2. Drying air was passed through the branches of the reaction zone at a flow rate of 83.9sccm while inert gas was passed through the reaction zone. The reaction zone was heated to a regeneration temperature of 800 ℃.3. Dry air was then passed through the reaction zone at a flow rate of 83.9sccm for 10min to regenerate the catalyst. 4. The system was purged with inert gas. 5. An H ₂ -containing gas having 10% H ₂ by volume and 90% Ar by volume was passed through the legs of the reaction zone at a flow rate of 46.6sccm for a period of time while an inert gas was passed through the reaction zone. This is then followed by flowing a gas containing H ₂ through the reaction zone at 800 ℃ for 3 seconds. 6. The system was purged with inert gas. During this process, the temperature of the reaction zone was changed from 800 ℃ to a reaction temperature of 670 ℃.7. A Hydrocarbon (HC) containing feed comprising 81% by volume C ₃H₈, 9% by volume inert gas (Ar or Kr) and 10% by volume steam was passed through the branch of the reaction zone at a flow rate of 35.2sccm for a period of time while inert gas was passed through the reaction zone. The hydrocarbonaceous feed was then passed through the reaction zone at 670 ℃ for 10min. GC sampling of the reaction effluent was started as soon as the feed was switched from branch to reaction zone.

The above method steps are cyclically repeated until stable performance is obtained. Tables 5 and 6 show that catalyst 8 containing only 0.025 wt% Pt and 1 wt% Sn has similar yields and similar selectivities as compared to catalyst 3 containing 0.4 wt% Pt and 1 wt% Sn, which is surprising and unexpected. Catalyst 10, which did not include any Pt, did not show significant propylene yield.

Catalysts 11-16 were also tested using the same method steps 1-7 described above with respect to catalysts 3-10. Table 7 shows that for optimal propylene yields for catalyst compositions comprising 0.1 wt% Pt based on the weight of the support, the Sn level should not be too low or too high.

Table 8 shows that for optimal propylene yields for catalyst compositions comprising 0.0125 wt% Pt based on the weight of the support, the Sn level should not be too high or too low.

Catalyst 8, containing only 0.025 wt% Pt and 1 wt% Sn, was also subjected to life testing using the same method steps 1-7 described above with respect to catalysts 3-10, except that a flow rate of 17.6sccm was used instead of 35.2sccm in step 7. Figure 8 shows that catalyst 8 maintains performance for 204 cycles (x-axis is time, y-axis is C ₃H₆ yield and C ₃H₆ selectivity, all in mole% carbon).

All documents described herein are incorporated by reference herein, including any priority documents and/or testing procedures, so long as they are not inconsistent with the present disclosure. As will be apparent from the foregoing general description and specific embodiments, while forms of the disclosure have been illustrated and described, various changes can be made without departing from the spirit and scope of the disclosure. Accordingly, it is not intended that the disclosure be limited thereby. Also, whenever a constituent, element or group of elements is preceded by the term "comprising", it should be understood that we also contemplate the same constituent or group of elements preceded by the term "consisting essentially of", "consisting of", "selected from the group consisting of" or "being" and vice versa.

While the present disclosure has been described with respect to a number of embodiments and examples, those skilled in the art, having benefit of this disclosure, will appreciate that other embodiments can be devised which do not depart from the scope and spirit of the present disclosure.

Claims

1. A process for forming an olefin, the process comprising:

introducing a hydrocarbon feedstream to a reactor comprising a dehydrogenation catalyst;

reacting the hydrocarbon feedstream with a dehydrogenation catalyst in a reactor to form a high temperature dehydrogenation product comprising at least a portion of the dehydrogenation catalyst;

separating at least a portion of the dehydrogenation catalyst from the high temperature dehydrogenation product in a primary separation device and a secondary separation device downstream of and in fluid communication with the primary separation device;

Combining the high temperature dehydrogenation product with a quench stream after the high temperature dehydrogenation product exits the secondary separation device to cool the high temperature dehydrogenation product and form an intermediate temperature dehydrogenation product, wherein the quench stream comprises hydrocarbons; and

The intermediate temperature dehydrogenation product is cooled to form a cooled dehydrogenation product.

2. The process of claim 1, wherein the quench stream comprises liquid hydrocarbons and the intermediate temperature product are entirely in the vapor phase.

3. The process of claim 1 or claim 2, wherein combining the high temperature dehydrogenation product with the quench stream occurs within a distribution chamber section of the reactor.

4. The method of claim 3, wherein combining comprises combining the high temperature dehydrogenation product from the plurality of secondary separation devices with the quench stream at a location within a distribution chamber that minimizes secondary catalytic cracking within the high temperature dehydrogenation product.

5. The process of any of the preceding claims, wherein the quench stream comprises aromatic compounds that are not formed in the high temperature dehydrogenation product.

6. The process of any one of the preceding claims, further comprising further cooling the intermediate temperature dehydrogenation product downstream of the reactor using a heat exchanger or a secondary quench stream to obtain a cooled dehydrogenation product,

Wherein the quench stream comprises at least a portion of the cooled dehydrogenation product after having been further cooled by the heat exchanger or the secondary quench stream.

7. The process of any of the preceding claims, wherein no quench stream is provided to the output zone of the primary separation device.

8. The process of any of the preceding claims, wherein the primary separation device and the secondary separation device are each cyclones and they together remove at least 98% of the dehydrogenation catalyst present in the high temperature dehydrogenation product.

9. The process of any of the preceding claims, wherein the hydrocarbon feed is propane, the olefin is propylene, and the dehydrogenation catalyst comprises a PtSn/MgO catalyst, wherein the Pt content ranges from 0.001 wt% to 6 wt%, the Sn content ranges from 0wt% to 10 wt%, and the catalyst meets the Geldar t A or Geldar t B classification requirements, and the reaction temperature within the reactor ranges from 600 ℃ to 700 ℃.

10. The process of any of the preceding claims, wherein the high temperature dehydrogenation product is at least 620 ℃.

11. The process of any of the preceding claims, wherein the intermediate temperature dehydrogenation product is no greater than 620 ℃.

12. The process of any of the preceding claims, wherein the residence time of the high temperature dehydrogenation product in each of the primary separation device and the secondary separation device is less than 1 second.

13. The method of any of the preceding claims, wherein:

The dehydrogenation catalyst comprises Pt, an inorganic support, and a contaminant, wherein the Pt has a concentration in the range of 0.001 wt% to 6wt% based on the weight of the inorganic support,

The inorganic support comprises at least 0.5 wt% of a group 2 element, based on the weight of the inorganic support,

The group 2 element comprises Mg and,

At least a portion of the group 2 element is in the form of MgO or a mixed metal oxide containing Mg, and

The catalyst further comprises up to 10 wt% of a promoter, based on the weight of the inorganic support, and wherein the promoter comprises one or more of the following elements: sn, ag, cu, combinations thereof, or mixtures thereof.

14. A system for forming olefins from a hydrocarbon feed stream, the system comprising:

A reactor configured to receive a hydrocarbon feedstream and a dehydrogenation catalyst under reaction conditions, wherein the reaction conditions produce a high temperature dehydrogenation product comprising at least a portion of the dehydrogenation catalyst;

A plurality of separation devices configured to separate at least a portion of the dehydrogenation catalyst from the high temperature dehydrogenation product, the plurality of separation devices comprising a primary separation device and a secondary separation device downstream of and in fluid communication with the primary separation device;

A quench system configured to combine the high temperature dehydrogenation product with a quench stream to cool the high temperature dehydrogenation product and form an intermediate temperature dehydrogenation product, wherein the quench stream comprises hydrocarbons; and

A heat transfer system configured to cool an intermediate temperature dehydrogenation product to form a cooled dehydrogenation product.

15. The system of claim 14, wherein the quench stream comprises liquid hydrocarbons and the intermediate temperature product are entirely in the gas phase.

16. The system of claim 14 or claim 15, wherein the quench system is configured to combine the high temperature dehydrogenation product with the quench stream within a distribution chamber section of the reactor.

17. The system of any of claims 14 to 16, wherein the quench system is configured to inject the quench stream at a location within the distribution chamber that minimizes secondary catalytic cracking within the high temperature dehydrogenation product.

18. The system of any one of claims 14 to 17, wherein the heat transfer system comprises a heat exchanger or a nozzle configured to supply a secondary quench stream.

19. The system of any one of claims 14 to 18, wherein the quench system is configured to not provide a quench stream to the output area of the primary separation device.

20. The system of any of claims 14 to 19, wherein the primary separation device and the secondary separation device are each cyclones and are collectively configured to remove at least 98% of the dehydrogenation catalyst present in the high temperature dehydrogenation product.