CN110951501A

CN110951501A - Catalytic conversion method of low-coke-formation raw material

Info

Publication number: CN110951501A
Application number: CN201811132827.2A
Authority: CN
Inventors: 崔守业; 白旭辉; 李泽坤
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: 2018-09-27
Filing date: 2018-09-27
Publication date: 2020-04-03
Anticipated expiration: 2038-09-27
Also published as: CN110951501B

Abstract

The invention relates to a catalytic conversion method of a low-coke-formation raw material, which comprises the following steps: (1) feeding the low-coke-formation raw material into a first reaction zone of a catalytic conversion reactor to contact with a catalytic conversion catalyst and perform a first catalytic conversion reaction; (2) continuously carrying out a second catalytic conversion reaction on the obtained first reaction product and the semi-spent catalyst in a second reaction zone of the catalytic conversion reactor to obtain a second reaction product and a spent catalyst; the oxygenate is fed to a second reaction zone of the catalytic conversion reactor to contact and carry out a second catalytic conversion reaction with the first reaction product and the semi-spent catalyst. The method of the invention can improve the heat distribution of the catalytic conversion reactor and improve the distribution of reaction products.

Description

Catalytic conversion method of low-coke-formation raw material

Technical Field

The invention relates to a catalytic conversion method of a low-coke-formation raw material.

Background

Lower olefins (C)₂-C₄Olefins) have been the dominant primary organic chemicals in the modern petroleum and chemical industries, especially ethylene and propylene. The methods for producing light olefins can be broadly divided into two major categories, namely the traditional petroleum route and the emerging non-petroleum route. The traditional method for preparing low-carbon olefin by petroleum route mainly comprises steam cracking and catalytic cracking process production. Since the 10 s in the 20 th century, various countries in the world have been dedicated to developing routes for preparing low-carbon olefins from non-petroleum resources, and some progress has been made.

Chinese patent CN92109905 discloses a catalyst and a reaction process for converting methanol into light olefin, wherein a ZSM-5 type zeolite catalyst containing phosphorus, rare earth elements and a pore structure regulator and a reaction process of a multistage adiabatic fixed bed cracking reactor adopting a dehydration reactor and 2-n reaction-regeneration switching operations are adopted, the non-cyclic operation is carried out at high temperature (more than 400 ℃), and the catalyst has high activity, high selectivity, high water resistance, high thermal stability and long reaction life. On a plant scale of 0.7 to 1 ton of methanol per day, the methanol conversion is 100%, C₂-C₄The olefin selectivity can be greater than 85%, the on-stream time can be greater than 600 hours, and the once-through operating period can be greater than 24 hours. However, because the heat transfer of the fixed bed reactor is slow, the reaction of preparing light olefins from methanol is a strong exothermic reaction, and hot spots are easy to occur, so that the device is damaged.

Chinese patent CN200810043239.1 discloses a method for producing low-carbon olefin from methanol, which mainly solves the problem of low selectivity of target products in the process of preparing low-carbon olefin from methanol. The method comprises the following steps: (a) heating an oxygen-containing compound raw material containing methanol at a reaction temperature of 300-600 ℃ and a raw material weight hourly space velocity of the oxygen-containing compound of 1-50 hoursTime of flight^-1Under the condition that the reaction pressure (gauge pressure) is 0.05-10 MPa, the raw materials are contacted with a silicoaluminophosphate molecular sieve catalyst in a reactor; (b) separating the catalyst from the reaction products; (c) exchanging heat between the reaction product and the raw material containing the methanol, and heating the raw material containing the methanol to 100-350 ℃ under the condition of ensuring effective recovery of heat of the reaction product; (d) the technical scheme of (a) to (c) is repeated, so that the problem is solved well, and the method can be used for industrial production of low-carbon olefin.

WO2006049864 discloses a process and corresponding apparatus for the production of light olefins from oxygenates, wherein the process comprises feeding an oxygenate feed stream through a feed stream distributor (3) into an OTO reactor; contacting the oxygenate with a catalyst to produce a mixture comprising light olefins, unreacted oxygenate, and other by-products; separating unreacted oxygenate and diolefins from said light olefins and said by-products; and returning unreacted oxygenate and diolefins to the OTO reactor. Unreacted oxygenate and diolefins are fed to the reactor through at least one feed nozzle at a point separate from the oxygenate feed stream. The process is believed to be able to carry out the oxygenate conversion over a wide range of pressures (10.1kPa to 10.1MPa), but not all pressures give good results, preferably from 101.3kPa to 1013.3 kPa.

After the dehydration reaction of the oxygen-containing compound to prepare the low-carbon olefin, a certain amount of water is also by-produced in addition to the hydrocarbon product, in the case of methanol, about 44% of the hydrocarbon product can be obtained after the reaction, and in the case of ethanol, 46% of the hydrocarbon product can be obtained after the reaction. It is known that the reaction of an oxygenate feedstock to produce lower olefins is a reaction in which the number of molecules increases, and that low reaction pressures favor chemical equilibrium toward the production of lower olefins. In view of this, in the prior art, lower reaction pressure is generally used to obtain the desired yield of lower olefins. The direct consequence of this low reaction pressure (typically 0.1-0.3MPa) is that if it is desired to increase the oxygenate feedstock throughput for the purpose of increasing the production of lower olefins, the prior art has had to increase the size or number of reactors for this purpose in order to maintain the yield of lower olefins at an acceptable level. Obviously, this would correspondingly increase the investment and maintenance costs of the equipment.

Because the reaction process of preparing the low-carbon olefin from the oxygen-containing compound is an exothermic reaction, if pure methanol is used for feeding, the total reaction heat of the methanol-to-olefin is generally 20-35 KJ/mol, the adiabatic temperature rise is above 200 ℃ at a small water-alcohol ratio, and if a side reaction in the MTO reaction process is considered, the adiabatic temperature rise is larger. Such high temperature rise not only affects the MTO reaction result and accelerates the carbon deposition rate of the catalyst, but also needs to consider the hydrothermal stability of the catalyst. Therefore, to reduce the temperature rise in the reactor, measures such as designing a heat removal system, reducing the initial composition of reactants, and reducing the temperature of the feed are generally required.

Chinese patent CN201110253681.9 discloses a method for reducing energy consumption of an apparatus for producing olefin from oxygen-containing compound. Preheating a liquid oxygen-containing compound raw material to a certain temperature by a raw material preheater, dividing the raw material into two strands, and heating and gasifying one strand by a raw material vaporizer to obtain a gas-phase raw material; the other is atomized into atomized liquid phase raw material; mixing a gas-phase raw material and an atomized liquid-phase raw material in front of a raw material/reaction gas heat exchanger, feeding the mixture into the raw material/reaction gas heat exchanger in a mist flow mode, fully exchanging heat with oil gas generated by high-temperature reaction from a reactor to recover high-temperature-level heat of the oil gas generated by the high-temperature reaction, completely gasifying the raw material after recovering the high-temperature-level heat to form high-temperature gasification raw material gas, and feeding the high-temperature gasification raw material gas into the reactor for reaction; oil gas generated by high-temperature reaction from the reactor is sent to a rear quenching water washing system after heat exchange of the raw material/reaction gas heat exchanger. The method can effectively improve the heat exchange effect of the heat exchanger, reduce the volume of the heat exchanger and reduce the energy consumption of the device, but the method does not solve the problem of cooling the reaction catalyst.

At present, the reactors adopted by the technology for preparing low-carbon olefin by using oxygen-containing compounds mainly comprise a fixed fluidized bed reactor, a dense-phase fluidized bed reactor and the like. The fixed bed reactor has poor bed heat transfer effect, and for strong exothermic reaction, if the heat released in the reaction process can not be transferred out in time, hot spots, temperature runaway and other problems are easy to occur, and the replacement and regeneration of the catalyst are relatively troublesome. The diameter and height of the fixed bed reactor are not strictly limited, but the height/diameter ratio adopted by the reactor design is 2.5-12 in consideration of the factors of fluid distribution, manufacturing cost, safety and the like. The fixed bed reactor is also adopted in the early catalytic cracking reaction, carbon deposition can occur on the catalyst in the catalytic cracking reaction, and the regeneration reaction is needed to recover the activity of the catalyst, so that the fixed bed reactor is needed to be discontinuously used for reaction and regeneration, and at least more than three reactors are needed to ensure the continuity of continuous feeding and other operations. The dense-phase fluidized bed reactor belongs to the category of bubbling bed and turbulent bed, the reaction residence time is generally longer, the residence time of the generated olefin in the reactor is longer, the hydrogen transfer reaction is increased, and the dense-phase fluidized bed reactor is also very unfavorable for the high yield of low-carbon olefin. Due to the low linear velocity of the dense phase fluidized bed reactor, back mixing tends to occur, affecting product distribution and quality, and on the other hand the diameter of the reactor is relatively large.

Since the birth of the 40 th generation of the 20 th century, catalytic cracking has been the most important process for lightening heavy oil. The reason is that the raw material source is wide, wax oil can be adopted, deasphalted oil of atmospheric residue oil and vacuum residue oil can also be adopted, or vacuum residue oil can be partially mixed; secondly, the product scheme is flexible, and can be fuel type or fuel chemical type, such as productive gasoline, productive diesel oil, productive propylene and the like; and thirdly, the product properties can be correspondingly adjusted through adjustment of a catalyst formula and change of process parameters, such as improvement of the octane number of the gasoline, reduction of the olefin content of the gasoline and the like. In recent years, with the development of the fischer-tropsch synthesis technology, how to utilize the fischer-tropsch synthesis oil is gradually paid attention by technical personnel, and the conversion of the fischer-tropsch synthesis oil by a catalytic cracking method is one of the ways.

Chinese patent CN201410543443.5 discloses a method for producing low-carbon olefins from hydrocarbon oil raw material, which comprises: the hydrocarbon oil raw material contacts with a catalytic cracking catalyst to carry out catalytic cracking reaction, wherein the hydrocarbon oil raw material contains 10-25 wt% of heavy distillate oil and 75-90 wt% of Fischer-Tropsch synthetic oil. According to the method for producing the low-carbon olefin from the hydrocarbon oil raw material, after the heavy distillate oil is doped into the Fischer-Tropsch synthetic oil, the generation of the low-carbon olefin can be promoted, and meanwhile, the heat balance of a self reaction-regeneration system can be met.

Chinese patent CN201180056719.1 discloses a process for converting heavy feedstocks into middle distillates, which requires a catalytic cracking unit, followed by one or more oligomerization units of olefins with a number of carbon atoms from C2 to C9, which may preferably produce an additional fraction called middle distillate. The light portion of the resulting oligomeric product that cannot be incorporated into the middle distillate is recycled to the FCC for cracking into light olefins that are returned to the oligomerization unit in addition to the olefins from the feedstock to preferably form heavy oligomeric products that can be incorporated into the middle distillate.

The heat required by the current catalytic conversion reactor is mainly provided by the heat regenerated catalyst circulated back, and the purpose can be achieved only by increasing the heat at the bottom of a riser in order to increase the reaction temperature, but the defects are that the contact temperature of the catalyst and the oil is too high, or the catalyst-oil ratio is too large, so that the yield of dry gas and coke is obviously increased, and the product distribution is poor. In addition, because the riser reactor dissipates heat, the raw material is vaporized, the hydrocarbon reacts, the temperature of the riser reactor is obviously reduced along with the increase of the height of the riser, particularly when the variable-diameter riser is adopted, the temperature of the secondary reaction is reduced by 10-15 ℃ compared with that of the primary reaction, and the cracking reaction of the secondary reaction is unfavorable. For high-quality raw materials, such as light Fischer-Tropsch synthetic oil, light hydrogenated diesel oil, catalytic gasoline, naphtha and the like, which are used as catalytic conversion raw materials, the reaction-regeneration heat balance cannot be met due to low coke formation, and the conventional method comprises the step of reacting the Fischer-Tropsch synthetic oil and the heavy raw materials together, so that the coke yield is improved; the regenerator is sprayed with combustion oil or combustion gas.

Disclosure of Invention

The invention aims to provide a catalytic conversion method of low-coke-formation raw materials, which can improve the heat distribution of a catalytic conversion reactor and improve the distribution of reaction products.

The invention provides a catalytic conversion method of a low-coke-formation raw material, which comprises the following steps:

(1) feeding the low-coke-formation raw material into a first reaction zone of a catalytic conversion reactor to contact with a catalytic conversion catalyst and perform a first catalytic conversion reaction to obtain a first reaction product and a semi-spent catalyst; wherein the low raw coke raw material is one or more selected from Fischer-Tropsch synthetic oil, hydrogenated diesel oil, hydrocracking tail oil, straight-run naphtha, straight-run diesel oil, catalytic gasoline, coker gasoline and low-carbon hydrocarbons; according to the flow direction of reaction materials, the catalytic conversion reactor sequentially comprises a first reaction zone and a second reaction zone which are communicated with each other by fluid;

(2) continuously carrying out a second catalytic conversion reaction on the obtained first reaction product and the semi-spent catalyst in a second reaction zone of the catalytic conversion reactor to obtain a second reaction product and a spent catalyst, and sending the second reaction product and the spent catalyst out of the catalytic conversion reactor through an outlet of the second reaction zone;

(3) feeding the obtained spent catalyst into a regenerator for regeneration, and returning the obtained regenerated catalyst serving as the catalytic conversion catalyst to the catalytic conversion reactor;

the method further comprises the following steps: in step (2), the oxygenate is fed through the oxygenate inlet into a second reaction zone of the catalytic conversion reactor to contact and carry out a second catalytic conversion reaction with the first reaction product and the semi-spent catalyst.

In some embodiments of the present invention, under the same catalyst-to-oil ratio condition, the method of the present invention can reduce the contact temperature of the regenerated catalyst and the raw material by more than 10 ℃, reduce thermal cracking reaction, reduce temperature difference in the second catalytic conversion reaction process, supplement heat required for cracking for the riser reactor, especially for the second reaction zone of the variable diameter riser reactor, reduce the partial pressure of oil gas and hydrocarbon, enhance selective cracking reaction, hydrogen transfer reaction and isomerization reaction, and improve the yield of low carbon olefins such as ethylene and propylene. It has been found through experiments that a better reaction effect can be obtained by controlling the temperature at the outlet of the second reaction zone to be 0.3 to 100 c higher, preferably 0.3 to 50 c higher, more preferably 0.5 to 20 c higher, and further preferably 0.6 to 8 c higher than the temperature at the inlet of the oxygen-containing compound (the temperature of the reaction zone).

In some embodiments of the invention, the first reaction zone and the second reaction zone are divided by the location of oxygenate injection, with the reaction zone upstream of the location of oxygenate injection being the first reaction zone and the reaction zone downstream of the location of oxygenate injection being the second reaction zone, in terms of reactant feed flow direction.

In some embodiments of the invention, the invention optimizes product selectivity by feeding an oxygenate to the reactor to improve reaction conditions in the latter half of the reactor. The catalytic conversion reaction is used for cracking the low-coke-formation raw material and the oxygen-containing compound under certain reaction conditions and the like to generate products such as dry gas, liquefied gas, gasoline and the like. The conditions of the first catalytic conversion reaction may include: the temperature (the temperature of the oxygen-containing compound feeding part) is 450-620 ℃, preferably 490-600 ℃, the reaction time is 0.5-2 seconds, preferably 0.8-1.5 seconds, the weight ratio of the catalytic conversion catalyst to the low-coke-formation raw material is 3-15, preferably 3-12, the reaction pressure is 130-450 kPa, and the weight ratio of the water vapor to the low-coke-formation raw material is 0.03-0.3; the conditions of the second catalytic conversion reaction may include: the temperature (temperature at the outlet of the second reaction zone) is 460-650 ℃, preferably 500-620 ℃ and the reaction time is 2-30 seconds, preferably 3-15 seconds.

In some embodiments of the present invention, the catalytic conversion reaction of the oxygenate produces water, which helps to reduce the partial pressure of the hydrocarbon reaction, the weight of the water produced by the second catalytic conversion reaction of the oxygenate is W1 (based on the total conversion of oxygen in the oxygenate to water), and the weight ratio of W1 to the steam injected in the first catalytic conversion reaction may be 200-500 wt%, preferably 300-400 wt%.

In some embodiments of the present invention, it is preferred to control the temperature prior to the introduction of the oxygenate into the catalytic conversion reactor in order to control the temperature of the second catalytic conversion reaction, for example, the temperature of the oxygenate introduced into the catalytic conversion reactor may be in the range of from 20 to 600 ℃, preferably 200 ℃ to 500 ℃, more preferably 360 ℃ to 450 ℃.

In some embodiments of the present invention, the low coke-forming raw material refers to raw material with less coke generated by reaction and less heat generated by regeneration of spent catalyst, wherein the low carbon hydrocarbon refers to hydrocarbons with less than C4, the low coke-forming raw material is preferably fischer-tropsch synthetic oil, and the carbon content of the spent catalyst can be 0.1-3 wt%, preferably 0.2-2 wt%, based on the weight of the spent catalyst.

In some embodiments of the present invention, the oxygen-containing compound is well known to those skilled in the art, and refers to an organic compound containing oxygen atoms, which may also be referred to as an organic oxygen-containing compound, and may be one or more selected from alcohols, ethers, and lipids, preferably one or more selected from methanol, ethanol, methyl ether, ethyl ether, methyl ethyl ether, dimethyl carbonate, and methyl formate, and more preferably methanol.

In some embodiments of the invention, the conditions of the second catalytic conversion reaction may be controlled by adjusting the feed rates of the oxygenate and the low coke forming feedstock, which may be in a ratio of from 0.02 to 50, preferably from 0.05 to 2.

In some embodiments of the present invention, catalytic conversion catalysts are well known to those skilled in the art, for example, the catalytic conversion catalysts may include a molecular sieve, which may be one or more selected from the group consisting of Y-type molecular sieves, ZSM-5 type molecular sieves, SAPO type molecular sieves, β type molecular sieves, and SRM type silicoaluminophosphate molecular sieves, a binder, and a matrix, and other conventional catalysts may also be used by those skilled in the art.

In some embodiments of the present invention, the catalytic conversion catalyst is well known to those skilled in the art, for example, the catalytic conversion reactor can be a riser reactor and/or a fluidized bed reactor, and the riser reactor can be one or more selected from the group consisting of a constant diameter riser, a variable diameter riser, and a constant linear velocity riser.

In some embodiments of the present invention, the inventors have assiduously studied and experimented that injecting an oxygen-containing compound at the rear of a riser reactor can increase the heat of the riser reactor; when the reducing riser reactor is adopted, the oxygen-containing compound can obtain a good conversion environment, the heat loss caused by the cracking and heat dissipation of the conventional hydrocarbons in the environment is matched with the heat provided by the oxygen-containing compound, and the water generated by the reaction of the oxygen-containing compound can reduce the hydrocarbon partial pressure of the second catalytic conversion reaction and improve the product distribution.

In a preferred embodiment, the catalytic conversion reactor may be a riser reactor, the first reaction zone is located below the second reaction zone, and the second reaction zone has an inner diameter greater than that of the first reaction zone.

In some embodiments of the invention, the method may further comprise: at least a portion of the low coke formation feedstock is fed to a second reaction zone for said second catalytic conversion reaction to further improve heat distribution.

In some embodiments of the invention, the method may further comprise: and (3) taking out at least part of the semi-regenerated catalyst containing carbon from the regenerator of the step (3), and sending the semi-regenerated catalyst into the catalytic conversion reactor to carry out the first catalytic conversion reaction and/or the second catalytic conversion reaction, so that the semi-regenerated catalyst can be used for supplementing the reaction heat.

In one embodiment, at the bottom of a conventional equal-diameter riser reactor, a preheated low-coke-formation raw material is contacted with a heat regeneration catalyst in a first reaction zone to perform a cracking reaction, generated oil gas and a carbon-containing catalyst ascend to be contacted with an introduced semi-regeneration catalyst, then a cracking reaction, an isomerization reaction and a hydrogen transfer reaction are performed, then an oxygen-containing compound is injected at the bottom of a second reaction zone, the preheated low-coke-formation raw material is contacted with the carbon-containing semi-regeneration catalyst and a first reaction product, the oxygen-containing compound is dehydrated to generate low-carbon olefin, the preheated low-coke-formation raw material is contacted with the carbon-containing semi-regeneration catalyst to perform the cracking reaction, the isomerization reaction and the hydrogen transfer reaction, and after the reaction, a second reaction product and the spent catalyst enter a settler; separating reaction products, feeding the spent catalyst into a regenerator for burning after steam stripping, and directly returning the regenerated catalyst from the regenerator to the bottom of the reactor without cooling.

In another embodiment, for a catalytic conversion device with a reducing riser reactor, a preheated low-coke-formation raw material enters from the lower part of a first reaction zone of the reactor to contact with a heat regeneration catalyst to perform a cracking reaction, a first reaction product and a semi-spent catalyst generated after the reaction move upwards to the lower part of a second reaction zone of the reactor, an injected oxygen-containing compound is dehydrated to generate low-carbon olefin, heat and water vapor are released, the injected preheated low-coke-formation raw material contacts with the carbon-containing semi-spent catalyst to perform a cracking reaction, an isomerization reaction and a hydrogen transfer reaction, and the second reaction product and the spent catalyst enter a settler after the reaction; separating reaction product, stripping catalyst to be regenerated, and regenerating.

Compared with the prior art, the invention has the advantages that:

1. the low-coke-formation raw material which is subjected to endothermic reaction during catalytic conversion reacts with the oxygen-containing compound which is subjected to exothermic reaction in the same reactor, different raw materials are reacted in a proper reaction zone by adjusting the operating conditions of different reaction zones, more optimized heat distribution is provided for catalytic conversion, cracking reaction is promoted, a uniform fluidized state is provided for the reaction of the oxygen-containing compound, the retention time is shorter, and the product distribution is improved.

2. The heat balance of the reactor and the regenerator is improved, and combustion oil/gas does not need to be injected into the regenerator when the low-coke-formation raw material is catalytically converted.

3. The reaction conditions in the reaction zone are improved, and the selectivity of high-quality products is improved.

Additional features and advantages of the invention will be set forth in the detailed description which follows.

Drawings

The accompanying drawings, which are included to provide a further understanding of the invention and are incorporated in and constitute a part of this specification, illustrate embodiments of the invention and together with the description serve to explain the principles of the invention and not to limit the invention. In the drawings:

FIG. 1 is a schematic diagram of the structure of one embodiment of a reactor used in the process of the present invention.

Description of the reference numerals

I first reaction zone II second reaction zone

1 line 2 riser reactor 3 lines

4 line 5 line 6 line

7 pipeline 8 settler 9 gas collection chamber

10 stripping section 11 pipeline 12 inclined tube

13 regenerator 14 line 15 line

16 inclined tube

Detailed Description

The following detailed description of embodiments of the invention refers to the accompanying drawings. It should be understood that the detailed description and specific examples, while indicating the present invention, are given by way of illustration and explanation only, not limitation.

In FIG. 1, a reducing riser reactor is used, and the shape and size of the equipment and lines are not limited by the attached drawings but determined according to specific conditions.

As shown in fig. 1, a pre-lifting medium enters a first reaction zone I from the bottom of a riser reactor 2 through a pipeline 1, a regenerated catalyst from an inclined pipe 16 moves upwards in an accelerated manner along the riser reactor under the lifting action of the pre-lifting medium, a low-coke-formation raw material is injected into the bottom of the first reaction zone I of the riser reactor 2 through a pipeline 3 together with atomized steam from a pipeline 4, the low-coke-formation raw material performs a first catalytic conversion reaction on a hot regenerated catalyst and moves upwards in an accelerated manner to obtain a semi-spent catalyst and a first reaction product. An oxygen-containing compound is injected into the bottom of a second reaction zone II of the riser reactor 2 through a pipeline 5 together with atomized steam from a pipeline 6, the oxygen-containing compound and a semi-spent catalyst containing certain carbon are dehydrated to generate a second catalytic conversion reaction of olefin, a first to-be-produced substance from a first reaction zone I continues the second catalytic conversion reaction in the second reaction zone II, a generated second reaction product and an inactivated spent catalyst enter a settler 8 through a pipeline 7, the spent catalyst and the second reaction product are separated through a cyclone separator, the second reaction product enters a gas collection chamber 9, and catalyst fine powder returns to the settler through a cyclone separator dipleg. Spent catalyst in the settler flows to the stripping section 10 where it is contacted with steam from line 11. The reaction product oil gas extracted from the spent catalyst enters a gas collection chamber 9 after passing through a cyclone separator. The stripped spent catalyst enters a regenerator 13 through an inclined pipe 12, the main air enters the regenerator through a pipeline 14 to burn off coke on the spent catalyst, the inactivated spent catalyst is regenerated, and the flue gas enters a smoke machine through a pipeline 15. The regenerated catalyst enters the riser reactor through the inclined tube 16. The reaction product oil gas in the gas collection chamber 9 enters a subsequent separation system through a large oil gas pipeline 17.

The invention is further illustrated by the following examples, but is not to be construed as being limited thereto.

The catalysts used in the examples and comparative examples were prepared as follows:

the molecular sieve slurry is prepared by the steps of adding 20 kg of ZSM-5 molecular sieve (ZSP-2, iron oxide content 2.2 wt%, phosphorus pentoxide content 2.5 wt%, product of Chinese petrochemical catalyst Qilu division) in a dry basis (weight basis of molecular sieve, the same below), adding 40 kg of deionized water for pulping, adding 8 kg of USY molecular sieve (DASY2.0, rare earth oxide content 1.8 wt%, unit cell constant 2.445 nm, crystallinity 68%, product of Chinese petrochemical catalyst Qilu division) and then adding 12 kg of SRM molecular sieve (SRM-8, product of Chinese petrochemical catalyst Qilu division) and pulping uniformly to obtain the molecular sieve slurry. Mixing the aluminum sol (commercially available, Al)₂O₃Content of 25 wt%) 80 kg of slurry, adding kaolin (product of China Kaolin Corp.) of 40 kg on a dry basis, uniformly pulping, adding the above molecular sieve slurry, uniformly pulping, spray-drying, and collecting microsphere particles in the range of 20-180 micrometers. The catalyst microspheres were calcined at 500 ℃ for 1 hour and then treated with 1 wt% NH₄And (3) carrying out exchange treatment on the Cl aqueous solution at 60 ℃ for 0.5 hour, filtering, washing with water, and drying at 120 ℃ for 16 hours to obtain the catalyst A adopted in the embodiment and the comparative example of the invention.

Catalyst a consisted of: 20 wt% of ZSP-2 zeolite, 8 wt% of DASY2.0 molecular sieve, 12 wt% of SRM-8 molecular sieve, 20 wt% of alumina sol (calculated by alumina) and 40 wt% of kaolin.

Catalyst A was aged in a fluidized bed aging apparatus at 800 ℃ under 100% steam for 17 hours prior to the catalytic reaction.

The methanol used in the examples and the comparative examples is pure grade methanol analyzed by Beijing chemical plant, and the methanol content is more than or equal to 99.5 weight percent.

Example 1

Example 1 in a specific embodiment the foregoing procedure is carried out in a medium-sized apparatus as shown in figure 1. The low green coke raw material is Fischer-Tropsch synthetic oil fraction, the properties are shown in Table 1, and the catalyst is used as catalyst A. The preheating temperature of the oxygen-containing compound was 148 ℃ and the carbon content of the spent catalyst was 0.25% by weight. The reaction conditions and the product distribution are shown in Table 2.

Comparative example 1

Comparative example 1 was carried out in the medium-sized apparatus shown in fig. 1 according to the aforementioned procedure, except that: the low-coking raw material and the methanol respectively enter the riser reactor from the bottoms of the riser reactors of the two sets of medium-sized devices for reaction, and the product distribution is respectively obtained. The product distributions of the low coke formation feedstock and methanol were multiplied by the respective feed ratios and then summed to obtain the product distribution of comparative example 1. The reaction conditions and the product distribution are shown in Table 2.

The method of the invention can reduce the mixing temperature of the oil in the first reaction area, and can further reduce the yield of dry gas and coke formation and improve the yield of ethylene, liquefied gas, propylene and butylene as can be seen from the table 2.

The preferred embodiments of the present invention have been described in detail with reference to the accompanying drawings, however, the present invention is not limited to the specific details of the above embodiments, and various simple modifications can be made to the technical solution of the present invention within the technical idea of the present invention, and these simple modifications are within the protective scope of the present invention.

It should be noted that the various technical features described in the above embodiments can be combined in any suitable manner without contradiction, and the invention is not described in any way for the possible combinations in order to avoid unnecessary repetition.

In addition, any combination of the various embodiments of the present invention is also possible, and the same should be considered as the content of the present invention as long as it does not depart from the gist of the present invention.

TABLE 1

Raw material numbering	A
		Density (20 ℃ C.)/(kg. m)^-3)	806.1
Freezing point/. degree.C	76
		Element composition/weight%
C	85.18
		H	14.47
O	0.35
		Nitrogen content/(microgram. g.g)^-1)	<1
Distillation range/. degree.C
		IBP	268
10% by volume	391
		30% by volume	465
50% by volume	521

TABLE 2

Temperature difference: the difference between the temperature at the outlet of the second reaction zone and the bottom of the second reaction zone (i.e., at the oxygenate inlet).

Claims

1. A process for the catalytic conversion of a low coke-forming feedstock, the process comprising:

characterized in that the method further comprises: in step (2), the oxygenate is fed through the oxygenate inlet into a second reaction zone of the catalytic conversion reactor to contact and carry out a second catalytic conversion reaction with the first reaction product and the semi-spent catalyst.

2. A process according to claim 1, wherein the temperature at the outlet of the second reaction zone is controlled to be 0.3-100 ℃, preferably 0.3-50 ℃, more preferably 0.5-20 ℃ higher than the temperature at the inlet of the oxygenate.

3. The method of claim 1, wherein the conditions of the first catalytic conversion reaction include: the temperature is 450-;

the conditions of the second catalytic conversion reaction include: the temperature is 460-650 ℃, and the reaction time is 2-30 seconds.

4. The process as claimed in claim 1, wherein the weight of the water generated by the second catalytic conversion reaction of the oxygen-containing compound is W1, and the weight ratio of W1 to the water vapor injected by the first catalytic conversion reaction is 200-500 wt%, preferably 300-400 wt%.

5. The process as claimed in claim 1, wherein the temperature of the oxygenate fed to the catalytic conversion reactor is 20-600 ℃, preferably 200-500 ℃, more preferably 360-450 ℃.

6. The process of claim 1, wherein the low green coke feedstock is a fischer-tropsch synthesis oil.

7. The process according to claim 1, wherein the amount of carbon contained in the semi-spent catalyst is from 0.1 to 3% by weight, preferably from 0.2 to 2% by weight, based on the weight of the semi-spent catalyst.

8. The method according to claim 1, wherein the oxygen-containing compound is one or more selected from the group consisting of alcohols, ethers, and lipids.

9. The method according to claim 1, wherein the oxygen-containing compound is one or more selected from methanol, ethanol, methyl ether, ethyl ether, methyl ethyl ether, dimethyl carbonate and methyl formate.

10. A process according to claim 1, wherein the ratio of the feed of oxygenate to the feed of low coke forming feedstock is in the range of from 0.02 to 50, preferably from 0.05 to 2.

11. The method of claim 1, wherein the catalytic conversion catalyst comprises a molecular sieve, a binder, and a matrix, the molecular sieve being one or more selected from the group consisting of a Y-type molecular sieve, a ZSM-5 type molecular sieve, a SAPO type molecular sieve, an β type molecular sieve, and an SRM type silicoaluminophosphate molecular sieve.

12. The process of claim 1, wherein the catalytic conversion reactor is a riser reactor and/or a fluidized bed reactor, the riser reactor being one or more selected from the group consisting of a constant diameter riser, a variable diameter riser, and a constant linear velocity riser.

13. The process of claim 1, wherein the catalytic conversion reactor is a riser reactor, the first reaction zone is located below the second reaction zone, and the second reaction zone has an inner diameter greater than the inner diameter of the first reaction zone.

14. The method of claim 1, further comprising: at least a portion of the low coke formation feedstock is fed to a second reaction zone for the second catalytic conversion reaction.

15. The method of claim 1, further comprising: and (3) taking out at least part of the semi-regenerated catalyst containing carbon from the regenerator of the step (3), and sending the semi-regenerated catalyst into a catalytic conversion reactor to carry out the first catalytic conversion reaction and/or the second catalytic conversion reaction.