阅读说明：本技术 利用循环流化床工序的烯烃的制备方法 (Process for the preparation of olefins using a circulating fluidized bed process ) 是由 朴德守 洪雄基 安亨赞 崔源春 朴容起 于 2018-05-01 设计创作，主要内容包括：本发明涉及一种利用高速流化床而从烃原料制备烯烃的循环流化床工序,提出一种用于更高效地增大烯烃生产所需的优选工序形态。根据本发明,利用能够在循环流化床工序中实现脱氢反应的催化剂和高速流化区域,从而提供能够发生脱氢反应的更充分的催化剂的体积比及分布,可以有效增进烯烃的生产量,特别是可以保持对丙烯的优秀的选择性。(The present invention relates to a circulating fluidized bed process for producing olefins from a hydrocarbon feedstock using a high velocity fluidized bed, and proposes a preferred process configuration for more efficiently increasing the size of the olefin production. According to the present invention, the volume ratio and distribution of the catalyst which can perform the dehydrogenation reaction are more sufficient to provide the dehydrogenation reaction, and the production of olefins can be effectively enhanced, and particularly, excellent selectivity to propylene can be maintained, by using the catalyst which can perform the dehydrogenation reaction in the circulating fluidized bed process and the high-speed fluidized zone.)

1. A method for producing olefins by a circulating fluidized bed process, comprising:

(a) a step of supplying a hydrocarbon mixture containing 90% by weight or more of LPG and the regenerated catalyst into a riser as a high-velocity fluidization region, and conducting dehydrogenation reaction in the presence of an alumina-based catalyst;

(b) a step of separating the catalyst as an effluent of the dehydrogenation reaction from the propylene mixture;

(c) a stripping step of removing hydrocarbon compounds contained in the catalyst separated in the (b);

(d) a step of continuously performing catalyst regeneration by mixing the catalyst stripped in the step (c) with a gas containing oxygen;

(e) recycling the catalyst regenerated in said (d) step to said (a) step for resupply to the riser; and

(f) a step of cooling, compressing and separating the propylene mixture as the reaction product separated in the step (b) to prepare a propylene product.

2. The process for producing olefins according to claim 1,

the high velocity fluidization region is a fluidization region in which a dense phase at the lower portion of the riser and a dilute phase at the upper portion of the riser exist in a normal state in which a fixed amount of catalyst is continuously flowed into the riser while maintaining the gas velocity in the riser higher than the turbulent fluidization region and lower than the dilute phase air-conveying fluidization region.

3. The process for producing olefins according to claim 2,

the high-velocity fluidization region (a) maintains the gas velocity at a gas velocity equal to or higher than that required for smoothly discharging the catalyst continuously flowing from the lower portion of the riser to the upper portion of the riser with the entrainment, and (b) adjusts the gas velocity and the catalyst inflow velocity so that the difference between the volume fractions of the catalyst at the two points is 0.02 or higher.

4. The process for producing olefins according to claim 3,

in the riser, the difference in the catalyst volume fractions between the lower portion 1/4 and 3/4 is maintained at 0.04 or more.

5. The process for producing olefins according to claim 1,

the hydrocarbon mixture contains 90% by weight or more of propane.

6. The process for producing olefins according to claim 1,

the alumina compound is added into Zr-Al₂O₃The carrier carries a metal component and an alkali metal at the same time.

7. The process for producing olefins according to claim 6,

the average size of the catalyst is 20-200 microns.

8. The process for producing olefins according to claim 7,

the average size of the catalyst is 60-120 microns.

9. The process for producing olefins according to claim 1,

the temperature of the lower part of the lifting pipe is 500-750 ℃, the temperature of the upper part of the lifting pipe is 450-700, and the temperature of the lower part of the lifting pipe is kept higher than that of the upper part of the lifting pipe.

10. The process for producing olefins according to claim 1,

the riser is-1 to 5kg/cm²The pressure of g.

11. The process for producing olefins according to claim 1,

the residence time of the hydrocarbon mixture for the dehydrogenation reaction in the riser is 0.1 to 500 seconds.

12. The process for producing olefins according to claim 11,

the residence time is 0.1-50 seconds.

13. The process for producing olefins according to claim 13,

the residence time is 0.5-5 seconds.

14. The process for producing olefins according to claim 1,

the weight ratio of the weight of the catalyst re-supplied to the lower part of the riser in the step (e) divided by the weight of the hydrocarbon mixture is 10-100.

15. The process for producing olefins according to claim 14,

the weight ratio is 20-60.

Technical Field

The present invention relates to a process for the preparation of olefins using a circulating fluidized bed process.

Background

Olefins such as ethylene, propylene are widely used in the petrochemical industry. Generally, such olefins are obtained in the pyrolysis process of naphtha. However, since the competitiveness of processes using lower hydrocarbons as raw materials has been gradually increased due to the shale gas revolution and the like, there is a demand for an On-purpose (On-pure) production process of olefins through a catalytic dehydrogenation process.

The catalytic contact dehydrogenation step required for olefin production uses various lower hydrocarbon compounds as a raw material, and exhibits excellent olefin production yield. However, in the fixed-bed commercial dehydrogenation process, although the olefin yield is high at the initial stage of the reaction in which the hydrocarbon is in contact with the catalyst, the catalyst is deactivated with time and excessive coking occurs, so that there is a problem that the conversion rate of the hydrocarbon and the yield of the olefin as a whole are reduced and the regeneration process consumes a large amount of energy. To solve this problem, a circulating fluidized bed process has been proposed which limits the contact time of the hydrocarbon with the catalyst to a short time.

However, even in the step of limiting the contact time between the hydrocarbon and the catalyst, the hydrocarbon reacts with the catalyst to rapidly form a by-product other than an olefin at the initial stage of the reaction, and thus the conversion rate of the reaction raw material is high, but the selectivity is low.

In a process for producing olefins from a hydrocarbon feedstock mixture through a circulating fluidized bed process, in order to selectively produce olefins such as ethylene and propylene at a high conversion rate and with a high degree of selectivity, the setting of the operating conditions of a Riser (Riser) that mainly performs a dehydrogenation reaction may be considered as an important factor. In particular, the fluidization and reaction phenomena in the riser can be more easily understood by the following theoretical examination, which will be described in more detail below.

As shown in fig. 1, if gas flows into a container filled with a solid catalyst from below, particles are fluidized, and if the Minimum Fluidization Velocity (Minimum Fluidization Velocity) or more is reached, a fluidized flow region (flowregion) is generally divided into 5 regions.

Specifically, they are called a Minimum Fluidization region (Minimum Fluidization region), a Bubbling Fluidization region (Bubbling Fluidization region), a slug Fluidization region (slug Fluidization region), a Turbulent Fluidization region (Turbulent Fluidization region), and a dilute phase air transport Fluidization region (dilute Fluidization with Pneumatic transport region), and there is a difference in particle motion characteristics among the regions.

Therefore, in the process using the fluidized bed reactor, a fluidized flow region conforming to the characteristics of each process is formed and operated.

Figure 2 shows the change in volume fraction of catalyst in the reactor for different riser heights, i.e., different fluidized flow zones, confirming that the volume fraction of catalyst in the reactor changes as the fluidized flow zone changes. However, as shown in the fluidized bed contact dehydrogenation reaction step, in the reaction involving the catalyst, the catalyst volume fraction has a significant influence on the performance of the step, and as a result, the step operation conditions for determining the fluidized flow region of the catalyst volume fraction in the left and right reactors play a very important role in the reaction result.

In order to determine the fluidized flow zone of the riser of such a circulating fluidized process, factors such as the size of the catalyst, the circulation velocity of the catalyst, the ratio of the supplied feedstock to the catalyst, the strength of the catalyst, etc. should be considered.

In addition, factors directly affecting the dehydrogenation reaction, such as reaction temperature, endothermic heat of reaction, reaction time, catalyst deactivation by Coke (Coke) formation, and the like, should be considered.

Therefore, in the present invention, during research into a method for producing olefins using a circulating fluidized bed process having superior economy and productivity compared to the conventional production process, a more efficient method for producing olefins has been developed by applying a catalyst having superior conversion rate, selectivity, and stability to the circulating fluidized bed process, and the present invention has been completed.

Disclosure of Invention

The present invention aims to provide a method for producing olefins in a fluidized bed, which is more economical and has better productivity than conventional processes.

The method for preparing olefins using a circulating fluidized bed process of the present invention comprises:

(a) a step of supplying a hydrocarbon mixture containing 90% by weight or more of LPG and a regenerated catalyst into a riser as a high-velocity fluidization region to perform dehydrogenation reaction in the presence of an alumina-based catalyst;

(b) a step of separating the catalyst as an effluent of the dehydrogenation reaction from the propylene mixture;

(c) a stripping step of removing hydrocarbon compounds contained in the catalyst separated in the (b);

(d) a step of continuously performing catalyst regeneration by mixing the catalyst stripped in the step (c) with a gas containing oxygen;

(e) recycling the catalyst regenerated in said (d) step to said (a) step for resupply to the riser; and

(f) a step of cooling, compressing and separating the propylene mixture as the reaction product separated in the step (b) to prepare a propylene product.

The high velocity fluidization region is a fluidization region in which a dense phase (dense region) in the lower portion of the riser and a dilute phase (dilute region) in the upper portion of the riser exist in a normal state in which a fixed amount of catalyst is continuously flowed into the riser while maintaining the gas velocity in the riser above the turbulent fluidization region and below the dilute phase air-conveying fluidization region.

Preferably, the high-velocity fluidization region (a) maintains the gas velocity at a gas velocity or higher required for smoothly discharging the catalyst continuously flowing from the lower portion of the riser to the upper portion of the riser with the mist, and (b) adjusts the gas velocity and the catalyst inflow velocity so that the difference in the volume fractions of the catalyst at the two points becomes 0.02 or higher.

More preferably, in the riser, the difference in the catalyst volume fractions between the point of the lower part 1/4 and the point of 3/4 is kept to 0.04 or more.

The hydrocarbon mixture containing LPG in an amount of 90 wt% or more as the feedstock of the present invention preferably contains propane in an amount of 90 wt% or more, and more preferably contains propane in an amount of 95 wt%.

The catalyst used in the olefin production method of the present invention is preferably Zr-Al as an alumina-based compound capable of effecting dehydrogenation reaction₂O₃The carrier carries a metal component and an alkali metal at the same time.

The average size of the catalyst is 20-200 microns, and preferably 60-120 microns.

In the olefin production method of the present invention, it is preferable that the temperature of the lower portion of the riser is 500 to 750, the temperature of the upper portion of the riser is 450 to 700, and the temperature of the lower portion of the riser is maintained higher than the temperature of the upper portion of the riser.

Preferably, the riser is maintained at-1 to 5kg/cm²The pressure of g.

The residence time of the hydrocarbon mixture for the dehydrogenation reaction in the riser is 0.1 to 500 seconds, preferably 0.1 to 50 seconds, and more preferably 0.5 to 5 seconds.

In the process for the preparation of olefins according to the present invention, the weight ratio of the weight of the catalyst re-supplied to the lower portion of the riser in the step (e) divided by the weight of the hydrocarbon mixture is 10 to 100, preferably 20 to 60.

The present invention relates to a circulating fluidized bed process for producing olefins from a hydrocarbon feedstock using a high velocity fluidized bed, which can more efficiently increase the production of olefins.

That is, the process of the circulating fluidized bed of the present invention has improved selectivity compared with the conventional commercial process, so that the incremental yield per unit raw material is increased, a smaller air flow rate and a smaller air compressor are required by means of a direct heat supply manner during the regeneration process, the fuel consumption is reduced by 10 to 15% compared with the commercial process, the compressor energy required for product separation and catalyst regeneration is reduced by 15 to 20%, and the investment cost is reduced as a whole. Further, since the operation is performed under positive pressure, the facility cost of the rear end product separation step is reduced as compared with the commercial step in which the operation is performed under vacuum.

Drawings

Fig. 1 is a graph illustrating changes in the internal characteristics of a fluidized bed with respect to a fluidized flow region under a change in the velocity of a normal gas.

FIG. 2 is a graph illustrating the volume fraction of catalyst inside a fluidized bed at different riser heights.

FIG. 3 is a view schematically illustrating a process of a circulating fluidized bed used in the present invention.

Fig. 4 illustrates a schematic view of a fluidized bed cold model (ColdModel) required for an experiment on a fluidized flow region at normal temperature.

FIG. 5 is a graph illustrating the volume fraction of catalyst inside a fluidized bed in a dilute phase air transport fluidization region as a result of a cold state model experiment of comparative example 1.

Fig. 6 is a graph illustrating a volume fraction of the catalyst inside the fluidized bed in the high-velocity fluidization region as a result of the cold state model experiment of example 1.

Reference numerals

1: lifting pipe (Riser)

2: gas lift tower (Strepper)

3: regenerator (Regenerator)

11: hydrocarbon feedstock supply line

13: regenerator vertical Pipe (Regenerator Stand Pipe)

15: gaseous reaction product

16: stripping Steam (striping Steam) supply line

17: gas lift tower riser (Stripper Standard Pipe)

18: stripper Slide Valve (Strepper Slide Valve)

19: exhaust Gas (fluent Gas)

20: oxygen-containing gas such as air

21: regenerator sliding Valve (Regenerator Slide Valve)

51: cold Model (Cold Model) riser

52: cold Model Cyclone separator (Cyclone)

53: cold Model (Cold Model) riser (Stand Pipe)

54: cold Model (Cold Model) feed back control valve (Loop Seal)

60: cold Model (Cold Model) main gas supply line

61: cold Model (Cold Model) gas supply line for catalyst circulation regulation

62: cold Model riser (Stand Pipe) gas supply line

63: cold Model (Cold Model) exhaust gas

Detailed Description

The process of the present invention, as mentioned above, comprises:

(a) a step of supplying a hydrocarbon mixture containing 90% by weight or more of LPG and a regenerated catalyst into a riser as a high-velocity fluidization region, and performing a dehydrogenation reaction in the presence of the catalyst;

(b) a step of separating the catalyst as an effluent of the dehydrogenation reaction from the propylene mixture;

(c) a stripping step of removing hydrocarbon compounds contained in the catalyst separated in the (b);

(d) a step of continuously performing catalyst regeneration by mixing the catalyst stripped in the step (c) with a gas containing oxygen;

(e) recycling the catalyst regenerated in said (d) step to said (a) step for resupply to the riser; and

(f) a step of cooling, compressing and separating the propylene mixture as the reaction product separated in the step (b) to prepare a propylene product.

Preferred embodiments of the present invention will be described below with reference to the accompanying drawings. However, the embodiment of the present invention may be modified into various different forms, and the scope of the present invention is not limited to the embodiments described below.

In describing the present embodiment, the same names and reference numerals are used for the same components, and therefore, overlapping additional description is omitted below. In the drawings referred to below, the scale is not applicable.

An embodiment of the contact decomposition step of the present invention will be described in more detail with reference to fig. 3, but the scope of the present invention is not limited thereto.

The above-described raw material may be supplied through the line 11 of FIG. 3, and in this case, the raw material may be heated to a temperature of 30 to 600 ℃ for smooth reaction. The supply material may be supplied in a gas state or a dispersed liquid state depending on the components of the supply material, but is not particularly limited thereto. The feed material of the line 11 flows into the riser 1 as a reaction zone, and is mixed with the regenerated catalyst supplied through a Regenerator Stand Pipe (Regenerator Stand Pipe) of a line 13 at the lower part of the riser 1. The mixing step of supplying the raw material and regenerating the catalyst may be performed by various methods known in the art, and such a configuration is included in the field of the present invention.

On the other hand, the catalyst used in the step is regenerated in the regenerator 3, and the regenerated catalyst is supplied to the riser 1 through the line 13, and the temperature of the lower portion of the riser at this time is preferably maintained at 500 to 750 ℃. That is, the feed 1 is raised to a temperature required for the dehydrogenation reaction by the heat supplied from the regenerated catalyst 13. If the temperature of the lower portion of the riser is 500 or less, the conversion rate of the catalyst decreases, and if the temperature is 750 or more, the selectivity of the catalyst decreases due to an increase in by-products caused by thermal decomposition of LPG as a raw material.

Subsequently, the feedstock and the catalyst mixed in the lower portion of the riser 1 are fluidized in the riser 1 into the upper portion in association with the dehydrogenation reaction. At this time, as the dehydrogenation reaction, which is an endothermic reaction, proceeds, the temperature of the mixture decreases, and the temperature of the upper portion of the riser 1 relatively decreases. The reaction product and the catalyst that have reached the upper part of the riser 1 flow into the stripper 2, and therefore the reaction product as a gas is separated from the catalyst as a solid in a short time. Optionally, a cyclone separator may also be used in order to increase the efficiency of the separation process. The separated reaction product in a gaseous state is discharged through the line 15, and the separated catalyst is deposited on the stripper 2 and moves in a downward direction. At this time, Stripping Steam (striping Steam) is supplied to the lower portion of the stripper 2 through a line 16, and the Stripping Steam 16 removes the unseparated hydrocarbon reaction product containing the catalyst while moving along the stripper 2 in the upper portion, and they are discharged to a gas reaction product line 15.

Inside said Stripper 2, the catalyst reaching the lower part is moved to the regenerator 3 through a Stripper riser (Stripper Stand Pipe) of line 17 by means of the adjustment of the slide valve 18. At this time, the catalyst may contain Coke (Coke) generated during the reaction. In the regenerator 3, a gas such as air containing oxygen is introduced through the line 20, coke contained in the catalyst is converted into carbon monoxide or carbon dioxide by reacting with oxygen at a high temperature of 500 or more, and is discharged as an exhaust gas through the line 19, and as a result, the content of coke contained in the catalyst can be significantly reduced.

On the other hand, the regenerated catalyst present in the lower part of the Regenerator 3 can be recirculated into the riser via the Regenerator Stand Pipe (line 13) by means of the regulation of the slide valve 21 and recycled in the process.

In the process of the present invention, a hydrocarbon compound, specifically, a hydrocarbon mixture containing 90 wt% or more of LPG may be used as a raw material. The hydrocarbon mixture containing LPG in an amount of 90 wt% or more as the feedstock of the present invention preferably contains propane in an amount of 90 wt% or more, and more preferably contains propane in an amount of 95 wt%. If the concentration of propane is low, the catalyst selectivity decreases due to side reactions with other impurities, and thus productivity decreases.

In the present invention, the catalyst that can be used for dehydrogenating the feed stock is not particularly limited as long as it can convert a hydrocarbon compound into an olefin by a dehydrogenation reaction, and is preferably a catalyst containing alumina. The catalyst may additionally contain a metal component as an auxiliary support component in the alumina support, and a transition metal and an alkali metal as an active component.

Preferably, the catalyst comprises one selected from zirconium, zinc and platinum as an auxiliary carrier component and one selected from oxides of chromium, vanadium, manganese, iron, cobalt, molybdenum, copper, zinc, cerium and nickel as a transition metal. More preferably, zirconium is included as an auxiliary carrier component, chromium is included as a transition metal, and potassium is included as an alkali metal.

Further, the average size of the catalyst is preferably 20 to 200 micrometers, and more preferably 60 to 120 micrometers. For the high efficiency catalytic reaction, a suitably high velocity fluidized zone flow is required between the turbulent fluidized zone and the dilute phase fluidized zone, and in the catalyst size region of 20 μm or less, the dilute phase flow dominates and the yield is lowered by the high space velocity. In addition, in the catalyst size region of 200 μm or more, since the circulating fluidized flow is too slow, the production rate of the product is low, and a very large catalytic reaction apparatus is required to maintain the same productivity, and the investment economy is low.

As already mentioned, the dehydrogenation reaction of the hydrocarbon feedstock compound to olefins takes place in the riser 1, and therefore, as the main reaction conditions which influence the yield of olefins, there are the riser temperature, the residence time of the reactants in the riser, the volume fraction and distribution of the catalyst in the riser, etc., as will be explained in more detail below.

First, the riser temperature exhibits the highest temperature in the lower portion, and the temperature decreases the more it moves to the upper portion. Therefore, in the present invention, it is most effective to maintain the temperature of the lower portion of the riser tube at 500 to 750 deg.f, and to maintain the temperature of the upper portion of the riser tube at 450 to 700 deg.f. However, for a smooth flow, the temperature in the lower part of the riser must be kept higher than in the upper part of the riser.

On the other hand, the riser is preferably maintained at-1 to 5kg/cm²The pressure of g. The pressure of the riser represents the reaction pressure, and is lower than-1 kg/cm²G, the compression energy required to separate the product in the product increases, and the investment cost for compression equipment also increases, thereby reducing the overall economy. In addition, when the pressure of the riser is higher than 5kg/cm²G, the capital cost and pressure energy of the compression equipment at the rear end of the reactor are reduced, but a high-pressure reaction is induced in the riser, reducing the product yield. Therefore, a suitable riser pressure in the range described is required.

In addition, in the dehydrogenation step for producing an olefin using the catalyst, the residence time of the reactant in the riser also becomes an important reaction condition for determining the yield and composition of the olefin. As the dehydrogenation reaction proceeds through the riser, the number of molecules of the gas and the flow rate are determined depending on how long the gas stays in the riser, and therefore, a criterion for determining the residence time is required. Therefore, in the present invention, the value of the volume of the riser divided by the volumetric velocity of the gas flowing out from the upper portion of the riser is used as a reference for the residence time of the reactants in the riser.

In the dehydrogenation step of the present invention, the effective residence time of the hydrocarbon feedstock compound in the riser is carried out in the range of 0.01 to 500 seconds, preferably 0.1 to 50 seconds, more preferably 0.5 to 5 seconds. When the residence time is less than 0.1 second, sufficient contact time of the catalyst with the hydrocarbon feedstock cannot be secured to lower the product yield, and when it is more than 500 seconds, the capital cost of the reactor equipment required for the high-velocity fluidization region required for embodying the present invention is excessive.

In the present invention, the fluidized bed dehydrogenation reaction is used as an endothermic reaction, and the heat required for the reaction is supplied by the recirculation of the catalyst at a high temperature. Therefore, in the present invention, the amount of the catalyst recycle more suitable for such purpose is most effective in the range of 10 to 100, preferably 20 to 60, in terms of the weight ratio of the weight of the catalyst recycle divided by the weight of the feed (LPG-containing hydrocarbon mixture).

When the weight ratio is less than 10, the space velocity with the hydrocarbon as the feedstock is too high to ensure a sufficient reaction contact time, and when it is more than 100, the capital requirement for the reactor equipment required for the high-velocity fluidization region required for embodying the present invention is too high. In addition, since an excessive flow rate is induced in the catalyst regeneration section and a sufficient regeneration time cannot be secured, a weight ratio obtained by dividing the appropriate catalyst weight in the above-described range by the hydrocarbon mixture weight is required.

On the other hand, as described above, the volume fraction and distribution of the catalyst in the riser are greatly affected by the fluidized flow region (fluidized region), which is determined by the gas velocity in the riser and the injection velocity of the catalyst into the riser.

In order to efficiently produce olefinic hydrocarbon compounds from a hydrocarbon feedstock mixture, it is important to maintain the fluidization region of the riser in a high velocity fluidization region to provide a sufficient volume fraction and distribution of catalyst that can be subjected to dehydrogenation reactions.

Therefore, the range of the high-velocity fluidization region needs to be more clearly defined, and for this reason, a description can be made in comparison with a turbulent fluidization region as an adjacent liquefaction region and a dilute-phase air-conveying fluidization region. First, in the turbulent fluidization region, as the gas velocity increases, it leaves the riser with a significant droplet of solid particles, transferring to the high velocity fluidization region. Therefore, at the gas velocity in the high velocity fluidization region, the catalyst is continuously introduced from the lower portion of the riser in order to maintain a predetermined amount of catalyst in the riser. In the high velocity fluidization region, the catalyst volume fraction varies with the height of the riser, with a dense phase (dense region) in the lower portion of the riser and a dilute phase (dilute region) in the upper portion of the riser.

Further, in the high-velocity fluidization region, if the velocity of the ascending gas is further increased or the inflow of the solid particles itself is decreased, the volume of the catalyst in the riser is decreased, and the catalyst is transferred to the dilute-phase air-conveying fluidization region. In the dilute phase air transport fluidization zone, the catalyst volume fraction has a very low value, with almost established values as a function of the height of the riser.

The catalyst volume is a volume occupied by the catalyst except for a space in a predetermined volume, and in the case of a porous catalyst, the volume includes macropores (macro pores) and micropores (micro pores) inside the catalyst.

Kunii and Levenspiel (1991, fluidization engineering) describe that in the high-velocity fluidization regime, the catalyst spray leaving the riser proceeds rapidly, so that in order to maintain normal operating conditions, continuous injection of catalyst is required, as shown in FIG. 3, the characteristics of the high-velocity fluidization regime being defined as follows.

The volume of catalyst corresponds to 0.2 to 0.4 fraction of the riser volume in the shorter part from the inlet in the lower part of the riser.

In the lower part of the riser, the volume of catalyst has a given value of about 0.2 fraction as the height increases up to a given height. This portion is referred to as dense phase (dense region).

-in the upper part of the riser above the dense phase, the volume of catalyst varies, being present in a fraction of 0.02 to 0.05.

The qualitative characteristics of the high velocity fluidization region are the same with process variations, but the quantitative value of the catalyst volume changes. The quantitative value of the catalyst volume varies depending on the physical properties of the catalyst, that is, the density and sphericity inherent to the catalyst, and also varies depending on physical substances such as the gas density and viscosity determined by the change in the gas type.

Thus, the preferred high velocity fluidization region that can be utilized in the hydrocarbon compound circulating fluidized bed dehydrogenation process of the present invention is formed by maintaining a fixed amount of catalyst continuously flowing into the normal state of the riser while maintaining the gas velocity in the riser above the turbulent fluidization region and below the dilute phase air transport fluidization region, where it can be stated that the catalyst volume fraction varies with the height of the riser, meaning the fluidization region where the dense phase (dense region) in the lower portion of the riser and the dilute phase (dilute region) in the upper portion of the riser exist. More specifically, it can be formed or defined as follows.

1) The gas velocity is maintained at a gas velocity or higher required for the catalyst to smoothly leave the upper part of the riser with the mist, and the catalyst needs to be continuously allowed to flow into the lower part of the riser.

2) Under these conditions, as the gas flow rate increases, the difference in the volume fractions of the catalyst in the riser between the point 1/4 and the point 3/4 from the lower part decreases, and therefore, the gas flow rate and the catalyst inflow rate need to be adjusted so that the difference in the volume fractions of the catalyst between the two points is maintained at 0.02 or more, preferably 0.04 or more. The catalyst of the present invention exhibits high catalytic reactivity particularly in a high-velocity fluidized zone, and if the difference in the volume fraction is 0.04 or less, the yield is lowered due to a high space velocity close to a dilute-phase fluidized zone.

According to the present invention, in the process of producing olefinic hydrocarbons from a hydrocarbon mixture as a feed feedstock, preferably, from a hydrocarbon mixture containing 90% by weight or more of LPG using a circulating fluidized bed process, the gas velocity in the riser and the catalyst inflow velocity into the riser are adjusted under the conditions as described above so as to operate in a high-velocity fluidization region, at which time the maximum catalyst concentration can be provided in the riser. Thus, by this principle, a high conversion and a high selectivity for olefinic hydrocarbons, preferably for propylene, can be provided.

Comparative example 1: dilute phase fluidized zone

A. Production of Cold Model

In order to grasp the fluidized flow region caused by changes in the gas velocity and the catalyst injection velocity at normal temperature, a fluidized bed cold model was prepared as shown in fig. 4. The solid particles passing through the solids circulation rate regulating return control valve 54 flow into the cold model riser 51 and travel along the riser to the upper section by means of the main gas supply of line 60. The gas and solids are separated in a cold model Cyclone (Cyclone), the gas is discharged as off-gas in line 63 and the solids move down the cold model riser (Stand Pipe) in line 53. At this time, the circulation of solids is smoothly performed by the gas supplied through the line 62. The cold model return control valve 54 regulates the catalyst circulation amount by means of a cold model catalyst circulation adjustment gas supplied through a line 61.

In the comparative example 1, the cold state model riser was made with a height of 2.5m and a diameter of 0.9cm, and the riser and the return control valve were made so that the catalyst circulation was smoothly realized.

B. Catalyst and process for preparing same

The catalyst used in the experiment was [ (5% Cr + 0.5% K)/5% Cr/Zr-Al₂O₃]The volume average diameter is 78 micrometers, and the particle size distribution is composed of 10% below 60 micrometers, 90% between 60 and 100 micrometers, and 10% above 100 micrometers.

C. Fluidized flow zone experiment

The experiment was carried out at normal temperature and pressure, the catalyst circulation rate being 20.2kg/hr injected through the inlet at the lower part of the riser, corresponding to 88.1kg/m2. sec inside the riser. Under the above conditions, the pressure drop was measured from the height of the riser, and the volume fraction (Solid fraction) of the catalyst was obtained (see fig. 5). In FIG. 5, the volume fractions of catalyst at 1/4 and 3/4 in the lower riser are 0.049 and 0.040 respectively, the difference between the two points being 0.009, which by definition of the invention is indicative of the fluidization regime of the dilute phase air transport fluidization region.