CN112708450B

CN112708450B - Method for producing propylene by catalytic cracking of hydrocarbons

Info

Publication number: CN112708450B
Application number: CN201911024569.0A
Authority: CN
Inventors: 杨超; 龚剑洪; 谢朝钢; 朱根权; 沙有鑫; 成晓洁
Original assignee: Sinopec Research Institute of Petroleum Processing; China Petroleum and Chemical Corp
Current assignee: Sinopec Research Institute of Petroleum Processing; China Petroleum and Chemical Corp
Priority date: 2019-10-25
Filing date: 2019-10-25
Publication date: 2022-11-15
Anticipated expiration: 2039-10-25
Also published as: CN112708450A

Abstract

The invention relates to the field of petroleum refining and petrochemical processing, and discloses a method for producing propylene by catalytic cracking of hydrocarbons, which comprises the following steps: (1) Carrying out first catalytic cracking reaction on heavy raw oil and a first regenerated catalyst to obtain a first oil mixture; (2) Carrying out second catalytic cracking reaction on the light raw oil and a second regenerated catalyst to obtain a second oil mixture; (3) Carrying out a third catalytic cracking reaction on the first oil mixture and the second oil mixture to obtain a third oil mixture, and recycling after separation, precondition and regeneration; the mass ratio of the heavy raw oil to the light raw oil is 0.5-5:1; the method further comprises the following steps: introducing a first cooling medium to an outlet of the first riser reactor; and/or introducing a second cooling medium into the stripper, and recycling part of the stripped spent catalyst to the first riser reactor. The method provided by the invention can improve the propylene selectivity and simultaneously reduce the yield of dry gas and coke.

Description

Method for producing propylene by catalytic cracking of hydrocarbons

Technical Field

The invention relates to the field of petroleum refining and petrochemical processing, in particular to a method for producing propylene by catalytic cracking of hydrocarbons.

Background

In 2015, the crude oil processing capacity of a domestic catalytic cracking unit is about 7.5 hundred million tons per year, the operating rate of the catalytic cracking unit is about 70 percent, and the consumption demand of diesel oil is turned. With the further expansion of oil refining capacity and the popularization of new energy automobiles, the demand of automobile fuel continues to decline, and by 2020, the operation rate of a catalytic cracking unit is reduced to 68%, and the demand of gasoline consumption will have inflection points. A large number of oil refining devices will be idle, and transformation to chemical engineering will be the exit of oil refining enterprises.

The catalytic cracking technology of heavy oil developed by the Chinese petrochemical science and chemistry institute is the first catalytic cracking technology which is verified by industry and takes the production of propylene as a main target product. The heavy oil is used as a raw material, and the yield of the propylene is up to 15-25%. The DCC technology for producing low carbon olefin by heavy oil catalytic cracking is described in CN1234426A, CN1388215A, CN1566272A and the like.

On the basis of the DCC technology, the reaction environment of a bed reactor is flexibly controlled by adopting a second lifting pipe, the primary cracking reaction of heavy oil and the secondary cracking reaction environment of gasoline are optimized, the yield of dry gas and coke is further reduced, and the DCC-plus technology is developed. For a detailed description of the DCC-plus technique see CN1333046C, CN1978411A, CN1986505A, CN100465250C, CN101029248A, etc.

In 2017, the related policies of the ethanol gasoline for the automobiles in China are continuously issued. According to the technical requirements of the ethanol gasoline (E10) (VIB) for vehicles, the content (mass fraction) of other organic oxygen-containing compounds except ethanol in the ethanol gasoline is not more than 0.5 percent, which means that etherification products represented by MTBE cannot be added into the gasoline as blending components, and the MTBE is stopped to be used in 2020. At present, the domestic MTBE capacity reaches 1750 ten thousand tons/year, and as a main way for utilizing the C-IV isobutene, after the production of an MTBE device is stopped, a processing and utilizing scheme for C-IV isobutene is the first problem facing oil refining enterprises.

The DCC-plus process recycles the C4/cracked naphtha to the second reaction zone for continuous reaction, and further increases the yield of propylene through oligomerization and re-cracking. After the MTBE unit is shut down, a large amount of the carbon four resource is available as a raw material for the DCC-plus unit, but simply increasing the amount of light hydrocarbon processed results in a decrease in the propylene yield and an increase in the coke yield.

Disclosure of Invention

The invention aims to provide a method for producing propylene by catalytic cracking of hydrocarbons. The method can improve the yield of propylene and reduce the yield of coke while processing a large proportion of light raw oil.

In order to achieve the above object, the present invention provides a method for producing propylene by catalytic cracking of hydrocarbons, the method comprising:

(1) Contacting heavy raw oil with a first regenerated catalyst in a first riser reactor to perform a first catalytic cracking reaction to obtain a first oil mixture;

(2) Contacting the light raw oil with a second regenerated catalyst in a second riser reactor to perform a second catalytic cracking reaction to obtain a second oil mixture;

(3) Introducing the first oil mixture and the second oil mixture into a fluidized bed reactor for a third catalytic cracking reaction to obtain a third oil mixture; separating the third oil agent mixture, introducing the separated carbon-deposited catalyst into a stripper for stripping, regenerating the stripped catalyst to be regenerated, and introducing the obtained regenerated catalyst into the first riser reactor and the second riser reactor for recycling;

the mass ratio of the heavy raw oil to the light raw oil is 0.5-5:1;

the method further comprises the following steps: introducing a first cooling medium to an outlet of the first riser reactor; and/or introducing a second cooling medium into the stripper, and recycling part of the stripped spent catalyst back to the first riser reactor.

Preferably, the mass ratio of the heavy raw oil to the light raw oil is 0.5-2:1, preferably the method further comprises: dividing the light raw oil into a light raw oil A and a light raw oil B, and introducing the light raw oil B into a first riser reactor; wherein the mass ratio of the total amount of the light raw oil B and the heavy raw oil to the light raw oil A is 4-5:1.

the inventors of the present invention found in the course of their studies that it is most advantageous to produce propylene in a large amount when the circulating amount of the first regenerated catalyst is 2 to 4 times as large as the circulating amount of the second regenerated catalyst. When the second riser reactor processes light raw oil in a large proportion, the circulation quantity of the catalyst of the second riser reactor is obviously increased, the severity of the first catalytic cracking reaction and the bed reaction environment are seriously influenced, the yield of dry gas coke is obviously increased, and the propylene selectivity is reduced. The invention provides a method for producing propylene by catalytic cracking of hydrocarbons, which comprises the steps of introducing a first cooling medium to an outlet of a first riser reactor; and/or introducing a second cooling medium into the stripper, and recycling part of the stripped spent catalyst back to the first riser reactor. Introducing a first cooling medium into the outlet of the first riser reactor, so that the outlet temperature of the first riser reactor can be reduced, and the agent-oil ratio of the first riser is not limited by the treatment capacity of the second riser; and introducing a second cooling medium into the stripper to reduce the temperature of the spent catalyst, returning the cooled spent catalyst to the first riser reactor to be mixed with the regenerated catalyst, reducing the contact temperature of the oil agent, greatly improving the oil-to-catalyst ratio and reducing the thermal cracking reaction degree. Compared with the prior art, the method provided by the invention can improve the propylene selectivity and simultaneously reduce the yield of dry gas and coke.

Drawings

FIG. 1 is a schematic flow diagram of the process for the production of propylene by catalytic cracking of hydrocarbons as described in example 1;

FIG. 2 is a schematic flow diagram of the process for the catalytic cracking of hydrocarbons to produce propylene as described in example 2.

Description of the reference numerals

Detailed Description

The endpoints of the ranges and any values disclosed herein are not limited to the precise range or value, and such ranges or values should be understood to encompass values close to those ranges or values. For ranges of values, between the endpoints of each of the ranges and the individual points, and between the individual points may be combined with each other to give one or more new ranges of values, and these ranges of values should be considered as specifically disclosed herein.

In the present invention, the terms "first" and "second" do not have any limiting effect, but are used only for clarity and convenience of description to distinguish the operations performed at different stages and the added materials.

The invention provides a method for producing propylene by catalytic cracking of hydrocarbons, which comprises the following steps:

(3) Introducing the first oil mixture and the second oil mixture into a fluidized bed reactor for a third catalytic cracking reaction to obtain a third oil mixture; separating the third oil agent mixture, introducing the carbon-deposited catalyst obtained by separation into a stripper for steam stripping, regenerating the catalyst to be regenerated after steam stripping, and introducing the obtained regenerated catalyst into the first riser reactor and the second riser reactor for recycling;

the mass ratio of the heavy raw oil to the light raw oil is 0.5-5:1;

The method provided by the invention can introduce a first cooling medium to the outlet of the first riser reactor, can also introduce a second cooling medium to the stripper, and can recycle part of the stripped spent catalyst to the first riser reactor, or can introduce the first cooling medium to the outlet of the first riser reactor and also introduce the second cooling medium to the stripper (and recycle part of the stripped spent catalyst to the first riser reactor).

The present invention is wide in the selection range of the types of the first cooling medium and the second cooling medium, and preferably, the first cooling medium and the second cooling medium are each independently water and/or gasoline, and more preferably water.

Preferably, the first cooling medium and the second cooling medium each independently have a temperature of 20 to 99 ℃.

According to the present invention, the amounts of the first cooling medium and the second cooling medium introduced may be appropriately selected according to the amount of heavy feedstock oil to be processed, and preferably, the amounts of the first cooling medium and the second cooling medium introduced are each independently 50 to 200 kg/ton of heavy feedstock oil, and more preferably 60 to 120 kg/ton of heavy feedstock oil.

According to a preferred embodiment of the present invention, the mass ratio of the first regenerated catalyst to the second regenerated catalyst is 2 to 4:1. with this preferred embodiment, it is more advantageous to produce more propylene.

According to the present invention, preferably, the operating conditions of said first riser reactor comprise: the agent-oil ratio is 8-20, the oil-gas retention time is 0.5-5s, and the outlet temperature is 550-580 ℃; preferably, the agent-oil ratio is 10-20, the oil-gas retention time is 0.8-2s, and the outlet temperature is 550-570 ℃.

According to the present invention, preferably, the operating conditions of said second riser reactor comprise: the outlet temperature is 610-650 ℃, and the oil gas retention time is 0.1-4s; further preferably, the outlet temperature is 620-630 ℃ and the oil gas residence time is 0.5-3s.

The present invention has a wide range of operating conditions for the fluidized bed reactor, preferably, the operating conditions for the fluidized bed reactor include: the reaction temperature is 560-580 ℃, and the weight hourly space velocity is 2-8h ^-1 Go forward toPreferably, the reaction temperature is 560-570 ℃, and the weight hourly space velocity is 4-8h ^-1 。

The first riser reactor, the second riser reactor, and the fluidized bed reactor may each independently be reactors conventionally used in the art. The first riser reactor and the second riser reactor can be of equal diameter, variable diameter, straight pipe or curved pipe. Preferably, the first riser reactor and the second riser reactor are straight pipes with equal diameters or variable diameters, or curved pipes with equal diameters or variable diameters. Further preferably, the first riser reactor and the second riser reactor are straight pipes with equal diameters.

According to a preferred embodiment of the present invention, the mass ratio of the heavy feedstock oil to the light feedstock oil is 1 to 5:1, preferably 2 to 5:1. the method provided by the invention can process a relatively large amount of light raw oil.

According to the invention, the circulation amount of the spent catalyst can be determined according to the processing amount of the heavy oil and the water amount injected into the stripper, and preferably, the circulation amount of the spent catalyst is 2-5 tons of catalyst per ton of heavy raw oil.

Preferably, a second cooling medium is introduced into the stripper and part of the stripped spent catalyst is recycled to the first riser reactor such that the bottom temperature of the first riser reactor is 610-670 ℃.

Preferably, the mass ratio of the second regenerated catalyst to the spent catalyst recycled to the first riser reactor is 2-4:1.

according to a preferred embodiment of the present invention, the mass ratio of the heavy feedstock oil to the light feedstock oil is 0.5 to 2:1. more preferably, the method further comprises: dividing the light raw oil into a light raw oil A and a light raw oil B, and introducing the light raw oil B into a first riser reactor; wherein the mass ratio of the total amount of the light raw oil B and the heavy raw oil to the light raw oil A is 4-5:1. in this preferred embodiment, when a larger amount of light raw oil is processed, the light raw oil is separated into two portions, light raw oil a and light raw oil B, and the light raw oil B is fed to the first riser reactor to be processed.

According to the invention, preferably, the light raw oil B is contacted with the first regenerated catalyst to carry out catalytic cracking reaction to obtain a mixture, and then the heavy raw oil is introduced into the first riser reactor to be contacted with the mixture to carry out reaction to obtain the first oil mixture. In this preferred embodiment, the light feedstock oil B is first contacted with the first regenerated catalyst to react, and then the heavy feedstock oil is introduced, and the mixture obtained by the contact reaction of the heavy feedstock oil with the light feedstock oil B and the first regenerated catalyst is further contacted to react. By adopting the preferred embodiment, the light raw oil is cracked by utilizing the high-temperature, high-activity and high-density catalyst reaction zone at the bottom of the first riser reactor, and meanwhile, the contact temperature of the heavy raw oil and the catalyst can be reduced. The preferred embodiment is more beneficial to flexibly adjusting the ratio of the catalyst circulation amount of the first riser reactor to the catalyst circulation amount of the second riser reactor, and is more beneficial to further improving the propylene selectivity and simultaneously reducing the yield of dry gas and coke.

According to a preferred embodiment of the present invention, the heavy feed oil is introduced 0.1 to 2 seconds, preferably 0.3 to 1 second after the introduction of the light feed oil B into the first riser reactor.

According to an embodiment of the present invention, the heavy feedstock oil is introduced into the upper portion of the light feedstock oil B, and the introduction of the heavy feedstock oil is to terminate the reaction of the light feedstock oil B, and the reaction time of the light feedstock oil B is to be the delayed introduction time of the heavy feedstock oil to the light feedstock oil B.

According to the invention, specifically, the third oil agent mixture is subjected to gas-solid separation, and the separated oil gas is sent to a subsequent separation system. The gas-solid separation may be carried out in a settler. The separated carbon deposited catalyst is introduced into a stripper for steam stripping.

The stripping method in step (3) is not particularly limited, and the stripping can strip out the adsorbed hydrocarbon products, and the stripping can be a conventional method in the art, and the stripping is not particularly limited, and is well known to those skilled in the art and will not be described herein again.

The regeneration method in the step (3) is not particularly limited in the present invention, and various regeneration methods conventionally used in the art, such as coke-burning regeneration, may be used. Preferably, the temperature of the regeneration is 660 to 720 ℃.

According to the present invention, the heavy raw oil and the light raw oil have meanings conventionally explained in the art, and preferably, the heavy raw oil is at least one selected from the group consisting of vacuum gas oil, atmospheric gas oil, coker gas oil, deasphalted oil, vacuum residue, atmospheric residue, hydrogenated heavy oil, oil sand oil, shale oil, synthetic oil, animal and vegetable fats and oils, and coal liquefied oil. Specifically, the distillation range of the heavy raw oil is above 300 ℃.

According to the present invention, the light raw oil is preferably a hydrocarbon fraction having 4 to 20 carbon atoms and a boiling point of 300 ℃ or lower, and more preferably, the light raw oil is at least one selected from the group consisting of a catalytically cracked C4 hydrocarbon fraction, a catalytically cracked light gasoline fraction, straight run naphtha and straight run diesel.

The catalyst of the present invention is not particularly limited in kind, and may be various catalytic cracking catalysts conventionally used in the art. Specifically, the first regenerated catalyst and the second regenerated catalyst contain molecular sieves. It is preferred for the present invention that the molecular sieve is selected from at least one of Y zeolite, ZSM-5 zeolite, high-silica zeolite having a pentasil structure, and beta type zeolite. More preferably, the molecular sieve selects at least one of Y or HY type zeolite with or without rare earth, ultrastable Y type zeolite with or without rare earth, ZSM-5 series zeolite, high silica zeolite having a pentasil structure, and beta zeolite.

The catalyst can be obtained commercially or prepared by the existing method.

The present invention will be described in detail below by way of examples.

The catalytic cracking catalysts used in the examples and comparative examples were industrially produced by the Qilu Branch of catalyst, of petrochemical Co., ltd., china, under the trade name of MMC-2. The catalyst contains ultrastable Y-type zeolite and ZSP zeolite with average pore diameter less than 0.7 nm, and is hydrothermally aged for 17 hours at 800 ℃ by saturated steam before use, and the main physicochemical properties of the catalyst are shown in Table 1. The properties of the heavy feed oil and the light feed oil used in examples and comparative examples are shown in tables 2 and 3.

TABLE 1 physicochemical Properties of the catalyst

TABLE 2

TABLE 3

Example 1

The experiment was carried out using a modified medium-sized apparatus for continuous reaction-regeneration operation, the flow of which is shown in fig. 1, the medium-sized apparatus having a first riser reactor 1 with an inner diameter of 16 mm and a length of 3800 mm, a second riser reactor 2 with an inner diameter of 12 mm and a length of 3200 mm, a fluidized bed reactor 3 with an inner diameter of 64 mm and a height of 600 mm. The high-temperature regenerated catalyst with the temperature of 700 ℃ is respectively introduced into the bottoms of the first riser reactor 1 and the second riser reactor 2 from the regenerator 6 through a first regenerated catalyst conveying pipe 11 and a second regenerated catalyst conveying pipe 21, and flows upwards under the action of pre-lifting steam. After being mixed with atomized water vapor, the heavy raw oil enters the first riser reactor 1 through the first feeding nozzle 12 to contact with a hot regenerant for catalytic conversion reaction, and oil gas after the reaction is introduced into the fluidized bed reactor through the outlet 13 of the first riser reactor for reaction. Water at 95 c is injected into the first riser reactor outlet through the first cooling medium feed nozzle 14. The light raw oil enters the lower part of a second riser reactor 2 through a second feeding nozzle 22 under an atomized water vapor medium to contact with a hot regenerant for catalytic conversion reaction, a mixture of reaction oil gas and a catalyst ascends along the second riser reactor 2 and enters a fluidized bed reactor through a second riser reactor outlet 23 to continuously participate in cracking reaction, the reaction oil is introduced into a settler 4 for oil separation, and the reaction oil gas is introduced into a product separation system to be separated into gas and liquid products. The spent catalyst containing coke from the fluidized bed reactor 3 enters a stripper 5, and the steam stripping steam is used for stripping hydrocarbon products adsorbed on the spent catalyst and then enters a settler through the fluidized bed reactor for gas-solid separation. The stripped spent catalyst enters a regenerator 6 through a spent catalyst conveying pipe 51 and contacts with air to be coked and regenerated at the high temperature of 700 ℃. The regenerated catalyst returns to the riser reactor for recycling through a first regenerated catalyst conveying pipe 11 and a second regenerated catalyst conveying pipe 21 respectively. The medium-sized devices use electrical heating to maintain the temperature of the reaction-regeneration system.

The main operating conditions and results are listed in table 4.

Comparative example 1

The process of example 1 was followed except that no cooling medium was injected into the outlet of the first riser reactor, resulting in an increase in the temperature of the outlet of the first riser reactor and an increase in the temperature of the bed of the fluidized bed. The reaction apparatus used was the same as in example 1, and the raw materials, catalysts and main experimental steps used were the same as in example 1, except that 95 ℃ water was not injected into the outlet of the first riser reactor.

The main operating conditions and results are listed in table 4.

Comparative example 2

The process described in comparative example 1 was followed except that the severity of the first riser reactor was reduced. The reaction apparatus used was the same as in comparative example 1, and the raw materials, catalysts and main experimental procedures used were the same as in example 1, except that the first riser reactor catalyst-to-oil ratio was decreased.

The main operating conditions and results are listed in table 4.

TABLE 4

Item	Example 1	Comparative example 1	Comparative example 2
				Heavy raw oil treatment amount/(kg/h)	10	10	10
Light raw oil treatment amount/(kg/h)	3	3	3
				Outlet temperature/. Degree.C. of the first riser	560	580	530
First riser agent to oil ratio	10	10	7
				First riser reaction time/s	1.5	1.5	1.5
Second riser outlet temperature/° c	620	620	620
				Second riser agent-to-oil ratio	15	15	15
Second riser reaction time/s	2	2	2
				Fluidized bed reaction temperature/° c	560	575	560
Fluidized bed weight hourly space velocity/h ^-1	6	6	6
				95 ℃ water injection/(kg/h)	0.8	0	0
Distribution of the product/%)
				Dry gas	9.9	12.3	10.4
Liquefied gas	41.8	40.1	39.5
				Gasoline (gasoline)	26.7	25.4	24.8
Diesel oil	9.2	8.9	11.8
				Heavy oil	3.3	3.1	4.2
Coke	9.1	10.2	9.3
				Total of	100	100	100
Propylene yield/%)	21.5	20.4	19.3
				Propylene selectivity	0.245	0.232	0.230

Example 2

The experiment was carried out using a modified medium-sized apparatus for continuous reaction-regeneration operation, the flow of which is shown in fig. 2, the medium-sized apparatus having a first riser reactor 1 with an inner diameter of 16 mm and a length of 3800 mm, a second riser reactor 2 with an inner diameter of 12 mm and a length of 3200 mm, a fluidized bed reactor 3 with an inner diameter of 64 mm and a height of 600 mm. The bottom of the first riser reactor 1 is provided with a catalyst mixer 7, 690 ℃ high-temperature regenerated catalyst is introduced into the catalyst mixer through a first regenerated catalyst conveying pipe 11, part of the spent catalyst is introduced into the catalyst mixer through a spent catalyst circulating pipe 52, and the catalyst flows upwards under the action of pre-lifting steam. After being mixed with atomized water vapor, the heavy raw oil enters the first riser reactor 1 through the first feeding nozzle 12 to contact with a catalyst from the catalyst mixer for catalytic conversion reaction, and oil gas after the reaction is introduced into the fluidized bed reactor through the outlet 13 of the first riser reactor for reaction. Part of 690 ℃ high-temperature regenerated catalyst is introduced into the bottom of the second riser reactor 2 from the regenerator 6 through a second regenerated catalyst conveying pipe 21, light raw oil enters the lower part of the second riser reactor 2 through a second feeding nozzle 22 under an atomized water vapor medium to contact with a hot regenerant to perform catalytic conversion reaction, a mixture of reaction oil gas and catalyst ascends along the second riser reactor 2 and enters the fluidized bed reactor through an outlet 23 of the second riser reactor to continuously participate in cracking reaction, the reaction oil gas is introduced into a settler 4 to perform oil separation, and the reaction oil gas is introduced into a product separation system to be separated into gas and liquid products. The spent catalyst containing coke from the fluidized bed reactor 3 enters a stripper 5, and the steam stripping steam is used for stripping hydrocarbon products adsorbed on the spent catalyst and then enters a settler through the fluidized bed reactor for gas-solid separation. Water with the temperature of 95 ℃ is sprayed into the stripper through a second cooling medium feeding nozzle 53, and the stripped spent catalyst enters the regenerator 6 through a spent catalyst conveying pipe 51 and contacts with air to be burnt and regenerated at the high temperature of 690 ℃. The regenerated catalyst returns to the riser reactor for recycling through a first regenerated catalyst conveying pipe 11 and a second regenerated catalyst conveying pipe 21 respectively. The medium-sized devices use electrical heating to maintain the temperature of the reaction-regeneration system.

The main operating conditions and results are listed in table 5.

Comparative example 3

The process described in example 2 was followed, except that the spent catalyst was not introduced into the bottom of the first riser reactor 1. The reaction apparatus used was the same as in example 2, and the raw materials, catalysts and main experimental steps used were the same as in example 2, except that the outlet temperature of the first riser reactor 1 was maintained by decreasing the solvent-to-oil ratio of the first riser reactor 1.

The main operating conditions and results are listed in table 5.

Comparative example 4

The process described in comparative example 3 was followed, except that no cooling medium was introduced into the stripper. The reaction apparatus used was the same as in comparative example 3, and the starting materials, catalysts and main experimental steps used were the same as in example 2, except that 95 ℃ water was not injected into the stripper. The main operating conditions and results are listed in table 5.

TABLE 5

Item	Example 2	Comparative example 3	Comparative example 4
				Heavy raw oil treatment amount/(kg/h)	10	10	10
Light raw oil treatment amount/(kg/h)	4	4	4
				Catalyst Mixer temperature/. Degree.C	640	690	660
Outlet temperature/. Degree.C. of the first riser	560	560	580
				First riser reaction time/s	1.5	1.5	1.5
First lift pipe regenerated catalyst circulation volume/(kg/h)	150	100	150
				Spent catalyst circulation volume/(kg/h)	50	0	50
Second riser outlet temperature/° c	630	630	630
				Second riser reaction time/s	2	2	2
Second lift pipe regenerated catalyst circulation volume/(kg/h)	80	80	80
				95 ℃ water injection/(kg/h)	0.8	0.8	0
Fluidized bed reaction temperature/° c	560	590	582
				Fluidized bed weight hourly space velocity/h ^-1	6	6	6
Distribution of products/%)
				Dry gas	9.8	12.4	12.5
Liquefied gas	40.3	38.9	39.5
				Gasoline (gasoline)	28.2	26.6	25.9
Diesel oil	9.5	9.2	9.4
				Heavy oil	3.3	3.1	3.2
Coke	8.9	9.8	9.5
				Total of	100	100	100
Propylene yield/%)	21.8	20.8	21.2
				Propylene selectivity	0.250	0.237	0.242

Example 3

The experiment was carried out using a modified medium-sized apparatus for continuous reaction-regeneration operation, the flow of which is shown in fig. 1, the first riser reactor 1 of which inner diameter is 16 mm and length is 3800 mm, and the heavy feedstock nozzle (not shown) is 600 mm above the light feedstock nozzle B (first feed nozzle 12). The second riser reactor 2 has an inner diameter of 12 mm and a length of 3200 mm, and the fluidized bed reactor 3 has an inner diameter of 64 mm and a height of 600 mm. The high-temperature regenerated catalyst with the temperature of 700 ℃ is respectively introduced into the bottoms of the first riser reactor 1 and the second riser reactor 2 from the regenerator 6 through a first regenerated catalyst delivery pipe 11 and a second regenerated catalyst delivery pipe 21, and flows upwards under the action of pre-lifting steam. After being mixed with atomized water vapor, the light raw oil B enters the first riser reactor 1 through the first feeding nozzle 12 to contact with a hot regenerant for catalytic conversion reaction, the heavy raw oil is sprayed into the first riser reactor 1 through the heavy raw oil nozzle (not shown) under the medium of atomized water vapor to contact with a mixture formed by contact reaction of the light raw oil B and the hot regenerant for reaction, and oil gas after the reaction is introduced into the fluidized bed reactor through the outlet 13 of the first riser reactor for reaction. Water at 95 c is injected into the outlet of the first riser reactor through the first cooling medium feed nozzle 14. The light raw oil A enters the lower part of a second riser reactor 2 through a second feeding nozzle 22 under an atomized water vapor medium to contact with a hot regenerant for catalytic conversion reaction, a mixture of reaction oil gas and a catalyst ascends along the second riser reactor 2 and enters a fluidized bed reactor through a second riser reactor outlet 23 to continuously participate in cracking reaction, the reaction oil is introduced into a settler 4 for oil separation, and the reaction oil gas is introduced into a product separation system to be separated into gas and liquid products. The spent catalyst containing coke from the fluidized bed reactor 3 enters a stripper 5, and the steam stripping steam is used for stripping hydrocarbon products adsorbed on the spent catalyst and then enters a settler through the fluidized bed reactor for gas-solid separation. The stripped spent catalyst enters a regenerator 6 through a spent catalyst conveying pipe 51 and contacts with air to be coked and regenerated at the high temperature of 700 ℃. The regenerated catalyst returns to the riser reactor for recycling through a first regenerated catalyst conveying pipe 11 and a second regenerated catalyst conveying pipe 21 respectively. The medium-sized devices use electrical heating to maintain the temperature of the reaction-regeneration system.

The main operating conditions and results are listed in table 6.

TABLE 6

Item	Example 3
		Heavy raw oil treatment amount/(kg/h)	10
Light raw oil A treatment amount/(kg/h)	4
		Light raw oil B treatment amount/(kg/h)	6
Outlet temperature/. Degree.C. of the first riser	570
		First riser agent to oil ratio	10
Reaction time/s of light raw oil B	0.8
		Reaction time/s of heavy raw oil	1.5
Second riser outlet temperature/° c	630
		Second riser agent to oil ratio	15
Second riser reaction time/s	2
		Fluidized bed reaction temperature/° c	580
Fluidized bed weight hourly space velocity/h ^-1	6
		95 ℃ water injection/(kg/h)	1.5
Distribution of the product/%)
		Dry gas	11.2
Liquefied gas	40.7
		Gasoline (gasoline)	32.9
Diesel oil	5.2
		Heavy oil	2.8
Coke	7.2
		Total of	100
Propylene yield/%)	21.8
		Propylene selectivity	0.236

It can be seen from the above data of the examples that the propylene selectivity can be significantly improved and the yields of dry gas and coke can be reduced by using the method for producing propylene by catalytic cracking of hydrocarbons provided by the present invention.

The preferred embodiments of the present invention have been described above in detail, but the present invention is not limited thereto. Within the scope of the technical idea of the invention, many simple modifications can be made to the technical solution of the invention, including various technical features being combined in any other suitable way, and these simple modifications and combinations should also be regarded as the disclosure of the invention, and all fall within the scope of the invention.

Claims

1. A process for the catalytic cracking of hydrocarbons to produce propylene, the process comprising:

the mass ratio of the heavy raw oil to the light raw oil is 0.5-5:1;

the method further comprises the following steps: introducing a first cooling medium to an outlet of the first riser reactor; and/or introducing a second cooling medium into the stripper, and recycling part of the stripped spent catalyst to the first riser reactor;

the mass ratio of the first regenerated catalyst to the second regenerated catalyst is 2-4:1;

the temperature of the first cooling medium and the second cooling medium is 20-99 ℃ respectively and independently;

the introduction amount of the first cooling medium and the second cooling medium is respectively and independently 50-200 kg/ton of heavy raw oil;

the operating conditions of the first riser reactor include: the agent-oil ratio is 8-20, the oil-gas retention time is 0.5-5s, and the outlet temperature is 550-580 ℃;

the operating conditions of the second riser reactor include: the outlet temperature is 610-650 ℃, and the oil gas retention time is 0.1-4s;

the operating conditions of the fluidized bed reactor include: the reaction temperature is 560-580 ℃, and the weight hourly space velocity is 2-8h ^-1 ；

The circulation quantity of the spent catalyst is 2-5 tons of catalyst per ton of heavy raw oil.

2. The method of claim 1, wherein the first cooling medium and the second cooling medium are each independently water and/or gasoline.

3. The method of claim 1, wherein,

the introduced amount of the first cooling medium and the second cooling medium is respectively and independently 60-120 kg/ton of heavy raw oil.

4. The process according to any one of claims 1 to 3, wherein the mass ratio of the heavy feedstock oil to the light feedstock oil is 1 to 5:1.

5. the process according to claim 4, wherein the mass ratio of the heavy raw oil to the light raw oil is 2-5:1.

6. a process according to any one of claims 1 to 3, wherein a second cooling medium is introduced into the stripper and part of the stripped spent catalyst is recycled to the first riser reactor such that the bottom temperature of the first riser reactor is in the range of 610 to 670 ℃.

7. The process according to any one of claims 1 to 3, wherein the mass ratio of the second regenerated catalyst to the spent catalyst recycled to the first riser reactor is from 2 to 4:1.

8. the process according to any one of claims 1 to 3, wherein the mass ratio of the heavy feedstock oil to the light feedstock oil is from 0.5 to 2:1.

9. the method of claim 8, wherein the method further comprises: dividing the light raw oil into a light raw oil A and a light raw oil B, and introducing the light raw oil B into a first riser reactor; wherein the mass ratio of the total amount of the light raw oil B and the heavy raw oil to the light raw oil A is 4-5:1.

10. the process according to claim 9, wherein the light raw oil B is contacted with the first regenerated catalyst to perform catalytic cracking reaction to obtain a mixture, and then the heavy raw oil is introduced into the first riser reactor to perform contact reaction with the mixture to obtain the first oil mixture.

11. The process according to claim 10, wherein the heavy feed oil is introduced 0.1-2 seconds after the light feed oil B is introduced into the first riser reactor.

12. A process according to any one of claims 1 to 3, wherein the temperature of the regeneration is 660 to 720 ℃.

13. The process of any of claims 1-3, wherein the heavy feedstock oil is selected from at least one of vacuum gas oil, atmospheric gas oil, coker gas oil, deasphalted oil, vacuum residue, atmospheric residue, hydrogenated heavy oil, oil sand oil, shale oil, synthetic oil, animal and vegetable fats and oils, and coal liquefied oil;

the light raw oil is at least one selected from catalytic cracking C4 hydrocarbon fraction, catalytic cracking light gasoline fraction, straight-run naphtha and straight-run diesel.

14. The process according to any one of claims 1 to 3, wherein the first regenerated catalyst and the second regenerated catalyst contain a molecular sieve selected from at least one of Y zeolite, ZSM-5 zeolite, high-silica zeolite having a pentasil structure, and beta zeolite.