阅读说明：本技术 用于处理气流的方法和装置 (Method and apparatus for treating a gas stream ) 是由 M·T·普雷兹 M·科达姆 于 2018-07-30 设计创作，主要内容包括：根据本文所公开的一个或多个实施例,可以通过包含将反应气体引入流化床反应器的方法来转化所述反应气体。所述流化床反应器的主反应器容器可以是锥形的,使得所述主反应器容器的上游部分包含比所述主反应器容器的下游部分更小的截面积。(According to one or more embodiments disclosed herein, the reaction gas may be converted by a method comprising introducing the reaction gas into a fluidized bed reactor. The main reactor vessel of the fluidized bed reactor may be tapered such that an upstream portion of the main reactor vessel comprises a smaller cross-sectional area than a downstream portion of the main reactor vessel.)

1. A method for converting a reactant gas, the method comprising:

introducing the reaction gas into a fluidized bed reactor such that the reaction gas contacts a catalyst, wherein the fluidized bed reactor comprises a main reactor vessel comprising an upstream portion and a downstream portion, and a transition section connected to the downstream portion of the main reactor vessel, and wherein the reaction gas enters the fluidized bed reactor at or near the upstream portion of the main reactor vessel;

catalytically reacting the reaction gas in the fluidized bed reactor to form a reaction product, wherein the reaction produces additional gas molecules relative to the reaction gas; and is

Passing the reaction products and any unreacted reaction gases through the transition section;

wherein the main reactor vessel is tapered such that an upstream portion of the main reactor vessel comprises a smaller cross-sectional area than a downstream portion of the main reactor vessel.

2. The method of claim 1, wherein a superficial velocity of the gas in the fluidized bed reactor at the downstream portion of the main reactor vessel is less than or equal to 140% of a superficial velocity of the gas in the fluidized bed reactor at the upstream portion of the main reactor vessel.

3. The process of claim 1, wherein the suspended density in the fluidized bed reactor at the downstream portion of the main reactor vessel is greater than or equal to 25% of the suspended density of the gas in the fluidized bed reactor at the upstream portion of the main reactor vessel.

4. The method of claim 1, wherein the reaction product is a dehydrogenation product of the reaction gas.

5. The method of claim 4, wherein the reaction gas comprises one or more of ethane, propane, n-butane, isobutane, and ethylbenzene.

6. The method of claim 1, wherein the main reactor vessel comprises a central portion between the downstream portion and the upstream portion, and wherein a superficial velocity of the gas in the fluidized bed reactor at the central portion of the main reactor vessel is 60% to 140% of a superficial velocity of the gas in the fluidized bed reactor at the upstream portion of the main reactor vessel.

7. The method of claim 1, wherein the main reactor vessel comprises a central portion between the downstream portion and the upstream portion, and wherein a superficial velocity of the gas in the fluidized bed reactor at the central portion of the main reactor vessel is 80% to 120% of a superficial velocity of the gas in the fluidized bed reactor at the upstream portion of the main reactor vessel.

8. The method of claim 1, wherein the superficial velocity of the gas in the fluidized bed reactor at the downstream portion of the main reactor vessel is 80% to 120% of the superficial velocity of the gas in the fluidized bed reactor at the upstream portion of the main reactor vessel.

9. The method of claim 1, wherein the primary reactor vessel comprises a substantially constant draft.

10. The method of claim 1, wherein the main reactor vessel comprises a lesser draft at or near the downstream portion than at or near the upstream portion.

11. The method of claim 1, wherein the primary reactor vessel comprises a central portion between the downstream portion and the upstream portion, and wherein a cross-sectional area of the central portion is less than a cross-sectional area of the downstream portion and greater than a cross-sectional area of the upstream portion.

12. The process of claim 1, wherein the fluidized bed reactor further comprises a riser connected to the transition section, and wherein the reaction products and any unreacted reaction gas pass through the transition section and into the riser.

13. A method for converting a reactant gas, the method comprising:

introducing the reaction gas into a fluidized bed reactor such that the reaction gas contacts a catalyst, wherein the fluidized bed reactor comprises a main reactor vessel comprising an upstream portion and a downstream portion, and a transition section connected to the downstream portion of the main reactor vessel, and wherein the reaction gas enters the fluidized bed reactor at or near the upstream portion of the main reactor vessel;

catalytically reacting the reaction gas to form a reaction product in the fluidized bed reactor, wherein the reaction produces additional gas molecules relative to the reaction gas; and is

Passing the reaction products and any unreacted reaction gases through the transition section;

wherein the main reactor vessel is tapered such that the upstream portion of the main reactor vessel comprises a smaller cross-sectional area than the downstream portion of the main reactor vessel, such that a superficial velocity of the gas in the fluidized bed reactor at the downstream portion of the main reactor vessel is less than or equal to 140% of a superficial velocity of the gas in the fluidized bed reactor at the upstream portion of the main reactor vessel, or such that a suspended density in the fluidized bed reactor at the downstream portion of the main reactor vessel is greater than or equal to 25% of a suspended density of the gas in the fluidized bed reactor at the upstream portion of the main reactor vessel, or both.

14. The method of claim 13, wherein the reaction product is a dehydrogenation product of the reaction gas.

15. The method of claim 14, wherein the reaction gas comprises one or more of ethane, propane, n-butane, isobutane, and ethylbenzene.

Technical Field

The present disclosure relates generally to chemical processes and, more particularly, to reactor designs and systems for chemical processes.

Background

Many reactions that utilize gas streams as reactants form additional moles of gas as a result of the reaction (i.e., when there are more gas molecules after the reaction than before). For example, dehydrogenation and cracking reactions produce additional moles of product compared to the product present in the reactant stream. When the reactants and products are gases, local gas velocity variations can be produced in the reactor by the formation of these additional molecules.

Disclosure of Invention

When the reaction produces additional gas molecules as compared to the reactant stream, pressure may build up in the reactor or other system components. These pressure variations can result in local variations in the velocity of the gas and/or solid catalyst in a reactor, such as a fluidized bed reactor. Variations in gas velocity can also affect the suspension density in localized portions of the reactor, which is generally related to the amount of catalyst per volume in the reactor at a localized location. Controlling the superficial velocity of the fluid and/or the suspension density in the reactor may be important to the overall chemical conversion of the reactor and/or reactor operating specifications (e.g., size, shape, etc.). Thus, there remains a need for methods and apparatus for treating gas streams under reaction conditions that increase the number of gas molecules while controlling the superficial velocity and/or the suspended density throughout the reactor.

More specifically, it has been found that the formation of excess gas molecules relative to the amount of gas fed to the reactants after the reaction can result in an increase in the superficial gas velocity and a decrease in the suspended density of the reactor contents. As used herein, "suspended density" refers to the density calculated from both the solids content (e.g., catalyst) and the gas content (e.g., gaseous reactants and products) in the reactor. For example, as more gas is produced by the reaction, the superficial velocity of the gas may increase, sometimes dramatically, in a fluidized bed reactor. In addition, the suspension density (including reactant and product gases and solid particulate catalyst) may be reduced such that the conversion to reactant gas is reduced due to the lack of catalyst. Because conversion is adversely affected, it may be necessary to increase the volume of the reactor, thereby increasing undesirable capital costs. In addition, high gas velocities in the reactor can make it difficult to control the amount of catalyst in the system at a given time.

To mitigate or completely prevent the increase in gas velocity and the decrease in suspension density, it has been found that fluidized bed reactors with increased cross-sectional area can be used for gas-based reactions that produce excess gas molecules. For example, a fluidized bed reactor that is narrower in its upstream portion than its downstream portion may allow for a relatively constant gas superficial velocity, suspension density, or both in the upstream and downstream portions of the fluidized bed reactor. In contrast, conventional fluidized bed reactors having a tubular shape generally have a gas velocity that varies with height. Appropriately designed geometries (e.g., tapered geometries) of the fluidized bed reactor can allow for a reduction in the increase in gas velocity and/or loss in suspension density as the reaction proceeds in the downstream portion of the reactor. That is, as the reaction proceeds along the height of the reactor (assuming the reaction products move upward), the increase in cross-sectional area compensates for the increase in gas volume and thus a relatively constant gas velocity can be maintained. The relative stability of the gas velocity, the suspension density, or both may allow the above-described problems to be alleviated.

In accordance with one or more embodiments, a reaction gas may be converted by a process comprising introducing the reaction gas into a fluidized bed reactor, contacting the reaction gas with a catalyst, catalytically reacting the reaction gas to form reaction products in the fluidized bed reactor, and passing the reaction products and any unreacted reaction gas through a transition zone. The fluidized bed reactor can include a main reactor vessel (including an upstream portion and a downstream portion), and a transition section connected to the downstream portion of the main reactor vessel. The reaction gas may enter the fluidized bed reactor at or near an upstream portion of the main reactor vessel. The reaction may produce additional gas molecules relative to the reactant gas. The main reactor vessel may be tapered such that an upstream portion of the main reactor vessel comprises a smaller cross-sectional area than a downstream portion of the main reactor vessel, such that a superficial velocity of gas in the fluidized bed reactor at the downstream portion of the main reactor vessel is less than or equal to 140% of a superficial velocity of gas in the fluidized bed reactor at the upstream portion of the main reactor vessel. In accordance with one or more additional embodiments, a reaction gas may be converted by a method comprising introducing a reaction gas into a fluidized bed reactor, contacting the reaction gas with a catalyst, catalytically reacting the reaction gas to form a reaction product in the fluidized bed reactor, and passing the reaction product and any unreacted reaction gas through a transition section. The fluidized bed reactor can include a main reactor vessel (including an upstream portion and a downstream portion), and a transition section connected to the downstream portion of the main reactor vessel. The reaction gas may enter the fluidized bed reactor at or near an upstream portion of the main reactor vessel. The reaction may produce additional gas molecules relative to the reactant gas. The main reactor vessel may be tapered such that an upstream portion of the main reactor vessel comprises a smaller cross-sectional area than a downstream portion of the main reactor vessel, such that a suspension density in the fluidized bed reactor at said downstream portion of the main reactor vessel is greater than or equal to 25% of a suspension density of the gas in said fluidized bed reactor at the upstream portion of the main reactor vessel.

It is to be understood that both the foregoing general description and the following detailed description present embodiments of the technology, and are intended to provide an overview or framework for understanding the nature and character of the technology as it is claimed. The accompanying drawings are included to provide a further understanding of the present technology and are incorporated in and constitute a part of this specification. The drawings illustrate various embodiments and together with the description serve to explain the principles and operations of the present technology. Additionally, the drawings and description are meant to be illustrative only and are not intended to limit the scope of the claims in any way.

Additional features and advantages of the technology disclosed herein will be set forth in the detailed description which follows, and in part will be readily apparent to those skilled in the art from that description or recognized by practicing the technology as described herein, including the detailed description which follows, the claims, as well as the appended drawings.

Drawings

The following detailed description of specific embodiments of the present disclosure can be better understood when read in conjunction with the following drawings, where like structure is indicated with like reference numerals and in which:

fig. 1 schematically depicts a fluidized bed reactor according to one or more embodiments described herein;

FIG. 2 schematically depicts another fluidized bed reactor in accordance with one or more embodiments described herein;

FIG. 3 schematically depicts an exemplary chemical treatment system that may utilize the presently described fluidized bed reactor in accordance with one or more embodiments described herein; and is

Fig. 4A-4D depict modeling data for three exemplary embodiments and comparative examples in accordance with one or more embodiments described herein.

It should be understood that the drawings are schematic in nature and do not contain some of the components of the reactor system that are commonly employed in the art, such as, but not limited to, temperature transmitters, pressure transmitters, flow meters, pumps, valves, and the like. Such components are known to be within the spirit and scope of the disclosed embodiments of the invention. However, operational components, such as those described in the present disclosure, may be added to the embodiments described in the present disclosure.

Reference will now be made in detail to various embodiments, some of which are illustrated in the accompanying drawings. Wherever possible, the same reference numbers will be used throughout the drawings to refer to the same or like parts.

Detailed Description

Embodiments are described herein relating to a method of treating a chemical stream in a fluidized bed reactor. In one or more embodiments, the treated chemical stream may be referred to as a feed stream or a reactant stream, and the chemical stream comprising the chemical reaction product may be referred to as a product stream. It is understood that when the conversion of the feed stream is incomplete, the product stream may include various components of the feed stream, which may be typical of many chemical reactions.

The systems and apparatus described herein (e.g., a fluidized bed reactor as described herein) can be used as a processing tool for various fluidized catalytic reactions. The described methods and apparatus can be used in reactions in which a gaseous feed is converted to a gaseous product stream by contact with a solid catalyst, such as a particulate catalyst. For example, hydrocarbons, as well as other chemical feedstocks, can be converted to desired products by using a fluidized bed reactor. Fluidized bed reactors provide a number of uses in industry, including dehydrogenation of paraffins and/or alkylaromatics, cracking of hydrocarbons (i.e., fluid catalytic cracking), chlorination of olefins, oxidation of naphthalene to phthalic anhydride, production of acrylonitrile from propylene, ammonia, and oxygen, Fischer-Tropsch synthesis (Fischer-Tropsch synthesis), polymerization of ethylene, dehydration of hydrocarbons to form light olefins, and some Methanol To Olefins (MTO) reactions.

In accordance with one or more embodiments, some of these reactions (such as, but not limited to, dehydrogenation, cracking, dehydration, and MTO) may form additional moles of gas molecules relative to the moles of feed gas molecules. When the product is gaseous, the partial pressure in the reactor may increase as the reaction proceeds. In some embodiments, such reactions may be represented by the formula aR → bP + cZ, where R represents a reaction species, P represents a product species, Z represents another product species, and a, b, and c each represent the relative amounts of each species utilized in the reaction. For example, a dehydrogenation reaction will produce hydrogen as Z, or Z will be water in a dehydration reaction. When a is less than b + c, additional molecules will be formed by the reaction that is used in the device for which the presently described method is directed. Thus, these reactions form two or more product molecules from each reactant molecule of the reaction. It should be noted that the above equation is merely one exemplary chemical reaction, and it should be understood that other reactions are within the scope of the present disclosure, such as those in which two or more products and/or reactants are present.

According to some embodiments, the fluidized bed reactor described herein can process reaction gases comprising ethane, propane, n-butane, isobutane, and ethylbenzene (e.g., any at least 80 wt%, 90 wt%, 95 wt%, or even 99 wt% or combinations thereof of these reaction gases) to form a product gas comprising ethylene, propylene, butene isomers, and butadiene, isobutylene, styrene, or combinations thereof (e.g., any at least 10 wt%, at least 20 wt%, at least 30 wt%, at least 40 wt%, at least 50 wt%, at least 70 wt%, or even at least 90 wt% or combinations thereof of these reaction gases). For example, gases suitable for dehydrogenation are contemplated herein.

Referring now to fig. 1, an embodiment of a fluidized bed reactor is depicted that can treat a feed stream by contact with a solid catalyst. According to one or more embodiments described herein, the fluidized bed reactor 202 may include a main reactor vessel 250, a transition section 258, and a downstream reactor section 230, such as a riser. The transition section 258 may connect the main reactor vessel 250 with the downstream reactor section 230. As depicted in fig. 1, the main reactor vessel 250 may be disposed below the downstream reactor section 230. Such a configuration may be referred to as an upflow configuration in the fluidized bed reactor 202. The transport riser 430 can supply one or more of reaction gas and catalyst into the fluidized bed reactor 202, and gaseous reactants and products and catalyst can move through the feed distributor 260, through the main reactor vessel 250, through the transition section 258, and into and through the downstream reactor section 230. As depicted in fig. 1, the movement of catalyst and product and reactant gases is upward (depicted by the x-axis). As described herein, "superficial velocity" refers to the apparent velocity of a substance in the entire direction of substance flow (through a plane perpendicular to the x-axis).

As described herein, the primary reactor vessel 250 may comprise a vessel, drum, cylinder, vat, or other vessel suitable for a given chemical reaction. In one or more embodiments, the main reactor vessel 250 may have a substantially circular cross-sectional shape (which represents the cross-sectional view of fig. 1). Alternatively, the cross-section of the main reactor vessel 250 may be non-circular. For example, the main reactor vessel 250 may comprise a cross-sectional shape that is triangular, rectangular, pentagonal, hexagonal, octagonal, elliptical, or other polygonal or curved closed shape, or a combination thereof. As used in this disclosure, the main reactor vessel 250 may generally include a metal frame, and may additionally include a refractory lining or other material for protecting the metal frame and/or controlling the process conditions. As depicted in fig. 1, the main reactor vessel 250 may include a lower reactor portion catalyst inlet port 252 defining a connection of a transport riser 430 to the main reactor vessel 250.

The main reactor vessel 250 may be connected to a transport riser 430, which in operation may provide treated catalyst and/or reactant chemicals in the feed stream. The treated catalyst and/or reactant chemicals may be mixed with a feed distributor 260 contained in the main reactor vessel 250. In one or more embodiments, the feed distributor 260 is operable to distribute the first feed stream and the second feed stream at all shroud distributor speeds of 250ft/s to 80 ft/s. In such embodiments, various feed streams may be utilized while maintaining desired reactor characteristics, such as operating as a fast fluidized, turbulent, or bubble bed reactor in the main reactor vessel 250 and as a dilute phase riser reactor in the downstream reactor section 230. For example, according to one or more embodiments, the primary reactor vessel 250 may utilize a shroud distributor velocity of about 80ft/s for naphtha feeds, and the primary reactor vessel 250 may utilize a shroud distributor velocity of about 250ft/s for propane feeds. In additional embodiments, some of the orifices in the fluidized bed reactor 202 may be closed when naphtha is used as the feed stream. "shroud distributor velocity" refers to the velocity of the gas exiting the distributor, sometimes through the shroud. For example, a suitable dispenser is disclosed in U.S. patent No. 9,370,759, the teachings of which are incorporated herein by reference in their entirety.

As depicted in fig. 1, the main reactor vessel 250 may be connected to the downstream reactor section 230 through a transition section 258. The transition section 258 may taper from the size of the cross-section of the main reactor vessel 250 to the size of the cross-section of the downstream reactor section 230 such that the transition section 258 protrudes inwardly from the main reactor vessel 250 to the downstream reactor section 230.

In one or more embodiments, the downstream reactor section 230 may be generally cylindrical in shape (i.e., having a substantially circular cross-sectional shape), or alternatively non-cylindrical in shape, such as prismatic in shape (having a cross-sectional shape that is triangular), rectangular, pentagonal, hexagonal, octagonal, elliptical, or other polygonal shape, or a curved closed shape, or a combination thereof. As used in this disclosure, the downstream reactor section 230 may generally comprise a metal frame, and may additionally comprise a refractory lining or other material for protecting the metal frame and/or controlling the process conditions.

In some embodiments, such as those in which the main reactor vessel 250 and the downstream reactor section 230 have similar cross-sectional shapes, the transition section 258 may have a frustum shape. For example, for embodiments of the reactor portion 200 comprising the circular cross-section main reactor vessel 250 and the cylindrical downstream reactor section 230, the transition section 258 may have a frustoconical shape. However, it should be understood that a wide variety of shapes for the main reactor vessel 250 are contemplated herein, connecting various shapes and sizes of the transition section 258 with the downstream reactor section 230.

In accordance with one or more embodiments, the main reactor vessel 250 may taper outwardly with respect to the general flow direction of the material in the fluidized bed reactor 202 (i.e., the direction of the x-axis). For example, fig. 1 depicts a linear expanded main reactor vessel 250 having a one-segment conical shape. Although the geometry of the taper may be linear in some embodiments, some embodiments disclosed herein are not linearly tapered, as depicted in fig. 2.

Referring to fig. 1 or 2, the main reactor vessel 250 may include an upstream portion 204 and a downstream portion 206. The upstream portion 204 may be the portion of the main reactor vessel 250 adjacent to the feed distributor 260, and the downstream portion 206 may be the portion of the main reactor vessel 250 adjacent to the transition section 258. The central portion 208 of the main reactor vessel 250 may be positioned equidistant between the upstream portion 204 and the downstream portion 206 (based on the height of the main reactor vessel 250 in the x-direction). In general, the upstream portion 204 of the main reactor vessel 250 may have a smaller cross-sectional area than the downstream portion 206 of the main reactor vessel 250, defining a taper in the cross-sectional geometry of the main reactor vessel 250. As described herein, unless expressly stated otherwise, "cross-sectional area" refers to the area of a cross-section of a portion of a reactor part in a plane substantially orthogonal to the average flow direction of reactants and/or products (i.e., a plane perpendicular to the x-axis in fig. 1). For example, in fig. 1, the cross-sectional areas of the main reactor vessel 250, the transition section 258, and the downstream reactor section 230, or any portion thereof, are in the direction of the plane defined by the horizontal direction and into the page (orthogonal to the direction of fluid movement, i.e., vertically upward in fig. 1). In additional embodiments, the upstream portion 204 may have a smaller cross-sectional area than the central portion 208, and the central portion 208 may have a smaller cross-sectional area than the downstream portion 206. Thus, in some embodiments, the cross-sectional area of the central portion 208 may be less than the cross-sectional area of the downstream portion 206 and greater than the cross-sectional area of the upstream portion 204.

In one or more embodiments, the downstream portion 206 of the main reactor vessel 250 can have a cross-sectional area that is 1.2 to 1.8 times (e.g., 1.2 to 1.4, 1.4 to 1.6, or 1.6 to 1.8) the cross-sectional area of the upstream portion 204 of the main reactor vessel 250. In accordance with one or more embodiments, the height of the main reactor vessel 250 may be 2ft to 12ft (e.g., 2ft to 4ft, 4ft to 6ft, 6ft to 8ft, 8ft to 10ft, or 10ft to 12 ft).

In one or more embodiments, the taper of the main reactor vessel 250 may be described by a slope at a particular portion of the main reactor vessel 250. For example, the main reactor vessel 250 may have a measurable slope (i.e., half of the width change divided by the height change) at the upstream portion 204, the downstream portion 206, and the central portion 208. That is, the slope of the constant sectional shape is 0. In the embodiment of fig. 1, the slope may be substantially constant throughout the main reactor vessel 250, such as with respect to the upstream portion 204, the downstream portion 206, and the central portion 208. Such a substantially constant slope corresponds to a linear profile of the main reactor vessel 250. In additional embodiments, such as the embodiment of fig. 2, the main reactor vessel 250 may include a slope at or near the downstream portion 206 that is less than the slope at or near the upstream portion 204 of the main reactor vessel 250. In such embodiments, the slope at the central portion 208 may be less than the slope of the upstream portion 204 and greater than the slope of the downstream portion 206. As explained in the examples below, the slope and relative shape of the main reactor vessel 250 may affect the superficial gas velocity and/or the suspended density at a local location within the main reactor vessel 250.

As depicted in fig. 1 and 2, the inwardly tapering transition section 258 and the outwardly tapering main reactor vessel 250 form a geometric configuration of the fluidized bed reactor 202, wherein the portion of the fluidized bed reactor 202 having the largest cross-sectional area is located at the junction of the transition section 258 and the main reactor vessel 250 (i.e., at or near the downstream portion 206 of the main reactor vessel 250). The downstream reactor section 230 may have a smaller cross-sectional area than the transition section 258 and the main reactor vessel 250. In one or more embodiments, the downstream reactor section 230 may have a smaller cross-sectional area than the upstream portion 204 of the main reactor vessel 250.

According to embodiments described herein, at least one use of the outwardly tapering main reactor vessel 250 is its effect on the superficial gas velocity and/or suspension density at a local location in the main reactor vessel 250. It may be desirable to have a relatively constant gas superficial velocity and/or suspension density along the height of the main reactor vessel 250. For example, it may be desirable to have a similar gas superficial velocity and/or suspension density at two or more of the upstream portion 204 as the downstream portion 206 or the central portion 208.

In accordance with one or more embodiments, the superficial velocity of the gas in the fluidized bed reactor at the downstream portion 206 of the primary reactor vessel 250 can be less than or equal to 140% (e.g., less than or equal to 130%, less than or equal to 120%, less than or equal to 110%, less than or equal to 100%, less than or equal to 90%, or even less than or equal to 80%) of the superficial velocity of the gas in the fluidized bed reactor at the upstream portion 204 of the primary reactor vessel 250. For example, in additional embodiments, the superficial velocity of the gas in the fluidized bed reactor at the downstream portion 206 of the primary reactor vessel 250 may be 60% to 140%, 70% to 130%, 80% to 120%, or even 90% to 110% of the superficial velocity of the gas in the fluidized bed reactor at the upstream portion 204 of the primary reactor vessel. In additional embodiments, the superficial velocity of the gas in the fluidized bed reactor at the downstream portion 206 of the main reactor vessel 250, at the upstream portion 204 of the main reactor vessel 250, or both, may be less than or equal to 140% (e.g., less than or equal to 130%, less than or equal to 120%, less than or equal to 110%, less than or equal to 100%, less than or equal to 90%, or even less than or equal to 80%) of the superficial velocity of the gas in the fluidized bed reactor at the central portion 208 of the main reactor vessel 250. For example, the superficial velocity of the gas in the fluidized bed reactor at the downstream portion 206 of the primary reactor vessel 250, at the upstream portion 204 of the primary reactor vessel 250, or both, may be 60% to 140%, 70% to 130%, 80% to 120%, or even 90% to 110% of the superficial velocity of the gas in the fluidized bed reactor at the central portion 208 of the primary reactor vessel 250.

According to additional embodiments, the suspension density in the fluidized bed reactor at the downstream portion 206 of the main reactor vessel 250 may be greater than or equal to 25% (e.g., greater than or equal to 35%, greater than or equal to 50%, or even greater than or equal to 75%) of the suspension density in the fluidized bed reactor at the upstream portion 204 of the main reactor vessel 250. For example, in additional embodiments, the suspension density in the fluidized bed reactor at the downstream portion 206 of the main reactor vessel 250 may be 25% to 175%, 50% to 150%, 80% to 120%, or even 90% to 110% of the suspension density in the fluidized bed reactor at the upstream portion 204 of the main reactor vessel. In additional embodiments, the suspension density in the fluidized bed reactor at the central portion 208 of the main reactor vessel 250 may be greater than or equal to 40%, 50%, 60%, 70%, or even 80% of the suspension density in the fluidized bed reactor at the upstream portion 204 of the main reactor vessel 250.

As described herein, the superficial velocities at the upstream portion 204 and the downstream portion 206 of the main reactor vessel 250 can be determined by utilizing known equations (e.g., the ideal gas law with measurable characteristics of the stream within the fluidized bed reactor 202). The superficial velocities at the upstream and downstream portions 204, 206 of the main reactor vessel 250 may be determined using the corresponding temperatures and pressures of the upstream and downstream portions 204, 206 of the main reactor vessel 250, as well as the mass flow rates and gas compositions entering and exiting the main reactor vessel 250. For example, temperature and pressure probes may be used within the fluidized bed reactor, and the slip stream at the height along the fluidized bed reactor may be determined as the gas composition at a particular height. In addition, the suspension density can be determined by comparing the pressure at the two reactor heights and the applied known equation. It should be appreciated that because two measurements may be required to determine the suspended density, the suspended density at the upstream portion 204 may be measured by data collected from the area adjacent the distributor and one foot downstream (e.g., above) the distributor. Similarly, the suspended density at the downstream portion 206 may be measured by data collected from the area adjacent the transition section 258 and one foot upstream (e.g., below) of the transition section 258.

In one or more embodiments, based on shape, size, and other processing conditions (e.g., temperature and pressure in the main reactor vessel 250 and the downstream reactor section 230), the main reactor vessel 250 may be operated in an isothermal or near isothermal manner (e.g., in a fast fluidized, turbulent, or bubbling bed reactor), while the downstream reactor section 230 may be operated in a more plug-flow manner (e.g., in a dilute phase riser reactor). For example, the fluidized bed reactor 202 of fig. 1 may comprise a main reactor vessel 250 operating as a fast fluidized, turbulent, or bubbling bed reactor and a downstream reactor section 230 operating as a dilute phase riser reactor, with the result that the average catalyst and gas streams move upward simultaneously. As used herein, the term "average flow rate" refers to the net flow rate, i.e., the total upward flow rate minus the counter-current flow rate or reverse flow rate, which is typical behavior of a typical fluidized particle. As described herein, a "fast fluidized" reactor may refer to a reactor that utilizes a fluidized flow regime, where the superficial velocity of the gas phase is greater than the choking velocity and may be semi-dense in operation. As described herein, a "turbulent" reactor may refer to a fluidized flow regime, where the superficial velocity is less than the choking velocity and is denser than the fast fluidized flow regime. As described herein, a "bubbling bed" reactor can refer to a fluidized flow regime, wherein well-defined bubbles in a highly dense bed exist in two distinct phases. "choke rate" means the minimum rate required to maintain solids in dilute phase mode in a vertical transfer line. As described herein, "dilute phase riser" may refer to a riser reactor operating at a transport velocity where the gas and catalyst have approximately the same velocity in the dilute phase.

In one or more embodiments, the pressure in the fluidized bed reactor 202 may be in the range of 6.0 to 44.7 pounds per square inch absolute (psia, about 41.4 kilopascals (kPa), to about 308.2kPa), although in some embodiments, a more narrow range of choice may be employed, such as 15.0psia to 35.0psia (about 103.4kPa to about 241.3 kPa). For example, the pressure may be 15.0psia to 30.0psia (about 103.4kPa to about 206.8kPa), 17.0psia to 28.0psia (about 117.2kPa to about 193.1kPa), or 19.0psia to 25.0psia (about 131.0kPa to about 172.4 kPa). A unit conversion herein from a standard (non-SI) to a metric (SI) expression includes "about" to indicate rounding that may exist in the metric (SI) expression due to the conversion.

In additional embodiments, the Weight Hourly Space Velocity (WHSV) for the disclosed process can range from 0.1 pounds (lb) to 100lb chemical feed per hour (hr) per lb catalyst (lb feed/hr/lb catalyst) in the reactor. For example, where the reactor comprises a main reactor vessel 250 acting as a fast fluidized, turbulent, or bubble bed reactor and a downstream reactor section 230 acting as a riser reactor, the superficial gas velocity may be in the range of 2 feet per second (ft/s, about 0.61 meters per second, m/s) to 80ft/s (about 24.38m/s) in the main reactor vessel 250, such as 2ft/s (about 0.61m/s) to 10ft/s (about 3.05m/s), and in the range of 30ft/s (about 9.14m/s) to 70ft/s (about 21.31m/s) in the downstream reactor section 230. In additional embodiments, an entirely riser-type reactor configuration may be operated at a single high superficial gas velocity, e.g., at least 30ft/s (about 9.15m/s) throughout in some embodiments.

In additional embodiments, the ratio of catalyst to feed stream in the fluidized bed reactor 202 may range from 5 to 100 on a weight/weight (w/w) basis. In some embodiments, the ratio may be in the range of 10 to 40, such as 12 to 36, or 12 to 24.

In additional embodiments, the catalyst flux in the main reactor vessel 250 may be 1 pound per square foot-second (1 b/ft)²S) (about 4.89 kg/m)²-s) to 20lb/ft²S (to about 97.7kg/m2-s) and the catalyst flux in the downstream reactor section 230 may be 10lb/ft²S (about 48.9kg/m2-s) to 100lb/ft²S (about 489kg/m 2-s).

According to additional embodiments, the fluidized bed reactor 202 may include internal structures such as those described in U.S. publication No. 2016/0375419, the contents of which are incorporated by reference in their entirety.

Referring now to fig. 3, an exemplary reactor system 102 incorporating a fluidized bed reactor (such as the fluidized bed reactor of fig. 1 or 2) is schematically depicted and is suitable for use with the methods described herein. It should be understood that the system of fig. 3 is merely an exemplary system and that other systems may be used with the fluidized bed reactors described herein.

The reactor system 102 generally includes a plurality of system components, such as a reactor section 200 and/or a catalyst treatment section 300. As used herein in the context of fig. 2, reactor portion 200 generally refers to the portion of reactor system 102 in which the primary treatment reaction is conducted, such as a dehydrogenation reaction, a cracking reaction, a dehydration reaction, or a methanol to olefins reaction, for example, to form light olefins. The reactor section 200 comprises a fluidized bed reactor 202, which may include a downstream reactor section 230 and a main reactor vessel 250. In accordance with one or more embodiments, as depicted in fig. 3, the reactor portion 200 can additionally include a catalyst separation section 210 for separating catalyst from chemical products formed in the fluidized bed reactor 202. Also, as used herein, the catalyst treatment section 300 generally refers to the portion of the reactor system 102 in which the catalyst is treated in some manner (e.g., by combustion). The catalyst treatment section 300 may comprise a combustion chamber 350 and a riser 330, and may optionally comprise a catalyst separation section 310. In some embodiments, the catalyst may be regenerated by burning contaminants (e.g., coke) in the catalyst treatment section 300. In additional embodiments, the catalyst may be heated in the catalyst treatment section 300. Supplemental fuel may be used to heat the catalyst in the catalyst treatment section 300 if coke or another combustible material is not formed on the catalyst, or if the amount of coke formed on the catalyst is not sufficient to burn to heat the catalyst to the desired temperature. In one or more embodiments, the catalyst separation section 210 can be in fluid communication with the combustor 350 (e.g., through the standpipe 426), and the catalyst separation section 310 can be in fluid communication with the primary reactor vessel 250 (e.g., through the standpipe 424 and the transport riser 430).

As described with respect to fig. 3, the feed stream may enter the transport riser 430, and the product stream may exit the reactor system 102 via the conduit 420. In accordance with one or more embodiments, the reactor system 102 can be operated by feeding a chemical feed (e.g., in a feed stream) and a fluidized catalyst into the main reactor vessel 250. The chemical feeds contact the catalyst in the main reactor vessel 250 and each chemical feed flows upward into and through the downstream reactor section 230 to produce a chemical product. From the downstream reactor section 230, the chemical product and the catalyst may flow out to a separation device 220 in the catalyst separation section 210, where the catalyst is separated from the chemical product, which is transported out of the catalyst separation section 210. The separated catalyst is transferred from the catalyst separation section 210 to the combustor 350. In the combustion chamber 350, the catalyst may be treated by, for example, combustion. For example, but not limiting of, the catalyst may be decoked and/or supplemental fuel may be combusted to heat the catalyst. The catalyst then flows from the combustion chamber 350 and through the riser 330 to a riser terminal separator 378, where the gaseous and solid components from the riser 330 are at least partially separated. The vapors and remaining solids are sent to a secondary separation device 320 in a catalyst separation section 310 where the remaining catalyst is separated from gases from the catalyst treatment (e.g., gases emitted by burning spent catalyst or supplemental fuel). The separated catalyst is then transferred from the catalyst separation section 310 (via portion 312) to the main reactor vessel 250 via standpipe 424 and transfer riser 430, where it is further used for catalytic reactions. Thus, catalyst may be circulated between reactor section 200 and catalyst treatment section 300 during operation. In general, the treated chemical stream, including the feed stream and the product stream, can be gaseous, and the catalyst can be a fluidized particulate solid.

Catalyst entering the primary reactor vessel 250 through the transport riser 430 may be transferred to the transport riser 430 through the standpipe 424 and thus arrive from the catalyst treatment section 300. In some embodiments, the catalyst may come directly from the catalyst separation section 210 via the standpipe 422 and into the transport riser 430 where the catalyst enters the main reactor vessel 250. This catalyst may be slightly deactivated, but in some embodiments may still be suitable for reaction in the main reactor vessel 250. As used herein, "deactivated" may refer to catalyst that is contaminated with substances such as coke or is at a lower temperature than desired. Regeneration may remove contaminants such as coke, raise the temperature of the catalyst, or both.

In operation, catalyst may move upward through the downstream reactor section 230 (from the main reactor vessel 250) and into the separation device 220. In some embodiments, the downstream reactor section 230 may include an inner portion 234 (i.e., within the catalyst separation section 210) and an outer portion 232. The separated vapors may be removed from the reactor system 102 via conduit 420 at the gas discharge port 216 of the catalyst separation section 210. According to one or more embodiments, the separation device 220 may be a cyclonic separation system (cyclonic separation system), which may include two or more cyclonic separation stages. In embodiments where the separation device 220 comprises more than one cyclonic separation stage, the first separation device into which the fluidising stream enters is referred to as the primary cyclonic separation device. The fluidized effluent from the primary cyclonic separating apparatus may enter a secondary cyclonic separating apparatus for further separation. The primary cyclonic separating apparatus may comprise, for example, a primary cyclone and systems commercially available under the names VSS (available from UOP), LD2 (available from schwann (Stone and Webster)) and RS2 (available from schwann). Primary cyclones are described, for example, in U.S. patent nos. 4,579,716, 5,190,650 and 5,275,641, each of which is incorporated herein by reference in its entirety. In some separation systems that utilize a primary cyclone as the primary cyclonic separation device, one or more additional sets of cyclones are employed, such as secondary and tertiary cyclones, for further separating the catalyst from the product gas. It will be appreciated that any primary cyclonic separating apparatus may be used in embodiments of the invention.

According to one or more embodiments, after separation from the vapors in the separation device 220, the catalyst may generally move through the stripper column 224 to the catalyst discharge port 222, where the catalyst is transferred out of the reactor portion 200 and into the catalyst treatment portion 300 via the standpipe 426. Optionally, the catalyst may also be transferred directly back into the main reactor vessel 250 via the standpipe 422. Alternatively, the catalyst may be premixed with the treated catalyst in the transport riser 430.

In accordance with one or more embodiments, operating a chemical process, such as in the reactor system 102, can include recycling catalyst for the chemical process by transferring the catalyst from the fluidized bed reactor 202 to a regeneration unit (such as the combustor 350 of the embodiment of fig. 2), treating the respective catalyst in the regeneration unit, and transferring the first catalyst from the regeneration unit to the fluidized bed reactor 202.

Referring now to the catalyst treatment section 300, as depicted in fig. 3, the combustion chamber 350 of the catalyst treatment section 300 may include one or more lower reactor section inlet ports 352 and may be in fluid communication with the riser 330. The combustor 350 may be in fluid communication with the catalyst separation section 210 via a standpipe 426, which may supply spent catalyst from the reactor section 200 onto the catalyst treatment section 300 for regeneration. The combustion chamber 350 may include an additional lower reactor section inlet port 352, with an air inlet 428 connected to the combustion chamber 350. The air inlet 428 may supply a reactant gas that may react with the spent catalyst or the supplemental fuel to at least partially regenerate the catalyst. For example, the catalyst may coke in the main reactor vessel 250 after reaction, and the coke may be removed from the catalyst (i.e., regenerated catalyst) by a combustion reaction. For example, an oxidant (e.g., air) may be fed into the combustion chamber 350 via the air inlet 428. Alternatively or additionally, supplemental fuel may be injected into the combustion chamber 350, such as when coke is not formed on the catalyst, which may combust to heat the catalyst. After combustion, the treated catalyst may be separated in the catalyst separation section 310 and delivered back to the reactor section 200 via standpipe 424.

Examples of the invention

The following examples are illustrative in nature and should not be used to limit the scope of the present application.

Models were constructed to calculate the expected suspension density, conversion and superficial gas velocity in the fluidized bed reactor as a function of reactor geometry. The fluidized bed reactor has a geometry similar to that of fig. 1 or fig. 2, but the main reactor vessel has a specified taper. In particular, the suspension density, conversion and superficial gas velocity are calculated as a function of the height of the various reactor geometries with outwardly tapering walls. The calculation of fluid dynamics is based on the methods disclosed in Kunii, D, and Levenspiel, O, Entrainment of solids from fluidized beds (Entrainment of solids from fluidized beds), Powder Technology (Powder Technology), 61,1990, 193-206. The model simulates a reaction in which each reactant gas molecule forms two product gas molecules. The gas velocity at the bottom of the reactor was 3.5 ft/s.

In the first test example (comparative example), a constant reactor diameter was tested. Lines 502 in fig. 4A, 4B and 4C depict the suspension density, conversion and superficial gas velocity, respectively, as a function of reactor height. Fig. 4D depicts in line 502 the constant diameter of the main reactor vessel. The results show that with a constant diameter, the suspension density decreases significantly with increasing height, while the apparent gas velocity increases significantly with increasing height.

The second test embodiment is depicted in fig. 4A, 4B, 4C, and 4D as line 504. In this embodiment, the geometry of the main reactor vessel allows for a constant superficial gas velocity as a function of reactor height. In such a configuration, the suspension density remains greater at all heights compared to the constant diameter comparative example. For such embodiments, the geometric profile of the main reactor vessel can be seen in fig. 4D, where the slope of the profile increases with increasing height, and the upper half of the main reactor vessel is almost cylindrical.

The third test embodiment is depicted in fig. 4A, 4B, 4C, and 4D as line 506. In this example, the geometry of the main reactor vessel was linear, with the diameter at the downstream portion of the main reactor vessel being equal to the diameter of the second test example. In such a configuration, both the suspension density and the apparent gas velocity are more desirable than in the comparative example.

The fourth test embodiment is depicted in fig. 4A, 4B, 4C, and 4D as line 508. In this embodiment, the geometry of the main reactor vessel is linear similar to that of the third exemplary embodiment, but wider at the downstream portion, so that a conversion of 0.5 can be achieved. This example also increased the suspension density and reduced the apparent gas velocity compared to the comparative example.

The residence times and average densities for the first, second, third, and fourth exemplary embodiments were calculated and are depicted in table 1 below.

Table 1:

as shown in fig. 4B, a reactor designed with a conical geometry has a higher conversion at a lower height. Thus, it is contemplated that the tapered geometric design disclosed herein may be used to build and operate a reactor that is shorter (relative to the total feed rate) than conventional reactors. In addition, without being bound by theory, it is believed that a shorter reactor may result in a shorter residence time, which may result in a higher selectivity for the dehydrogenation reaction due to the reduction in the thermal cracking reaction.

It will be apparent to those skilled in the art that various modifications and variations can be made in the present invention without departing from the spirit and scope of the invention. Since modifications combinations, sub-combinations and variations of the disclosed embodiments incorporating the spirit and substance of the invention may occur to persons skilled in the art, the invention should be construed to include everything within the scope of the appended claims and their equivalents.