Process for the preparation of olefins using a new catalyst and circulating fluid bed technology

Abstract

The present application relates to a circulating fluidized bed process for producing olefins from hydrocarbon feedstock by applying the catalyst for producing olefins to a high-speed fluidized bed, which can effectively increase the production of olefins, in particular, can keep the high output rate of propylene. That is, the output rate of the circulating fluidized bed process according to the present application is higher than that of the conventional process, so that the incremental profit per unit of feedstock can be increased. Moreover, the use of the continuous reaction-regeneration fluidized bed reactor can make the regeneration process by means of direct heating, only need to be equipped with less air flow and smaller air compressor, so compared with the conventional process, the fuel consumption can be reduced by 10 to 15%, the compressor energy required for product separation and catalyst regeneration can be reduced by 15 to 20%, thereby reducing the overall production cost.

Description

Technical Field

[0001] This invention relates to a method for preparing olefins using a circulating fluidized bed process. Background Technology

[0002] Olefins are widely used in the petrochemical industry, such as in ethylene and propylene. Generally, olefins are obtained from the pyrolysis of naphtha. However, with the surge in shale gas production and the price competitiveness of gaseous feedstocks compared to naphtha, ethane pyrolysis has become increasingly common. Consequently, the supply of ethylene has increased, but propylene production has relatively slowed, leading to an imbalance between propylene supply and demand. Therefore, to regulate the supply and demand of propylene, "on-purpose propylene" technology is becoming increasingly popular. This technology, which involves the production of specialty propylene, utilizes catalysts and a dehydrogenation process involving lower hydrocarbons to produce propylene.

[0003] Among the existing commercial propane dehydrogenation (PDH) processes, the most representative are the fixed-bed reactor and the moving-bed reactor.

[0004] In contrast, there is currently no commercial precedent for PDH (Fast-fluidized Propane Dehydrogenation) technology using high-speed fluidized bed (hereinafter referred to as fluidized bed) reactors.

[0005] The biggest difference between the fixed-bed reactor and the fluidized-bed reactor is the contact time between the catalyst and the reactant (propane). That is, the fluidized-bed reactor injects propane and catalyst together at a very high speed for reaction, and then the catalyst flows into the regeneration section, while the product flows into the separation section.

[0006] In the circulating fluidized bed process for producing olefins from hydrocarbon feedstock mixtures, the high conversion and selectivity allow for careful consideration of the operating conditions of the riser, which primarily performs the dehydrogenation reaction, to facilitate the selective production of olefins such as ethylene and propylene. Specifically, the fluidization and reaction phenomena within the riser can be more readily understood through the following theoretical examination, which will be described in more detail below.

[0007] like Figure 1 As shown, if gas flows into a container filled with a solid catalyst from below, the particles are fluidized. If the minimum fluidization velocity is reached, the fluidization flow region is generally divided into 5 regions.

[0008] Specifically, they are called the Minimum Fluidization Regime, Bubbling Fluidization Regime, Slugging Fluidization Regime, Turbulent Fluidization Regime, and Lean Phase Fluidization with Pneumatic Transport Regime, and the particle motion characteristics differ in each region.

[0009] Therefore, in processes using fluidized bed reactors, fluidized flow zones that conform to the characteristics of each process are formed and operations are carried out.

[0010] Figure 2 The diagram illustrates the change in catalyst volume percentage within the reactor as the riser height, i.e., different fluidized flow regions, changes, confirming that the catalyst volume percentage within the reactor also changes with the fluidized flow region. However, as demonstrated in the dehydrogenation reaction process using a contact fluidized bed, the catalyst volume percentage significantly impacts process performance when the reaction is carried out with entrained catalyst. Consequently, controlling the catalyst volume percentage within the reactor to determine the process operating conditions of the fluidized flow region plays a crucial role in the reaction outcome.

[0011] To determine the fluidization flow region of the riser in this circulating fluidization process, the following factors should be considered, such as catalyst size, catalyst circulation rate, feed-to-catalyst ratio, and catalyst strength.

[0012] In addition, the following factors that directly affect the dehydrogenation reaction should be considered, such as reaction temperature, reaction endothermic heat, reaction time, and catalyst deactivation caused by coke formation.

[0013] The goal of the currently developing FPDH process is to reduce the catalyst residence time to below 10 seconds. A shorter catalyst residence time results in a faster propane delivery rate, allowing the catalyst to regenerate and re-participate in the reaction. Therefore, when developed into a commercial process, this would significantly increase propylene production compared to stationary bed processes.

[0014] However, due to the relatively short contact time between the catalyst and propane, the catalyst efficiency becomes very important. That is, maximizing the two efficiency metrics of the catalyst, namely selectivity and conversion, is crucial.

[0015] Furthermore, since the currently used propane dehydrogenation technology is based on precious metal catalysts or consists of discontinuous processes, the following problems exist in propylene production operations: reactor blockage due to excessive activity of precious metal catalysts (generating coke), or valve sequence problems in fixed-bed reactors, etc.

[0016] Furthermore, due to the reversible nature of the hydrogen-induced reaction, the propane conversion rate in the propane dehydrogenation reaction is thermodynamically limited. To overcome this problem, most processes use external oxidants such as oxygen, halogens, sulfur compounds, carbon dioxide, or water vapor to convert hydrogen into water.

[0017] Therefore, in order to achieve efficient large-scale production of propylene, it is necessary to solve the problems existing in the above-mentioned continuous process and develop a new propane dehydrogenation process that does not use oxidants but uses direct dehydrogenation catalysts, thereby saving production costs.

[0018] In catalysts used for propane dehydrogenation, noble metal catalysts react via direct dehydrogenation, i.e., hydrogen adsorption to active sites. However, due to the incompleteness of active sites caused by electron mobility, the mechanism of transition metal oxides has not yet been fully elucidated.

[0019] Due to the above reasons, the most commonly used catalysts for PDH are Pt-Sn, VOx, and CrOx catalysts. Although CrOx catalysts are excellent in terms of propane conversion and selectivity, their use is limited due to environmental pollution, health hazards, and the difficulty in controlling the oxidation reaction in the early stages. Platinum catalysts have good selectivity, but they are expensive and produce coke very quickly, requiring precise control. Furthermore, the inherent activity of the catalyst changes when combined with co-catalyst components such as Sn and other metals, and because Sn is harmful to the environment, there is a continuous need to develop new multi-component catalysts for platinum catalysts.

[0020] However, in fluidized bed reactors, the catalyst residence time is less than 10 seconds. Therefore, in the initial stage of the reaction, the transition metal oxide catalyst encounters significant problems when propane is completely oxidized to CO2. To ensure propylene selectivity, the oxidation level of the transition metal must be controlled. Consequently, when implementing the FPDH process using a fluidized bed reactor, selectivity is extremely important due to the short reaction time.

[0021] Therefore, this invention studies a method for preparing olefins using a circulating fluidized bed process that offers better economic benefits and production efficiency than traditional preparation processes. During this process, a catalyst with excellent conversion rate and selectivity for olefin preparation was applied to the circulating fluidized bed process, and the reaction conditions were improved, thereby developing a more efficient method for olefin preparation and completing the invention. Summary of the Invention

[0022] The purpose of this invention is to provide a fluidized bed olefin preparation method that offers better economic benefits and production efficiency than traditional processes.

[0023] The olefin preparation method using the circulating fluidized bed process of this invention includes:

[0024] (a) The step of feeding a hydrocarbon mixture containing propane and a dehydrogenation catalyst into a riser that serves as a high-speed fluidization zone to cause a dehydrogenation reaction;

[0025] (b) The step of separating the catalyst from the product of the dehydrogenation reaction, i.e., the propylene mixture;

[0026] (c) Stripping step, which removes unseparated hydrocarbons remaining in the catalyst separated in step (b);

[0027] (d) The step of continuously regenerating the catalyst by mixing the catalyst stripped in step (c) with oxygen-containing gas;

[0028] (e) Recycling the catalyst regenerated in step (d) back to step (a) for resupply in the riser; and

[0029] (f) The step of cooling, compressing, and separating the reaction product separated in step (b), i.e., the propylene mixture, to prepare propylene.

[0030] The boron-modified dehydrogenation catalyst is impregnated with an active metal containing cobalt and platinum on an alumina support.

[0031] Preferably, the high-speed fluidization region is in a normal state, that is, the gas flow velocity in the riser is maintained higher than that in the turbulent fluidization region and lower than that in the dilute phase air transport fluidization region, so that the catalyst continuously and quantitatively flows into the riser. It is a fluidization region with a dense phase at the bottom of the riser and a dilute phase at the top of the riser.

[0032] Preferably, in the high-speed fluidization region,

[0033] a) Maintain the gas velocity at a rate above the required gas velocity for the catalyst mist continuously flowing in from the bottom of the riser to flow smoothly out of the top of the riser.

[0034] at the same time,

[0035] b) Adjust the gas flow rate and catalyst inflow rate to keep the difference in catalyst volume percentage between the two locations between 0.02 and 0.04.

[0036] Preferably, the difference in catalyst volume percentage between the lower 1 / 4 and the 3 / 4 of the riser tube is maintained at 0.02 to 0.04.

[0037] Preferably, the propane content of the hydrocarbon mixture is 90% by weight or more, and more preferably, it is 95% or more.

[0038] The catalyst temperature at the inlet of the riser is 550 to 700°C, more preferably 620 to 680°C, and preferably, the temperature at the lower part of the riser is kept higher than the temperature at the upper part of the riser.

[0039] Preferably, the pressure of the riser tube is 24 to 26 psig.

[0040] The residence time of the hydrocarbon mixture is 1 to 6 seconds, more preferably 2.5 to 4.5 seconds, to facilitate the dehydrogenation reaction within the riser.

[0041] In step (e), the weight ratio of the catalyst, which is again fed to the lower part of the riser, to the hydrocarbon mixture is 10 to 50, more preferably 30 to 45.

[0042] The gas space velocity (WHSV, h) relative to the weight of the catalyst flowing into the riser is... -1 Preferably, the value is 2 to 40, and more preferably, it is 7 to 13.

[0043] Preferably, the alumina support prepared at a temperature of 550 to 850°C above the dehydrogenation reaction temperature has a γ to θ phase, and within this range, the surface area is 100 to 300 m². 2 / g.

[0044] Preferably, the pore size of the alumina carrier is 0.1 to 5 μm.

[0045] Preferably, the pore volume of the alumina support is 0.4 to 0.6 cm³. 3 / g.

[0046] Preferably, the catalyst comprises 0.1 to 2 wt% boron, 2 to 10 wt% cobalt, and 0.001 to 0.05 wt% platinum.

[0047] Preferably, the catalyst has an average size of 20 to 200 micrometers, and more preferably, 60 to 120 micrometers.

[0048] This invention relates to a circulating fluidized bed process for preparing olefins from hydrocarbon feedstocks by applying the catalyst described herein to a high-speed fluidized bed. Because the catalyst has good contact with the feedstock in the high-speed fluidized zone, the conversion rate can be improved, the yield increased, and the production of olefins more efficiently increased.

[0049] In other words, the circulating fluidized bed process described in this invention has a higher yield than commonly used processes in conventional technologies, thus increasing the incremental profit per unit of raw material. Furthermore, the use of a continuous reaction-regeneration fluidized bed reactor allows the regeneration process to utilize direct heating, requiring only a smaller airflow rate and air compressor. Therefore, compared to conventional processes, fuel consumption can be reduced by 10% to 15%, and the compressor energy required for product separation and catalyst regeneration can be reduced by 15% to 20%, thereby lowering overall production costs. Attached Figure Description

[0050] Figure 1 The changes in the internal characteristics of the fluidized bed in the fluidized flow region under normal gas velocity variations are shown.

[0051] Figure 2 The volume percentage of catalyst inside the fluidized bed is shown based on the riser height.

[0052] Figure 3 This is a schematic diagram of the circulating fluidized bed process used in this invention.

[0053] Figure 4 This is a schematic diagram showing the conversion and selectivity test results of the CPB / KRICT-2 catalyst under temperature changes according to the present invention.

[0054] Figure 5 This is a schematic diagram showing the conversion and selectivity test results of the CPB / KRICT-2 catalyst with varying catalyst / propane ratio according to the present invention.

[0055] Figure 6 This is a schematic diagram showing the conversion and selectivity test results of the CPB / Puralox catalyst under temperature changes according to the present invention.

[0056] Figure 7 This is a schematic diagram showing the conversion and selectivity test results of the CPB / Puralox catalyst with varying catalyst / propane ratio according to the present invention.

[0057] Figure 8 This is a schematic diagram showing the conversion and selectivity test results of the CPB / Puralox catalyst residence time (RT) variation according to the present invention.

[0058] Figure 9This is a schematic diagram showing the conversion and selectivity test results during the continuous process of the CPB / Puralox catalyst according to the present invention.

[0059] Figure 10 This is a schematic diagram showing the conversion and selectivity test results during the continuous process of the CPB / KRICT-2 catalyst according to the present invention.

[0060] Label Explanation

[0061] 1: Riser

[0062] 2: Stripper

[0063] 3: Regenerator

[0064] 11: Hydrocarbon feedstock transportation pipeline

[0065] 13: Regenerator Stand Pipe

[0066] 15: Gas reaction products

[0067] 16: Stripping Steam Delivery Pipeline

[0068] 17: Stripper Stand Pipe

[0069] 18: Stripper Slide Valve

[0070] 19: Flue Gas

[0071] 20: Oxygen-containing gases such as air

[0072] 21: Regenerator Slide Valve Detailed Implementation

[0073] The olefin preparation method using the circulating fluidized bed process of this invention includes:

[0074] (a) The step of feeding a hydrocarbon mixture containing propane and a dehydrogenation catalyst into a riser that serves as a high-speed fluidization zone to cause a dehydrogenation reaction;

[0075] (b) The step of separating the catalyst from the product of the dehydrogenation reaction, i.e., the propylene mixture;

[0076] (c) Stripping step, which removes unseparated hydrocarbons remaining in the catalyst separated in step (b);

[0077] (d) The step of continuously regenerating the catalyst by mixing the catalyst stripped in step (c) with oxygen-containing gas;

[0078] (e) Recycling the catalyst regenerated in step (d) back to step (a) for resupply in the riser; and

[0079] (f) The step of cooling, compressing, and separating the reaction product separated in step (b), i.e., the propylene mixture, to prepare propylene.

[0080] The boron-modified dehydrogenation catalyst is impregnated with an active metal containing cobalt and platinum on an alumina support.

[0081] Preferably, the high-speed fluidization region is in a normal state, that is, the gas flow velocity in the riser is maintained higher than that in the turbulent fluidization region and lower than that in the dilute phase air transport fluidization region, while the catalyst continuously and quantitatively flows into the riser, forming a fluidization region with a dense region at the bottom of the riser and a dilute region at the top of the riser.

[0082] [Implementation Forms of the Invention]

[0083] Hereinafter, preferred embodiments of the present invention will be described with reference to the accompanying drawings. However, the embodiments described herein can be modified in many other ways, and the scope of the present invention is not limited to the following embodiments.

[0084] In describing this embodiment, the same components will be referred to by the same names and symbols, and therefore, repeated additional descriptions will be omitted below. The accompanying drawings will not be to scale.

[0085] The olefin preparation method using the circulating fluidized bed process of this invention includes:

[0086] (a) The step of feeding a hydrocarbon mixture containing propane and a dehydrogenation catalyst into a riser that serves as a high-speed fluidization zone to cause a dehydrogenation reaction;

[0087] (b) The step of separating the catalyst from the product of the dehydrogenation reaction, i.e., the propylene mixture;

[0088] (c) Stripping step, which removes unseparated hydrocarbons remaining in the catalyst separated in step (b);

[0089] (d) The step of continuously regenerating the catalyst by mixing the catalyst stripped in step (c) with oxygen-containing gas;

[0090] (e) Recycling the catalyst regenerated in step (d) back to step (a) for resupply in the riser; and

[0091] (f) The step of cooling, compressing, and separating the reaction product separated in step (b), i.e., the propylene mixture, to prepare propylene.

[0092] The dehydrogenation catalyst is impregnated in an alumina support modified with 0.1 to 2 wt% boron with an active metal comprising 2 to 10 wt% cobalt and 0.001 to 0.05 wt% platinum.

[0093] Preferably, the high-speed fluidization region is in a normal state, that is, the gas flow velocity in the riser is maintained higher than that in the turbulent fluidization region and lower than that in the dilute phase air transport fluidization region, while the catalyst continuously and quantitatively flows into the riser, forming a fluidization region with a dense region at the bottom of the riser and a dilute region at the top of the riser.

[0094] The following is for reference. Figure 3 The specific implementation example of the contact pyrolysis process according to the present invention is described, but the scope of the present invention is not limited thereto.

[0095] pass Figure 3 The pipeline 11 shown facilitates the transport of the hydrocarbon mixture as raw material. To promote a smoother reaction, the mixture can be heated to 300-600°C before transport. Furthermore, depending on the composition of the transported raw material, it can be transported as a gas or as a dispersed liquid; no particular limitation is made here.

[0096] The raw material transported by pipeline 11 flows in through riser 1, which serves as the reaction zone, and at the lower part of riser 1, it mixes with the regeneration catalyst transported through regenerator stand pipe 13. Alternatively, the mixing process for transporting the raw material and the regeneration catalyst can be configured using various methods known in the art, all of which are included within the scope of this invention.

[0097] Furthermore, the catalyst used in the process is regenerated by regenerator 3, and the regenerated catalyst is transported to riser 1 through pipeline 13. At this time, the catalyst temperature at the riser inlet is 550 to 700°C, more preferably maintained at 620 to 680°C. That is, the heat provided by the regenerated catalyst 13 will raise the temperature of the transported feedstock 1 to the temperature required for the dehydrogenation reaction. When the temperature at the lower part of the riser is below 550°C, the catalyst conversion rate will decrease, while when the temperature at the lower part of the riser exceeds 700°C, the thermal decomposition of the hydrocarbon mixture as feedstock will lead to an increase in by-products, resulting in a decrease in catalyst selectivity.

[0098] Next, the feedstock and dehydrogenation catalyst mixed in the lower part of riser 1 undergo a dehydrogenation reaction in riser 1, and then flow upwards in a fluidized manner. At this time, as the dehydrogenation reaction proceeds further, as an endothermic reaction, the temperature of the mixture decreases, and the temperature of the upper part of riser 1 decreases relatively.

[0099] The reaction products and catalyst reach the upper part of the riser 1 and flow into the stripper 2, whereby the gaseous reaction products and solid catalyst are separated in a short time. To improve the efficiency of this separation process, a hydrocyclone is selectively used.

[0100] The separated gaseous reaction products are discharged through pipe 15, and the separated catalyst accumulates in stripper 2 and moves downward. At this time, stripping steam is delivered to the lower part of stripper 2 through pipe 16, and as stripping steam 16 moves upward along stripper 2, it removes the unseparated hydrocarbon reaction products contained in the catalyst and discharges them into gaseous reaction product pipe 15.

[0101] Inside the stripper 2, the catalyst reaching the lower part is regulated by the slide valve 18 and moved to the regenerator 3 through the stripper stand pipe 17. At this time, the catalyst may contain coke produced in the reaction. Oxygen-containing gases such as air flow into the regenerator 3 through the pipe 20. The coke contained in the catalyst reacts with oxygen at a high temperature above 500°C to be converted into carbon monoxide or carbon dioxide, and is discharged as waste gas through the pipe 19. As a result, the coke content in the catalyst can be significantly reduced.

[0102] In addition, the regenerated catalyst located at the bottom of the regenerator 3 is regulated by the slide valve 21 and flows back into the riser pipe through the regenerator stand pipe of the pipeline 13 to achieve recycling in the process.

[0103] The process according to the present invention can use a hydrocarbon mixture, specifically, a hydrocarbon mixture containing more than 90% by weight of propane, more preferably, a hydrocarbon mixture containing 95% by weight of propane, as a feedstock. When the concentration of propane is low, the catalyst selectivity decreases due to side reactions of other impurities, thereby leading to a reduction in production efficiency.

[0104] In this invention, the dehydrogenation catalyst can be used to dehydrogenate the transported feedstock. Generally, it is not particularly limited as long as it is known in the art to convert hydrocarbons into olefins through a dehydrogenation reaction. However, preferably, it is an active metal impregnated in a boron-modified alumina support containing 2 to 10 wt% cobalt and 0.001 to 0.05 wt% platinum.

[0105] When the boron content is less than 0.1 wt%, there is a drawback of decreased propylene selectivity; when the boron content exceeds 2 wt%, there is a drawback of decreased conversion rate. Therefore, it is not advisable to use this method.

[0106] When the cobalt content is less than 2 wt%, there is a drawback of decreased conversion rate; when the cobalt content exceeds 10 wt%, there is a drawback of decreased conversion rate and propylene selectivity, and therefore it is not advisable to use it.

[0107] When the platinum content is less than 0.001 wt%, platinum does not have the effect of co-catalyst. When the platinum content exceeds 0.05 wt%, the sharp increase in by-products (methane, ethane, etc.) leads to a decrease in propylene selectivity, making it unsuitable for use.

[0108] Preferably, the alumina support prepared at a temperature of 550 to 850°C above the dehydrogenation reaction temperature has a γ to θ phase, and within this range, the surface area is 100 to 300 m². 2 / g.

[0109] When the support is prepared at a temperature below the dehydrogenation reaction temperature, the dehydrogenation reaction will cause thermal deformation of the catalyst. When prepared at a temperature above 850°C, the surface area of ​​the catalyst is small due to the crystallization of the support, which will hinder the transfer of substances that express catalytic activity when in contact with the reactants.

[0110] Preferably, the pore size of the alumina support is 0.1–5 μm. When the pore size is smaller than this range, propane gas does not easily diffuse as an active species within the pores, which leads to reduced reactivity and decreased fluidity. On the other hand, when the pore size is too large, the catalyst strength weakens, making it unsuitable.

[0111] Further, preferably, the average size of the catalyst is 20 to 200 micrometers, more preferably, in the range of 60 to 120 micrometers. To achieve the highly efficient catalytic reaction, a moderately high-speed fluidization region is required between the turbulent fluidization region and the dilute phase fluidization region. However, in catalyst size regions smaller than 20 micrometers, dilute phase flow dominates, resulting in reduced yield due to high space velocities. Furthermore, in catalyst size regions exceeding 200 micrometers, excessively slow circulating fluidization flow reduces product production rates, thus requiring very large catalytic reaction equipment to maintain the same production efficiency, which reduces the economic benefits of investment.

[0112] As described above, a hydrocarbon reaction that converts hydrocarbon feedstock compounds into olefins occurs in the riser 1. Therefore, the main reaction conditions that affect the olefin yield include the riser temperature, the residence time of the reactants in the riser, and the volume percentage and distribution of the catalyst in the riser. These will be described in more detail below.

[0113] First, the riser temperature represents the highest temperature at the bottom; the temperature decreases as you move upwards. Furthermore, in a continuous reaction-regeneration circulating fluidized bed process, the temperature at the top of the riser determines the regeneration temperature. Therefore, in this invention, the catalyst temperature at the bottom of the riser is 550 to 700°C, and more preferably, 580 to 650°C. Temperatures below 550°C indicate a decrease in the initial activity of the catalyst, while temperatures above 700°C can trigger thermal decomposition of the feedstock within the riser, leading to feedstock loss. Moreover, a temperature of 500 to 650°C at the top of the riser, more preferably 550 to 620°C, is effective. When the top temperature is below 500°C, the regeneration process in the circulating fluidized bed reactor hinders temperature recovery; when the temperature is above 650°C, overheating occurs in the catalyst coke removal step of the regeneration process, making it impossible to control the temperature at the bottom of the riser. However, to achieve smooth flow, it is necessary to maintain a temperature at the bottom of the riser higher than the temperature at the top.

[0114] Furthermore, preferably, the pressure in the riser is maintained at 24 to 26 psig. The riser pressure represents the reaction pressure; below 24 psig, the compression energy of the product in the separation process increases, as does the investment cost of the compression equipment, which reduces overall economic efficiency. Conversely, above 26 psig, although the investment cost and compression energy of the compression equipment at the reactor downstream decrease, the riser can trigger a high-pressure reaction, reducing product yield. Therefore, a suitable pressure within the aforementioned range is required for the riser.

[0115] Furthermore, in the dehydrogenation process for producing olefins using the aforementioned catalyst, the residence time of the reactants in the riser can also be a crucial reaction condition determining the olefin yield and composition. As the dehydrogenation reaction proceeds through the riser, the amount of gas molecules remaining in the riser determines the number of gas molecules and the flow rate; therefore, a standard for determining the residence time is needed. In view of this, the present invention uses the ratio of the riser volume to the volumetric velocity of the gas flowing out from the upper part of the riser as a standard to determine the residence time of the reactants in the riser.

[0116] In the dehydrogenation process according to the present invention, the effective residence time of the hydrocarbon feedstock compound in the riser is 1 to 6 seconds, preferably 2.5 to 4.5 seconds. When the residence time is less than 1 second, sufficient contact time between the catalyst and the hydrocarbon feedstock cannot be guaranteed, which will reduce the product yield. When the residence time is longer than 6 seconds, excessive investment in reactor equipment is required to achieve the high-speed fluidization zone required by the present invention.

[0117] The fluidized bed dehydrogenation reaction of the present invention is an endothermic reaction, which utilizes high-temperature catalyst recycling to deliver the heat required for the reaction. Therefore, in the present invention, the catalyst recycling amount is more suitable for achieving the above-mentioned objective when the weight ratio of the recycled catalyst weight to the weight of the feedstock (a hydrocarbon mixture containing LPG) is 10 to 50, preferably in the range of 30 to 45.

[0118] When the weight ratio is less than 10, the space velocity between the catalyst and the hydrocarbon feedstock is too high, thus insufficient contact time for the reaction cannot be guaranteed. When the weight ratio is greater than 50, excessive investment in reactor equipment is required to achieve the high-speed fluidization zone required by this invention. Furthermore, the catalyst regeneration section will generate excessive flow, failing to guarantee sufficient regeneration time. Therefore, a suitable weight ratio of catalyst weight divided by the weight of the hydrocarbon mixture within the aforementioned range is required.

[0119] Furthermore, preferably, the gas space velocity (WHSV,h) is relative to the weight of the catalyst flowing into the riser. -1 The value is 2 to 40, more preferably 7 to 13. The space velocity is less than 2 h. -1 At that time, the reaction conversion rate increased, but the selectivity decreased. The space velocity was 40 h⁻¹. -1 At that time, selectivity increased, but conversion rate decreased.

[0120] In addition, as mentioned above, the flow fluidization region has a significant impact on the volume percentage and distribution of the catalyst in the riser. At this point, the gas velocity in the riser and the injection velocity of the catalyst flowing into the riser determine the flow fluidization region.

[0121] According to the circulating fluidized bed process of the present invention, in order to effectively prepare olefinic hydrocarbons from a mixture of hydrocarbon feedstocks, it is necessary to maintain the fluidization zone of the riser as a high-speed fluidization zone to provide sufficient catalyst volume percentage and distribution for hydrocarbon reactions to occur.

[0122] Therefore, the extent of the high-speed fluidization region needs to be defined more precisely. To this end, adjacent fluidization regions can be described by comparison: the turbulent fluidization region and the dilute-phase air transport fluidization region. First, as the gas velocity increases in the turbulent fluidization region, solid particles are obviously carried through the riser and migrate to the high-speed fluidization region. Therefore, to maintain a certain amount of catalyst in the riser at the gas velocity of the high-speed fluidization region, a continuous inflow of catalyst is required in the lower part of the riser. The catalyst volume percentage in the high-speed fluidization region changes with the height of the riser; thus, there exists a dense region in the lower part of the riser and a dilute region in the upper part.

[0123] Furthermore, in the high-speed fluidization region, when the booster gas accelerates further or the inflow of solid particles decreases, the catalyst volume inside the riser decreases and migrates to the dilute-phase air transport fluidization region. The catalyst volume percentage in the dilute-phase air transport fluidization region has a very small value and remains almost constant depending on the riser height.

[0124] Catalyst volume refers to the volume occupied by the catalyst excluding space within a given volume. In porous catalysts, catalyst volume refers to the volume including both macropores and micropores inside the catalyst.

[0125] According to Kunii and Levenspiel (1991, Fluidization Engineering), in the high-speed fluidization region, the catalyst exiting the riser rapidly undergoes mist entrainment. Therefore, to maintain normal operating conditions, continuous catalyst injection is required. Figure 3 As shown, the characteristics of the high-speed fluidization region are defined as follows.

[0126] - Starting from the lower inlet of the riser, in the shorter section, the percentage of catalyst volume in the riser volume is equivalent to 0.2 to 0.4%.

[0127] From the bottom of the riser, as the height increases, the percentage of catalyst volume remains constant at approximately 0.2% until a certain height is reached. This portion is referred to as the dense region.

[0128] - In the upper part of the riser tube above the dense phase, the volume of the catalyst gradually changes, ranging from 0.02 to 0.05 percent.

[0129] Even if the process changes, the qualitative characterization of the high-speed fluidization region will remain the same; however, the quantitative value of the catalyst volume will change. The quantitative value of the catalyst volume varies with the physical properties of the catalyst, namely, its inherent density and sphericity, and also with the type of gas, depending on its physical properties such as density and viscosity.

[0130] Therefore, the preferred method for forming the high-speed fluidization region in the circulating fluidized bed hydrocarbon process of the present invention can be utilized as follows: The gas flow velocity inside the riser is maintained above the turbulent fluidization region and below the dilute phase air transport fluidization region, while simultaneously maintaining a normal state, i.e., allowing the catalyst to continuously and quantitatively flow into the riser. At this time, the catalyst volume percentage changes with the height of the riser, thus implying that the fluidization region contains a dense region below the riser and a dilute region above the riser. More specifically, the formation method can be defined as follows.

[0131] 1) The gas velocity is maintained above the gas velocity required for the catalyst mist to flow smoothly out of the upper part of the riser, and the gas continuously flows into the catalyst from the lower part of the riser.

[0132] 2) Under the above conditions, as the gas flow rate increases, the difference in catalyst volume percentage between the bottom 1 / 4 and 3 / 4 positions within the riser decreases. Therefore, it is necessary to adjust the gas flow rate and catalyst velocity to maintain the difference in catalyst volume percentage between the two positions at 0.02 to 0.04. In particular, the catalyst of this invention exhibits highly efficient catalytic reactivity in the high-speed fluidization region. When the volume percentage difference is less than 0.02, it approaches the dilute-phase fluidization region, where the high space velocity reduces the yield. Conversely, when the volume percentage difference exceeds 0.04, it becomes turbulent fluidization or bubbling fluidization, which is unsuitable.

[0133] The process of this invention uses a hydrocarbon mixture as the feedstock. Preferably, a circulating fluidized bed process is employed to prepare olefin-based hydrocarbons from a hydrocarbon mixture containing more than 90% by weight of propane. This process adjusts the gas velocity and catalyst inflow velocity within the riser under the aforementioned conditions, thereby operating in a high-speed fluidization region, where the catalyst concentration within the riser can be maximized. Therefore, based on the above principle, olefin-based hydrocarbons are provided, and more preferably, high conversion and high selectivity of propylene are achieved.

[0134] <Catalyst>

[0135] The catalyst used in the experiment was [(4wt% Co + 0.01wt% Pt) / 0.7wt% B + alumina], wherein the alumina is commonly used Puralox or KRICT-2 alumina, preferably with a pore size of 0.1 to 5 μm, and other physical properties are shown in Table 1 below.

[0136] Table 1

[0137]

[0138] The following describes examples of olefin preparation using the catalyst.

[0139] <Example 1>

[0140] In Example 1, a circulating fluidized bed process was used to detect... Figure 3 The activity of the CPB / KRICT-2 catalyst in the olefin preparation process shown.

[0141] The circulating fluidized bed process consists of a riser, a regenerator, a stripper, and a stabilizer. The riser is 7m high and 0.94cm in diameter; the regenerator is 1.5m high and 12cm in diameter; the stripper is 2m high and 10cm in diameter; and the stabilizer is 1.7m high and 15cm in diameter.

[0142] A hydrocarbon mixture containing more than 90% propane was injected into the riser inlet to transport the raw materials, dilution gas, and catalyst and mix them. The test conditions are shown in Table 2 below.

[0143] Table 2

[0144] Catalyst temperature at riser inlet [°C] 640 650 660 Feedstock delivered [g / hr] 420 420 420 Catalyst circulation rate [g / hr] 18000 18000 18000 Catalyst / propane ratio 43 43 43 Pressure [psig] 24.9 25 25.5 Gas residence time RT [sec] 3.80 3.76 3.77 Solid residence time RT [sec] 7.93 7.93 8.63 WHSV [h-1] 10.6 10.6 9.7 Riser upper temperature [°C] 585 602 612 Riser lower temperature [°C] 610 623 635 Conversion (wt%) 25.45 32.71 33.21 Propylene selectivity (wt%) 89.54 88.20 85.23

[0145] Furthermore, in Example 1, a dehydrogenation reaction occurs in the fluidized bed while the mixture passes through the riser. The mixture passing through the riser is then stripped in a stripper at 500 to 520°C, separating into catalyst and distillate. The catalyst is recycled to the regenerator, and the distillate flows into the stabilizer. The catalyst injected into the regenerator is contacted with air and regenerated at 630 to 690°C. The regenerated catalyst then flows back into the riser. Additionally, the distillate flowing into the stabilizer is separated into gaseous and liquid components at 5 to 20°C.

[0146] The selectivity and conversion rate detected in Example 1 are approximately shown below. Figure 4 In the middle stage, as the catalyst temperature at the riser inlet changed, the conversion increased from 25% to 33%, but the selectivity decreased from 89% to 85%.

[0147] <Example 2>

[0148] Example 2 differs only in the operating conditions described in Table 3 below, and was conducted using the same method as Example 1.

[0149] Table 3

[0150]

[0151]

[0152] The selectivity and conversion rate detected in Example 2 are approximately shown below. Figure 5 In the middle stage, as the catalyst / propane ratio changed, the conversion increased from 26% to 33%, but the selectivity decreased from 92% to 88%.

[0153] <Example 3>

[0154] Example 3 used a CPB / Puralox catalyst, with the only difference being the operating conditions described in Table 4 below. The experiment was conducted using the same method as in Example 1.

[0155] Table 4

[0156] Catalyst temperature at riser inlet [°C] 640 650 Feedstock delivered [g / hr] 420 420 Catalyst circulation rate [g / hr] 18000 18000 Catalyst / propane ratio 43 43 Pressure [psig] 25.3 25.2 Gas residence time RT [sec] 3.83 3.78 Solid residence time RT [sec] 11.66 11.20 WHSV [h-1] 7.2 7.5 Riser upper temperature [°C] 589 599 Riser lower temperature [°C] 614 624 Conversion (wt%) 32.51 35.06 Propylene selectivity (wt%) 88.07 86.99

[0157] The selectivity and conversion rate detected in Example 3 are approximately shown below. Figure 6 In the middle section, as the catalyst temperature at the riser inlet changed, the conversion increased from 32.5% to 35.1%, but the selectivity decreased from 88.1% to 87%.

[0158] <Example 4>

[0159] Example 4 used a CPB / Puralox catalyst, with the only difference being the operating conditions described in Table 5 below. The experiment was conducted using the same method as in Example 1.

[0160] Table 5

[0161]

[0162]

[0163] The selectivity and conversion rate detected in Example 4 are approximately shown below. Figure 7 In the middle stage, as the catalyst / propane ratio changed, the conversion increased from 32.7% to 35.1%, while the selectivity decreased from 89.6% to 87%.

[0164] <Example 5>

[0165] Example 5 used a CPB / Puralox catalyst, with the gas residence time changed. Only the operating conditions described in Table 6 below were different, and the experiment was conducted using the same method as in Example 1.

[0166] Table 6

[0167] Gas residence time RT [sec] 3.8 5.1 Feedstock delivered [g / hr] 420 160 Catalyst circulation rate [g / hr] 18000 6500 Catalyst / propane ratio 43 41 Pressure [psig] 25.2 25.2 Catalyst temperature at riser inlet [°C] 650 650 Solid residence time RT [sec] 11.20 27.13 WHSV [h-1] 7.5 3.3 Riser upper temperature [°C] 599 593 Riser lower temperature [°C] 624 603 Conversion (wt%) 35.06 38.88 Propylene selectivity (wt%) 86.99 75.17

[0168] The selectivity and conversion rate detected in Example 5 are approximately shown below. Figure 8 In the medium-high concentration, as the catalyst / propane ratio changed, the conversion increased from 35.1% to 38.9%, but the selectivity decreased from 87% to 75.2%.

[0169] <Example 6>

[0170] In Example 1, the catalyst temperature at the riser inlet is fixed at 650°C, and the catalyst / propane ratio is fixed at 31. When implementing a continuous process, as follows... Figure 9 and Figure 10 As shown, the selectivity is 90%, and the conversion rate is 31% to 33%. When comparing the yield of the reaction products, the calculated value is about 28% to 30%, resulting in increased profit from the incremental raw materials and a significant improvement in economic benefits.

[0171] The foregoing has described the embodiments of the present invention in detail. However, the scope of the present invention is not limited by the foregoing. It will be apparent to those skilled in the art that various modifications and variations can be made without departing from the technical concept of the present invention as set forth in the claims.

[0172] [Industrial Applications]

[0173] This invention relates to a method for preparing olefins using a circulating fluidized bed process.

Claims

1. A method for preparing olefins, characterized in that: In olefin preparation methods utilizing circulating fluidized bed technology, include: (a) The step of feeding a hydrocarbon mixture containing propane and a dehydrogenation catalyst into a riser that serves as a high-speed fluidization zone to cause a dehydrogenation reaction; (b) The step of separating the catalyst from the product of the dehydrogenation reaction, i.e., the propylene mixture; (c) Stripping step, which removes unseparated hydrocarbons remaining in the catalyst separated in step (b); (d) The step of continuously regenerating the catalyst by mixing the catalyst stripped in step (c) with oxygen-containing gas; (e) Recycling the catalyst regenerated in step (d) back to step (a) for resupply in the riser; and (f) The step of cooling, compressing, and separating the reaction product separated in step (b), i.e., the propylene mixture, to prepare propylene. The boron-modified dehydrogenation catalyst is impregnated with an active metal containing cobalt and platinum on an alumina support. The catalyst comprises 0.1 to 2 wt% boron, 2 to 10 wt% cobalt, and 0.001 to 0.05 wt% platinum; In the high-speed fluidized region a) Maintain the gas velocity at a rate above the required gas velocity for the catalyst mist continuously flowing in from the bottom of the riser to flow smoothly out of the top of the riser. at the same time, b) Adjust the gas flow rate and catalyst inflow rate to keep the difference in catalyst volume percentage between the lower 1 / 4 and 3 / 4 of the riser between 0.02 and 0.

04.

2. The olefin preparation method according to claim 1, characterized in that: The hydrocarbon mixture has a propane content of 90% by weight or more.

3. The olefin preparation method according to claim 1, characterized in that: The catalyst temperature at the riser inlet is 550 to 700°C, and the temperature at the bottom of the riser is kept higher than the temperature at the top of the riser.

4. The olefin preparation method according to claim 1, characterized in that: The catalyst temperature at the riser inlet is 620 to 680°C, and the temperature at the bottom of the riser is kept higher than the temperature at the top of the riser.

5. The olefin preparation method according to claim 1, characterized in that: The pressure in the riser tube is 24 to 26 psig.

6. The olefin preparation method according to claim 1, characterized in that: The residence time of the hydrocarbon mixture is 1 to 6 seconds to facilitate the dehydrogenation reaction within the riser.

7. The olefin preparation method according to claim 1, characterized in that: The residence time of the hydrocarbon mixture is 2.5 to 4.5 seconds to facilitate the dehydrogenation reaction within the riser.

8. The olefin preparation method according to claim 1, characterized in that: In step (e), the weight ratio of the catalyst, which is again fed to the lower part of the riser, to the hydrocarbon mixture is 10 to 50.

9. The olefin preparation method according to claim 1, characterized in that: In step (e), the weight ratio of the catalyst, which is again fed to the lower part of the riser, to the hydrocarbon mixture is 30 to 45.

10. The olefin preparation method according to claim 1, characterized in that: The gas space velocity (WHSV) relative to the weight of the catalyst flowing into the riser is 2 to 40 h⁻¹. -1 .

11. The olefin preparation method according to claim 1, characterized in that: The gas space velocity (WHSV) relative to the weight of the catalyst flowing into the riser is 7 to 13 h⁻¹. -1 .

12. The olefin preparation method according to claim 1, characterized in that: The alumina support, prepared at temperatures above the dehydrogenation reaction temperature of 550 to 850°C, exhibits a γ to θ phase, with a surface area of ​​100 to 300 m² within this range. 2 / g.

13. The olefin preparation method according to claim 1, characterized in that: The alumina carrier has a pore size of 0.1 to 5. .

14. The olefin preparation method according to claim 1, characterized in that: The alumina support has a pore volume of 0.4 to 0.6 cm³. 3 / g.

15. The olefin preparation method according to claim 1, characterized in that: The catalyst has an average size of 20 to 200 micrometers.

16. The olefin preparation method according to claim 1, characterized in that: The catalyst has an average size of 60 to 120 micrometers.

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