RU2812664C1 - Reactor for control of coke content, as well as device and method for obtaining low-carbon olefins from oxygen-containing compound - Google Patents

Abstract

FIELD: oil and gas industry.

SUBSTANCE: group of inventions relates to a reactor for controlling coke content, a device for producing low-carbon olefins from an oxygen-containing compound and their use, and also relates to the field of chemical devices for catalysis. Reactor (1) contains riser reactor (1-1) and bed reactor (1-2). In this case, the bed reactor contains bed reactor shell (1-3), and the bed reactor shell covers reaction zone I, the transition zone and the gas-solid separation zone I from the bottom up. Moreover, in the inner lower part of reaction zone I, there is bed reactor distributor (1-4), and outside reaction zone I, there is pipe (1-9) supplying a catalyst with controlled coke content. Next, the upper section of the riser reactor passes through the lower part of the bed reactor, which is inserted into the bed reactor in the axial direction, the outlet end of the riser reactor is located in the transition zone, and, in reaction zone I, a plurality of perforated plates (1-5) are installed in series located in the axial direction at the periphery of the riser reactor.

EFFECT: development of a device and method for production of DMTO, in which it is possible to effectively use a new generation DMTO catalyst with high productivity in relation to methanol processing and with controlled coke content with high selectivity of low-carbon olefins.

11 cl, 1 dwg, 5 ex

Description

TECHNICAL FIELD

The present application relates to a reactor for controlling coke content, an apparatus for producing low-carbon olefins from an oxygen-containing compound and their use, and also relates to the field of chemical apparatus for catalysis.

BACKGROUND OF THE ART

Methanol-to-olefin technology (MTO) mainly includes the DMTO (methanol-to-olefin) technology of the Dalian Institute of Chemical Physics of the Chinese Academy of Sciences and the MTO technology of UOP from the United States. In 2010, the construction and commissioning of the Shenhua Baotou methanol to olefin processing plant using DMTO technology was completed. This is the world's first industrial application of MTO technology. As of the end of 2019, 14 commercial DMTO plants were operational with a total production capacity of approximately 8 million tons of low-carbon olefins per year.

In recent years, DMTO technology has been further developed and new generation DMTO catalysts with better performance have gradually been introduced into industry, bringing greater benefits to DMTO plants. New generation DMTO catalyst provides higher methanol throughput and low carbon olefin selectivity. It is difficult for existing industrial DMTO devices to take full advantage of the new generation DMTO catalyst. Therefore, there is a need to develop a DMTO production apparatus and method that can efficiently utilize a new generation DMTO catalyst with high methanol processing capacity and high selectivity for low carbon olefins.

DISCLOSURE OF THE INVENTION

According to a first aspect of the present application, there is provided a coke control reactor configured to control the conversion and formation of coke particles in the catalyst. On the one hand, the inactive large molecular coke particles remaining in the regenerated catalyst are converted into low molecular weight coke particles; and on the other hand, the feedstock of the riser reactor and the feedstock of the bed reactor can also enter the catalyst to form high-activity low molecular weight coke particles, and the low molecular weight coke particles are mainly polymethylbenzene and polymethylnaphthalene, which can improve the selectivity for ethylene.

The coke content control reactor includes a riser reactor and a bed reactor; the bed reactor includes a shell of the bed reactor, and the shell of the bed reactor covers the reaction zone I, the transition zone and the gas-solid separation zone I from bottom to top; in the inner lower part of reaction zone I there is a reactor distributor with a bed; outside reaction zone I there is a pipe for supplying catalyst with a controlled coke content; the upper section of the lift reactor passes through the lower part of the bed reactor, which is inserted into the bed reactor in the axial direction; and in the transition zone the output end of the elevator reactor is located.

Specifically, the coke-controlled catalyst supply pipe is configured to supply the coke-controlled catalyst to a next-level reactor, such as a methanol shift reactor.

If necessary, at least one perforated plate can be installed in reaction zone I; and if there are a plurality of perforated plates, the plurality of perforated plates may be successively arranged in an axial direction at the periphery of the riser reactor; the outlet end of the lift reactor may be located above the perforated plate; and the bed reactor distributor may be located below the perforated plate.

In particular, the perforated plate may be positioned horizontally, and the peripheral portion of the perforated plate is configured to abut the inner wall of the bed reactor so that flow can pass through the holes in the perforated plate.

If necessary, the perforated plate can have a porosity from 1% to 30%.

If necessary, gas-solid separation zone I may comprise a bed reactor gas-solid separation separator and a bed reactor gas collection chamber; a gas outlet of the gas-solid separation separator of the bed reactor is configured to communicate with a gas collection chamber of the bed reactor; a catalyst outlet from the gas-solid separation separator of the bed reactor may be formed above the perforated plate; and the bed reactor gas collection chamber is configured to communicate with a coke control product gas supply pipe located outside the bed reactor.

Specifically, the bed reactor gas-solid separator may be a bed reactor gas-solid separator cyclone.

In particular, the inner upper part of the bed reactor may be provided with a chamber for collecting gas from the bed reactor; and the catalyst outlet from the cyclonic gas-solid separation separator of the bed reactor may be located below the outlet end of the riser reactor.

If necessary, the bed reactor distributor may be configured to supply raw material of the bed reactor; and the bed reactor feedstock may include hydrogen in an amount of from 0 wt%. up to 20% wt., methane in amounts from 0% wt. up to 50% wt., ethane in amounts from 0% wt. up to 50% wt., ethylene in an amount from 0% wt. up to 20% wt., propane in amounts from 0% wt. up to 50% wt., propylene in amounts from 0% wt. up to 20% wt., butane in amounts from 0% wt. up to 90% wt., butene in amounts from 0% wt. up to 90% wt., pentane in an amount from 0% wt. up to 90% wt., pentene in an amount from 0% wt. up to 90% wt., hexane in an amount from 0% wt. up to 90% wt., hexene in amounts from 0% wt. up to 90% wt., methanol in an amount from 0% wt. up to 50% wt., ethanol in an amount from 0% wt. up to 50% wt. and water in an amount of 0% wt. up to 50% by weight, and the total content of methanol, ethanol and water may exceed or be equal to 10% by weight.

If necessary, the elevator reactor can be configured to supply catalyst and raw material of the elevator reactor; and the raw material of the lift reactor may include hydrogen in an amount of from 0% by weight. up to 20% wt., methane in amounts from 0% wt. up to 50% wt., ethane in amounts from 0% wt. up to 50% wt., ethylene in an amount from 0% wt. up to 20% wt., propane in amounts from 0% wt. up to 50% wt., propylene in amounts from 0% wt. up to 20% wt., butane in amounts from 0% wt. up to 90% wt., butene in amounts from 0% wt. up to 90% wt., pentane in an amount from 0% wt. up to 90% wt., pentene in an amount from 0% wt. up to 90% wt., hexane in an amount from 0% wt. up to 90% wt., hexene in amounts from 0% wt. up to 90% wt., methanol in an amount from 0% wt. up to 50% wt., ethanol in an amount from 0% wt. up to 50% wt. and water in an amount of 0% wt. up to 50% by weight, and the total content of methanol, ethanol and water may exceed or be equal to 10% by weight.

A second aspect of the present application also provides an apparatus for producing low carbon olefins from an oxygen-containing compound, including a methanol conversion reactor and a coke control reactor described above.

Low carbon olefins mentioned in this application refer to ethylene and propylene.

Optionally, the methanol conversion reactor may include a methanol conversion reactor shell and a feed pipe; the shell of the methanol conversion reactor may include a lower shell and an upper shell; the lower shell is configured to surround the reaction zone II, and the distributor of the methanol conversion reactor can be located in the inner lower part of the reaction zone II; the feed pipe may be positioned axially above reaction zone II; the feed pipe may have one end closed and the other end in communication with reaction zone II; the upper shell may be located at the periphery of the supply pipe; the upper shell and the wall of the supply pipe may form a cavity; this cavity can be divided from bottom to top into a spent catalyst zone and a gas-solid separation zone II; and the spent catalyst zone may be provided with a gas distributor for the spent catalyst zone.

In particular, the upper casing may be located at the periphery of the supply pipe and configured to wrap around it. The distributor of the methanol conversion reactor may be configured to supply raw material with an oxygen-containing compound; and the gas distributor for the spent catalyst zone may be configured to supply fluidizing gas from the spent catalyst zone.

If necessary, gas-solid separation zone II can be provided with a first gas-solid separation unit of the methanol conversion reactor; an upper portion of the supply pipe may be connected to an inlet of the first gas-solid separation unit of the methanol conversion reactor; a catalyst outlet from the first gas-solid separation unit of the methanol conversion reactor may be formed in a spent catalyst zone; a gas outlet from the first gas-solid separation unit of the methanol conversion reactor is configured to communicate with a gas collection chamber of the methanol conversion reactor; and the gas collecting chamber of the methanol conversion reactor may be connected to a product gas supply pipe.

If necessary, gas-solid separation zone II can also be provided with a second gas-solid separation unit of the methanol conversion reactor; a gas inlet of the second gas-solid separation unit of the methanol conversion reactor may be formed in the gas-solid separation zone II; a catalyst outlet of the second gas-solid separation unit of the methanol conversion reactor may be formed in a spent catalyst zone; and a gas outlet of the second gas-solid separation unit of the methanol conversion reactor is configured to communicate with a gas collection chamber of the methanol conversion reactor.

If necessary, a gas distributor for the spent catalyst zone may be located below the first gas-solid separation unit of the methanol conversion reactor and the second gas-solid separation unit of the methanol conversion reactor; and a methanol conversion reactor cooler may be additionally provided in the spent catalyst zone.

If necessary, the spent catalyst circulation pipe and the spent catalyst inclined pipe can also be installed outside the spent catalyst area; the spent catalyst circulation pipe may be configured to connect the spent catalyst zone and the reaction zone II; and the spent catalyst slant may be configured to discharge the spent catalyst.

In particular, the spent catalyst circulation pipe may be configured to supply a portion of the spent catalyst from the spent catalyst zone to reaction zone II. The spent catalyst circulation pipe may be provided with a sliding valve for circulating the spent catalyst.

If necessary, the gas-solid separation zone II is configured to communicate with the gas collection chamber of the bed reactor through a pipe for supplying a gaseous product to control the coke content; and reaction zone II is configured to communicate with reaction zone I through a pipe for supplying catalyst with a controlled coke content.

In particular, a sliding valve for the coke controlled catalyst may be further provided on the coke controllable catalyst supply pipe.

If necessary, said device may further include a regenerator; this regenerator can be connected to the spent catalyst slant pipe, so that the spent catalyst can be supplied to the regenerator; the regenerator can be connected to the riser reactor so that the regenerated catalyst can be supplied to the reactor to control the coke content; and a regenerator distributor can be installed in the inner lower part of the regenerator.

In particular, the regenerator distributor may be configured to supply regeneration gas.

If necessary, a regenerator stripping column can be additionally installed in the lower part of the regenerator; the upper section of the regenerator stripper may be located within the regenerator, and the inlet of the upper section of the regenerator stripper may be located above the regenerator distributor; the lower section of the regenerator stripping column may be located outside the regenerator, and the outlet of the lower section of the regenerator stripping column may be connected to the riser reactor; and the regenerator stripper column may be additionally provided with a regenerator cooler.

If necessary, the regenerator can be connected to the spent catalyst inclined pipe through the spent catalyst supply pipe and the stripper column of the methanol conversion reactor; and the regenerator may be connected to the inlet of the riser reactor by means of a regenerator stripper column and a regenerated catalyst inclined pipe.

Specifically, the spent catalyst sliding valve may be installed between the spent catalyst supply pipe and the stripper column of the methanol conversion reactor; and the inlet of the spent catalyst sliding valve may be piping connected to the bottom of the stripper column of the methanol conversion reactor, and the outlet of the spent catalyst sliding valve may be piping connected to the inlet of the spent catalyst supply pipe.

A sliding valve for the regenerated catalyst may be installed between the regenerator stripper column and the regenerated catalyst inclined pipe; and the inlet of the regenerated catalyst sliding valve is connected by a pipeline to the bottom of the stripping column of the regenerator, and the outlet of the regenerated catalyst sliding valve is connected by a pipeline to the inlet of the regenerated catalyst slant pipe.

If necessary, the regenerator can be additionally equipped with a block for separating gas and solid matter of the regenerator, as well as a chamber for collecting regenerator gas; a catalyst outlet of the gas-solid separation unit of the regenerator may be formed above the regenerator distributor; a gas outlet of the gas-solid separation unit of the regenerator may be connected to a gas collection chamber of the regenerator; and the gas collection chamber of the regenerator may be connected to a flue gas supply pipe located outside the regenerator.

According to a third aspect of the present application, there is also provided a method for on-line modification of a DMTO catalyst, comprising supplying riser reactor raw material and catalyst to a transition zone from the riser reactor, and supplying bed reactor feedstock to reaction zone I; and allowing the catalyst to contact and react with the riser reactor feedstock and the bed reactor feedstock to form a coke-controlled catalyst and a coke-controlled product gas, wherein the catalyst is a DMTO catalyst; and the coke controlled catalyst is a modified DMTO catalyst.

If desired, the catalyst active component may be a SAPO-34 molecular sieve.

In the present application, the catalyst entering the riser reactor may be fresh catalyst or regenerated catalyst, and preferably regenerated catalyst, so that both coke control and catalyst regeneration can be realized on-line.

If desired, the catalyst may be a regenerated catalyst; and the coke content of the regenerated catalyst may be less than or equal to 3 wt%.

If necessary, the coke content in the catalyst with controlled coke content can range from 4 wt.%. up to 9% wt.

If necessary, the quartile deviation of the distribution of coke content in a catalyst with controlled coke content can be less than 1% wt. Specifically, in the present application, the coke content of the coke-controlled catalyst is controlled in the range of 4 wt.%. up to 9% wt. by using a coke content control reactor and selecting a coke content control process. Since the catalyst is granular, the coke content of the catalyst refers to the average coke content of the catalyst granules, but the actual coke content of different catalyst granules may vary. In the present application, the quartile deviation of the coke content distribution of a coke controlled catalyst can be controlled to be less than 1 wt% to narrow the overall coke content distribution of the catalyst, thereby increasing catalyst activity and selectivity for low carbon olefins.

If desired, the coke particles in the controlled coke catalyst may include polymethylbenzene and polymethylnaphthalene; the total mass of polymethylbenzene and polymethylnaphthalene can be more than 70% wt. from the total mass of coke or may be equal to the specified amount; the mass of coke particles with a molecular weight of more than 184 may be less than 25% wt. from the total mass of coke or may be equal to the specified amount; and the total mass of coke may refer to the total mass of coke particles.

In the present application, the types and contents of coke particles are also very important, and are also one of the control objectives in the present application. In the present application, the total weight of polymethylbenzene and polymethylnaphthalene is controlled to be greater than or equal to 70% by weight. of the total coke mass, using a coke control mechanism and selection of coke control process parameters to increase catalyst activity and improve selectivity for low carbon olefins.

If necessary, the operating conditions for carrying out the process in a lift reactor can be as follows: apparent linear gas velocity: from 3 m/s to 10 m/s; reaction temperature: from 400 °C to 700 °C; reaction pressure: from 100 kPa to 500 kPa; and layer density: from 10 kg/ ^m3 to 150 kg/ ^m3 .

If necessary, the operating conditions for the process in reaction zone I of the bed reactor can be as follows: apparent linear gas velocity: from 0.1 m/s to 1.0 m/s; reaction temperature: from 300 °C to 650 °C; reaction pressure: from 100 kPa to 500 kPa; and layer density: from 150 kg/ ^m3 to 800 kg/ ^m3 .

According to a fourth aspect of the present application, there is also provided a method for producing low carbon olefins from an oxygen-containing compound, including the above-described method of on-line modification of a DMTO catalyst.

If necessary, the method may also include: supplying a gaseous coke control product to a gas-solid separation zone of the methanol conversion reactor; and supplying catalyst with a controlled coke content to reaction zone II of the methanol conversion reactor.

If desired, in reaction zone II, the oxygenate feedstock may contact and react with a controlled coke content catalyst to form stream A with low carbon olefins and spent catalyst.

If necessary, stream A can be divided into gas-phase stream B and solid-phase stream C after performing gas-solid separation in gas-solid separation zone II of the methanol conversion reactor; gas-phase stream B may enter a gas collection chamber of the methanol conversion reactor; solid-phase stream C can enter the spent catalyst zone; and gas phase stream B may include low carbon olefins and solid phase stream C may include spent catalyst.

If necessary, fluidizing gas from the spent catalyst zone can be supplied to the spent catalyst zone; fluidizing gas from the spent catalyst zone and coke control product gas may be mixed, the mixture containing a portion of the spent catalyst, to produce stream D; stream D can be divided into gas-phase stream E and solid-phase stream F after gas-solid separation; the gas-phase stream E may be supplied to a gas collection chamber of the methanol conversion reactor; solid phase flow F can enter the spent catalyst zone; the gas phase stream E may be a gas mixture of fluidizing gas from the spent catalyst zone and a product gas for coke control; and the solid stream F may be spent catalyst.

If necessary, the gas-phase stream B and the gas-phase stream E can be mixed in the gas collection chamber of the methanol shift reactor to produce a gaseous product, and the gaseous product can be supplied to the downstream operating section through a gaseous product supply pipe.

If necessary, part of the spent catalyst from the spent catalyst zone can be returned to the lower part of reaction zone II through the spent catalyst circulation pipe; and the remainder of the spent catalyst can be discharged through the spent catalyst inclined pipe.

If necessary, the spent catalyst discharged through the spent catalyst inclined pipe can be fed to the regenerator; and the regeneration gas may be supplied to the regenerator to interact and react with the spent catalyst to produce a stream G with flue gas and regenerated catalyst.

If necessary, the G stream can be divided into gas and solid; the separated flue gas can be supplied to a gas collection chamber of the regenerator, and then supplied to a downstream flue gas treatment system through a flue gas supply pipe; and the separated regenerated catalyst can be stripped and cooled and then fed to the reactor for coke control.

In particular, the separated regenerated catalyst may be supplied to a riser reactor for stripping and cooling, and then may be fed to a bed reactor from the riser reactor.

If desired, the oxygen-containing compound may include methanol and/or dimethyl ether (DME).

If necessary, the coke content in the spent catalyst can be from 9% wt. up to 13% wt.

If desired, the fluidizing gas from the spent catalyst zone may include nitrogen and/or water vapor.

If necessary, the regeneration gas can include from 0% wt. up to 100% wt. air, from 0% wt. up to 50% wt. oxygen, from 0% wt. up to 50% wt. nitrogen and from 0% wt. up to 50% wt. water vapor; and the values of air, oxygen, nitrogen and water vapor content cannot be equal to zero at the same time.

If necessary, the operating conditions for the process in reaction zone II of the methanol conversion reactor can be as follows: apparent linear gas velocity: from 0.5 m/s to 7.0 m/s; reaction temperature: from 350 °C to 550 °C; reaction pressure: from 100 kPa to 500 kPa; and layer density: from 100 kg/ ^m3 to 500 kg/ ^m3 .

If necessary, the operating conditions for the process in the spent catalyst zone of the methanol conversion reactor can be as follows: apparent linear gas velocity: from 0.1 m/s to 1.0 m/s; reaction temperature: from 350 °C to 550 °C; reaction pressure: from 100 kPa to 500 kPa; and layer density: from 200 kg/ ^m3 to 800 kg/ ^m3 .

If necessary, the operating conditions for the process in the regenerator can be as follows: apparent linear gas velocity: from 0.5 m/s to 2.0 m/s; regeneration temperature: from 600 °C to 750 °C; regeneration pressure: from 100 kPa to 500 kPa; and layer density: from 150 kg/ ^m3 to 700 kg/ ^m3 .

As used herein, the term “coke content” refers to the weight ratio of coke particles to catalyst with a controlled coke content.

The subscripts in each C ₄ -C ₆ represent the number of carbon atoms in the corresponding group, for example, C ₄ -C ₆ hydrocarbon compounds are hydrocarbon compounds with 4-6 carbon atoms.

In this application, when expressing the specific production rate, the mass of dimethyl ether in the oxygen-containing compound is equivalently converted to the mass of methanol based on the mass of element C, and the unit of specific production rate is ton of methanol per ton of low-carbon olefins.

In the method according to this application, the specific production consumption can be from 2.50 to 2.60 tons of methanol per ton of low-carbon olefins.

A lift reactor is similar to a plug flow reactor (PFR). Thus, by using a lift reactor to treat the catalyst to control the coke content, a narrow coke content distribution can be achieved. The residence time of the catalyst in the riser reactor is short, typically from 1 to 20 seconds. Therefore, it is difficult to significantly increase the coke content of the regenerated catalyst by treating the catalyst only using a riser reactor. The coke control reactor described herein includes a riser reactor and a bed reactor. On the one hand, due to the advantages of the lift reactor, it is possible to obtain a catalyst with a narrow distribution of coke content; and, on the other hand, a bed reactor is used to further increase the coke content of the catalyst and improve selectivity for low carbon olefins. The main feature of the bed reactor described herein is that the perforated plate is configured to prevent backmixing of the catalyst between the beds and improve the uniformity of coke distribution in the catalyst. The catalyst first enters the upper layer of the reactor with a layer from the riser reactor, gradually flows down into the lower layer and then from the lower layer enters the reaction zone II of the methanol conversion reactor.

Possible beneficial effects of this application:

(1) The coke control reactor described herein is configured to control the conversion and formation of coke particles in the catalyst. On the one hand, the inactive large molecular coke particles remaining in the regenerated catalyst are converted into low molecular weight coke particles; and on the other hand, the feedstock of the riser reactor and the feedstock of the bed reactor can also enter the catalyst to form high-activity low molecular weight coke particles, and the low molecular weight coke particles are mainly polymethylbenzene and polymethylnaphthalene, which can improve the selectivity for ethylene.

(2) By using the on-line modification method of DMTO catalyst by carrying out the coke control reaction described in this application, a coke controllable catalyst with a high coke content, a narrow coke content distribution, and polymethylbenzene and polymethylnaphthalene as the main components can be obtained coke, i.e., the regenerated catalyst with low low carbon olefin selectivity is converted to a controlled coke catalyst with high low carbon olefin selectivity.

(3) The regenerated catalyst described in this application can be directly used in the process of producing low carbon olefins from an oxygenate compound without treatment to control the coke content, in which case the selectivity for low carbon olefins in the resulting gaseous product is between 80% wt. up to 83% wt. When the regenerated catalyst described herein is subjected to coke control treatment and then used in a process for producing low carbon olefins from an oxygenate, the selectivity for low carbon olefins in the resulting product gas is 92 wt%. up to 96% wt.

(4) The methanol conversion reactor described in the present application uses a composite fluidized bed reactor including a fast fluidized bed zone and a bubbling fluidized bed zone. The fast fluidized bed zone is a reaction zone that can produce a high flow of methanol, increase the methanol processing capacity per unit volume of the device, and achieve a weight hourly space velocity (WHSV) of methanol ranging from 5 h ^-1 to 20 h ^-1 . The bubbling fluidized bed zone is a spent catalyst zone that is configured to reduce the temperature of the spent catalyst, supply low-temperature spent catalyst to the reaction zone, increase the density of the bed in the reaction zone, and control the temperature of the bed in the reaction zone. When the apparent linear gas velocity is from 0.5 m/s to 7.0 m/s, the corresponding layer density is from 500 kg/m ³ to 100 kg/m ³ .

(5) The methanol shift reactor described in the present application has a structure in which the first gas-solid separation unit of the methanol shift reactor is directly connected to the feed pipe, which enables rapid separation of the low carbon olefin containing gas and the spent catalyst in the stream A, and prevents low carbon olefins from further reacting under the influence of the spent catalyst to form high molecular weight hydrocarbon by-products.

BRIEF DESCRIPTION OF THE DRAWINGS

In fig. 1 is a schematic block diagram of a DMTO apparatus for producing low-carbon olefins from an oxygen-containing compound according to an embodiment of the present application.

List of item numbers:

1 is a reactor for controlling coke content; 1-1 is a lift reactor; 1-2 represents a bed reactor; 1-3 represents a reactor shell with a layer; 1-4 represents a bed reactor distributor; 1-5 represents a perforated plate; 1-6 is a cyclone separator for separating gas and solids of a bed reactor; 1-7 is a gas collection chamber of a bed reactor; 1-8 is a product gas supply pipe for controlling coke content; 1-9 represents a pipe for supplying a catalyst with a controlled coke content; 1-10 is a sliding valve for a controlled coke catalyst;

2 is a methanol conversion reactor; 2-1 is a methanol conversion reactor shell; 2-2 is a distributor of a methanol conversion reactor; 2-3 represents a supply pipe; 2-4 represents a first gas-solid separation unit of the methanol conversion reactor; 2-5 represents a gas collection chamber of a methanol conversion reactor; 2-6 represents a gas distributor for the spent catalyst zone; 2-7 represents a methanol conversion reactor cooler; 2-8 represents a second gas-solid separation unit of the methanol conversion reactor; 2-9 represents a pipe for supplying a gaseous product; 2-10 represents a spent catalyst circulation pipe; 2-11 is a sliding valve for circulating spent catalyst; 2-12 is an inclined pipe for spent catalyst; 2-13 represents a stripping column of a methanol conversion reactor; 2-14 is a sliding valve for spent catalyst; 2-15 represents a spent catalyst supply pipe; 3 is a regenerator;

3-1 represents the regenerator shell; 3-2 is a regenerator distributor; 3-3 is a gas-solid separation unit of the regenerator; 3-4 represents a chamber for collecting regenerator gas; 3-5 represents a flue gas supply pipe; 3-6 represents a regenerator stripping column; 3-7 represents a regenerator cooler; 3-8 represents a sliding valve for regenerated catalyst; 3-9 is a slant pipe for regenerated catalyst.

IMPLEMENTATION OF THE INVENTION

The present application will be described in detail below with reference to examples, but the present application is not limited to these examples.

The main characteristic of the dimethyl ether/methanol to olefins (DMTO) catalyst is that the selectivity of the low carbon olefin to methanol increases with the coke content of the catalyst. Low carbon olefins mentioned in this application refer to ethylene and propylene.

Applicants have discovered through research that the main factors affecting DMTO catalyst activity and low carbon olefin selectivity are coke content, coke content distribution, and coke particles in the catalyst. Given the same average coke content in the catalysts, the narrower the distribution of coke content, the higher the selectivity and activity of low-carbon olefins. The coke particles in the catalyst may include polymethylaromatic hydrocarbons, polymethylcycloalkanes, etc., and polymethylbenzene and polymethylnaphthalene can contribute to the formation of ethylene. Thus, controlling the coke content, coke content distribution, and coke particles in the catalyst is key to controlling the activity of the DMTO catalyst and improving low carbon olefin selectivity.

To improve the performance of a DMTO catalyst, the present application proposes a method for on-line modification of a DMTO catalyst by performing a coke control reaction, comprising the following steps:

a) the regenerated catalyst is fed to the reactor to control the coke content;

b) coke control raw material including hydrogen, methane, ethane, ethylene, propane, propylene, butane, butene, pentane, pentene, hexane, hexene, methanol, ethanol and water are supplied to the coke control reactor;

c) the coke control raw material contacts and reacts with the regenerated catalyst in the coke control reactor such that the coke control raw material cokes the regenerated catalyst, wherein the coked catalyst is called a coke control catalyst; The coke content in the catalyst with controlled coke content is from 4% wt. up to 9% wt., and the quartile deviation of the distribution of coke content is less than 1% wt.; coke particles include polymethylbenzene and polymethylnaphthalene, and the total mass of polymethylbenzene and polymethylnaphthalene is more than 70% by weight. from the total mass of coke or equal to the specified amount; and the mass of coke particles with a molecular weight of more than 184 is less than 25% wt. from the total mass of coke or equal to the specified amount; And

d) catalyst with controlled coke content is fed to the methanol conversion reactor.

The regenerated catalyst may be a DMTO catalyst with a coke content of less than or equal to 3 wt.%, and the active component of the DMTO catalyst may be a SAPO-34 molecular sieve.

The raw material for coke control can consist of hydrogen in amounts ranging from 0% wt. up to 20% wt., methane in amounts from 0% wt. up to 50% wt., ethane in amounts from 0% wt. up to 50% wt., ethylene in amounts from 0% wt. up to 20% wt., propane in amounts from 0% wt. up to 50% wt., propylene in amounts from 0% wt. up to 20% wt., butane in amounts from 0% wt. up to 90% wt., butene in amounts from 0% wt. up to 90% wt., pentane in an amount from 0% wt. up to 90% wt., pentene in an amount from 0% wt. up to 90% wt., hexane in an amount from 0% wt. up to 90% wt., hexene in amounts from 0% wt. up to 90% wt., methanol in an amount from 0% wt. up to 50% wt., ethanol in an amount from 0% wt. up to 50% wt. and water in an amount of 0% wt. up to 50% by weight, and the total content of methanol, ethanol and water may exceed or be equal to 10% by weight.

The coke control reaction temperature can range from 300°C to 700°C.

In the present application, both the raw material for the bed reactor and the raw material for the riser reactor are raw materials for controlling coke content.

According to one aspect of the present application, there is provided a method for producing low carbon olefins from an oxygen-containing compound, which includes a method for on-line modification of a DMTO catalyst by carrying out the coke control reaction described above, and an apparatus used therefor. The device includes a reactor (1) for controlling coke content, a methanol conversion reactor (2) and a regenerator (3).

A reactor (1) for controlling coke content is proposed, configured to modify the DMTO catalyst on-line, wherein the reactor (1) for controlling coke content includes a riser reactor (1-1) and a bed reactor (1-2); The bed reactor (1-2) includes a bed reactor shell (1-3), a bed reactor distributor (1-4), a perforated plate (1-5), a cyclone separator (1-6) for gas separation, and bed reactor solids, bed reactor gas collection chamber (1-7), product gas supply pipe (1-8) for coke control, coke control catalyst supply pipe (1-9) and sliding valve (1-10) for a catalyst with controlled coke content;

the bed reactor (1-2) is divided from bottom to top into a reaction zone, a transition zone and a gas-solid separation zone; the lift reactor (1-1) passes through the shell (1-3) of the bed reactor, is partially located below the bed reactor (1-2) and is partially located in the bed reactor (1-2); the distributor (1-4) of the reactor with the layer is located in the lower part of the reaction zone and n perforated plates (1-5) are installed in the reaction zone, where 0 ≤ n ≤ 9;

The cyclone separator (1-6) for separating gas and solid matter of the bed reactor is located in the gas-solid separation zone, the entrance to the cyclone separator (1-6) for separating gas and solid matter of the bed reactor is formed in the gas-solid separation zone substances, the catalyst outlet of the gas-solid separation cyclone separator (1-6) of the bed reactor is formed in the reaction zone, and the gas outlet of the cyclone separator (1-6) for gas-solid separation of the bed reactor is connected to the chamber ( 1-7) to collect bed reactor gas;

a chamber (1-7) for collecting bed reactor gas is located in the upper part of the bed reactor (1-2); the inlet of the coke content control product gas supply pipe (1-8) is connected to the gas collection chamber (1-7) of the bed reactor, and the coke content control product gas supply pipe (1-8) outlet is connected to the upper part methanol conversion reactor (2);

a sliding valve (1-10) for the coke-controlled catalyst is located in the pipe (1-9) for supplying the coke-controlled catalyst; the inlet of the pipe (1-9) for supplying the catalyst with a controlled content of coke is connected to the lower part of the reaction zone, and the outlet of the pipe (1-9) for supplying the catalyst with a controlled content of coke is connected to the lower part of the methanol conversion reactor (2).

In a preferred embodiment, the perforated plate may have a porosity from 1% to 30%.

Methanol conversion reactor (2) for producing low-carbon olefins by methanol conversion includes a methanol conversion reactor shell (2-1), a methanol conversion reactor distributor (2-2), a supply pipe (2-3), a first block (2- 4) for separating gas and solid matter of the methanol conversion reactor, a chamber (2-5) for collecting gas from the methanol conversion reactor, a gas distributor (2-6) of the spent catalyst zone, a cooler (2-7) of the methanol conversion reactor, a second unit (2- 8) for gas-solid separation of the methanol conversion reactor, pipe (2-9) for supplying the gaseous product, pipe (2-10) for circulating the spent catalyst, sliding valve (2-11) for circulating the spent catalyst, inclined pipe (2 -12) for spent catalyst, stripper column (2-13) of the methanol conversion reactor, sliding valve (2-14) for spent catalyst and pipe (2-15) for supplying spent catalyst;

the lower part of the methanol conversion reactor (2) is a reaction zone, its middle part is a spent catalyst zone, and its upper part is a gas-solid separation zone; the distributor (2-2) of the methanol conversion reactor is located in the lower part of the reaction zone of the methanol conversion reactor (2); the supply pipe (2-3) is located in the central zones of the middle and upper parts of the methanol conversion reactor (2); the lower end of the supply pipe (2-3) is connected to the upper end of the reaction zone, and the upper part of the supply pipe (2-3) is connected to the inlet of the first gas-solid separation unit (2-4) of the methanol conversion reactor; the first gas-solid separation unit (2-4) of the methanol conversion reactor is located in the gas-solid separation zone of the methanol conversion reactor (2); and the gas outlet of the first unit (2-4) for separating gas and solid of the methanol shift reactor is connected to the gas collection chamber (2-5) of the methanol shift reactor, and the catalyst outlet of the first unit (2-4) for separating the gas and methanol conversion reactor solids formed in the spent catalyst zone;

the gas distributor (2-6) of the spent catalyst zone is located in the lower part of the spent catalyst zone, the cooler (2-7) of the methanol conversion reactor is located in the spent catalyst zone, and the second block (2-8) for separating gas and solid matter of the methanol conversion reactor is located in the gas-solid separation zone of the methanol conversion reactor (2); the inlet of the second block (2-8) for separating gas and solid matter of the methanol conversion reactor is formed in the gas-solid separation zone of the methanol conversion reactor (2), the gas outlet of the second block (2-8) for separating gas and solid matter of the conversion reactor methanol is connected to a gas collection chamber (2-5) of the methanol conversion reactor, and a catalyst outlet of the second unit (2-8) for separating gas and solids of the methanol conversion reactor is formed in the spent catalyst zone; a methanol shift reactor gas collection chamber (2-5) is located at the top of the methanol shift reactor (2), and a product gas supply pipe (2-9) is connected to the top of the methanol shift reactor gas collection chamber (2-5) ;

the inlet of the spent catalyst circulation pipe (2-10) is connected to the spent catalyst zone, and the outlet of the spent catalyst circulation pipe (2-10) is connected to the lower part of the reaction zone of the methanol conversion reactor (2); a sliding valve (2-11) for circulating the spent catalyst is located in the pipe (2-10) for circulating the spent catalyst; the outlet of the controlled coke catalyst supply pipe (1-9) is connected to the bottom of the reaction zone of the methanol conversion reactor (2), the inlet of the spent catalyst slant pipe (2-12) is connected to the spent catalyst zone, and the outlet of the slant pipe (2 -12) for spent catalyst is connected to the upper part of the stripping column (2-13) of the methanol conversion reactor; the stripping column (2-13) of the methanol conversion reactor is located outside the shell (2-1) of the methanol conversion reactor; and the inlet of the sliding valve (2-14) for the spent catalyst is connected by pipeline to the bottom of the stripper column (2-13) of the methanol conversion reactor, the outlet of the sliding valve (2-14) for the spent catalyst is connected by pipeline to the inlet of the pipe (2-15) for supply of spent catalyst, and the outlet of the pipe (2-15) for supplying spent catalyst is connected to the middle part of the regenerator (3).

In a preferred embodiment, the first gas-solid separation unit (2-4) of the methanol conversion reactor may have one or more sets of gas-solid separation cyclone separators, and each set of gas-solid separation cyclone separators may include a cyclone a first-stage gas-solid separator and a second-stage cyclone gas-solid separator.

In a preferred embodiment, the second gas-solid separation unit (2-8) of the methanol conversion reactor may have one or more sets of gas-solid separation cyclone separators, and each set of gas-solid separation cyclone separators may include a cyclone a first-stage gas-solid separator and a second-stage cyclone gas-solid separator.

The methanol conversion reactor (2) may be a fluidized bed reactor.

A regenerator (3) for regenerating a catalyst is proposed, which includes a shell (3-1) of the regenerator, a distributor (3-2) of the regenerator, a block (3-3) for separating gas and solid matter of the regenerator, a chamber (3-4) for collecting regenerator gas, flue gas supply pipe (3-5), regenerator stripper column (3-6), regenerator cooler (3-7), sliding valve (3-8) for regenerated catalyst and inclined pipe (3-9) for regenerated catalyst;

the regenerator distributor (3-2) is located at the bottom of the regenerator (3), and the regenerator gas-solid separation unit (3-3) is located at the top of the regenerator (3); The inlet of the block (3-3) for separating gas and solid matter of the regenerator is formed in the upper part of the regenerator (3), the gas outlet of the block (3-3) for separating gas and solid matter of the regenerator is connected to the chamber (3-4) for collecting gas regenerator, and the outlet for the catalyst block (3-3) for separating gas and solid matter of the regenerator is formed in the lower part of the regenerator (3); a regenerator gas collection chamber (3-4) is located at the top of the regenerator (3), and a flue gas supply pipe (3-5) is connected to the top of the regenerator gas collection chamber (3-4);

the stripping column (3-6) of the regenerator is located outside the shell (3-1) of the regenerator, and the inlet pipe of the stripping column (3-6) of the regenerator passes through the shell (3-1) of the regenerator and is open above the distributor (3-2) of the regenerator; the regenerator cooler (3-7) is located in the regenerator stripping column (3-6); and the inlet of the sliding valve (3-8) for the regenerated catalyst is connected by pipeline to the bottom of the stripping column (3-6) of the regenerator, the outlet of the sliding valve (3-8) for the regenerated catalyst is connected by pipeline to the inlet of the inclined pipe (3-9) for the regenerated catalyst, and the outlet of the inclined pipe (3-9) for the regenerated catalyst is connected to the lower part of the riser reactor (1-1).

In a preferred embodiment, the gas-solid separation unit (3-3) of the regenerator may have one or more sets of gas-solid separation cyclone separators, and each gas-solid separation cyclone separator set may include a gas-solid separation cyclone separator. gas and solids of the first stage and a cyclone separator for separating gas and solids of the second stage.

The regenerator (3) may be a fluidized bed reactor.

In the present application, the inner diameter of the transition zone of the coke control reactor is gradually increased from bottom to top; in the methanol conversion reactor, the inner diameter of the connection between reaction zone II and the feed pipe gradually decreases from bottom to top; and the inner diameter of the connection between reaction zone II and the spent catalyst zone gradually increases from bottom to top.

Where appropriate, the inlet of a pipe for supplying catalyst with a controlled coke content may be formed above the distributor of the bed reactor; the outlet of a pipe for supplying catalyst with a controlled coke content may be formed above the distributor of the methanol conversion reactor; The inlet of the spent catalyst circulation pipe may be formed above the gas distributor of the spent catalyst zone, and the outlet of the spent catalyst circulation pipe may be formed above the distributor of the methanol conversion reactor; and the spent catalyst slant may be positioned above the spent catalyst zone gas distributor.

According to yet another aspect of the present application, there is also provided a method of converting methanol to olefin (Methanol-to-olefin technology, MTO), including a method of modifying a DMTO catalyst on-line by carrying out a coke control reaction, comprising the following steps:

a. from the bottom of the riser reactor (1-1) the riser reactor raw material is supplied, which contains the regenerated catalyst from the inclined pipe (3-9) for the regenerated catalyst and which moves upward into the transition zone of the bed reactor (1-2); the raw material of the bed reactor is supplied from the distributor (1-4) of the bed reactor to the reaction zone of the bed reactor (1-2); In the riser reactor (1-1) and bed reactor (1-2), the regenerated catalyst comes into contact with the feedstock of the riser reactor and the feedstock of the bed reactor, and a chemical reaction occurs to form a catalyst with a controlled content of coke and gaseous coke control product; The gaseous coke control product containing the coke control catalyst portion enters the gas-solid separation cyclone separator (1-6) of the gas-solid separation bed reactor, and then the gaseous coke control product enters a gas-solid separation zone of the methanol conversion reactor (2) through a chamber (1-7) for collecting bed reactor gas and a pipe (1-8) for supplying product gas to control the coke content; and the coke-controlled catalyst enters the reaction zone of the methanol conversion reactor (2) through the coke-controlled catalyst supply pipe (1-9) and through the coke-controlled catalyst sliding valve (1-10);

b. the oxygen-containing compound feedstock is supplied from the distributor of the methanol conversion reactor (2-2) to the reaction zone of the methanol conversion reactor (2), where it contacts a catalyst with a controlled coke content to create a stream A with low carbon olefins and spent catalyst; stream A enters the first gas-solid separation unit (2-4) of the methanol conversion reactor through a gas-solid separation feed pipe (2-3) to produce gas-phase stream B and solid-phase stream C, wherein gas-phase stream B is a gas with low carbon olefins, and the solid phase stream C is a spent catalyst; gas-phase stream B enters the gas collection chamber (2-5) of the methanol conversion reactor, and solid-phase stream C enters the spent catalyst zone; The fluidizing gas from the spent catalyst zone is supplied from the gas distributor (2-6) of the spent catalyst zone to the spent catalyst zone, where it comes into contact with the spent catalyst, and the fluidizing gas from the spent catalyst zone and the coke control product gas are mixed, wherein the mixture will contain a portion of the spent catalyst to produce stream D; stream D enters the second gas-solid separation unit (2-8) of the methanol conversion reactor to perform gas-solid separation to produce a gas-phase stream E and a solid-phase stream F, wherein the gas-phase stream E is a mixture of gases from the fluidizing gas from spent catalyst and product gas zones to control the coke content, and the solid phase stream F represents the spent catalyst; gas-phase stream E enters the gas collection chamber (2-5) of the methanol conversion reactor, and solid-phase stream F enters the spent catalyst zone; gas-phase stream B and gas-phase stream E are mixed in the gas-collecting chamber (2-5) of the methanol conversion reactor to obtain a gaseous product, and this gaseous product is supplied to the downstream working section through a pipe (2-9) for supplying the gaseous product; part of the spent catalyst located in the spent catalyst zone is returned to the lower part of the reaction zone of the methanol conversion reactor (2) through a pipe (2-10) for circulating the spent catalyst and through a sliding valve (2-11) for circulating the spent catalyst, and the remaining part spent catalyst enters the stripper column (2-13) of the methanol conversion reactor through the spent catalyst slant pipe (2-12) for stripping, and then enters the middle part of the regenerator (3) through the sliding valve (2-14) for spent catalyst and through a pipe (2-15) for supplying spent catalyst; And

c. regeneration gas is supplied from the distributor (3-2) of the regenerator to the lower part of the regenerator (3), and in the regenerator the regeneration gas comes into contact with the spent catalyst and a chemical reaction occurs, while part of the coke in the spent catalyst is burned and removed, obtaining a stream G with flue gas and regenerated catalyst; stream G enters the gas-solid separation unit (3-3) of a regenerator for separating gas and solids to produce flue gas and regenerated catalyst; the flue gas enters the regenerator gas collecting chamber (3-4), and then enters the downstream flue gas treatment system through the flue gas supply pipe (3-5); the regenerated catalyst is returned to the lower part of the regenerator (3); and the regenerated catalyst in the regenerator (3) enters the stripper column (3-6) of the regenerator for stripping and cooling, and then enters the reactor (1) to control the coke content through the sliding valve (3-8) for the regenerated catalyst and through the inclined pipe (3-9) for the regenerated catalyst.

In a preferred embodiment, the method according to the present application can be implemented using the above-mentioned apparatus including a coke control reactor (1), a methanol conversion reactor (2) and a regenerator (3).

In a preferred embodiment, the raw material of the lift reactor according to this method can consist of hydrogen in an amount of 0% by weight. up to 20% wt., methane in amounts from 0% wt. up to 50% wt., ethane in amounts from 0% wt. up to 50% wt., ethylene in amounts from 0% wt. up to 20% wt., propane in amounts from 0% wt. up to 50% wt., propylene in amounts from 0% wt. up to 20% wt., butane in amounts from 0% wt. up to 90% wt., butene in amounts from 0% wt. up to 90% wt., pentane in an amount from 0% wt. up to 90% wt., pentene in an amount from 0% wt. up to 90% wt., hexane in an amount from 0% wt. up to 90% wt., hexene in amounts from 0% wt. up to 90% wt., methanol in an amount from 0% wt. up to 50% wt., ethanol in an amount from 0% wt. up to 50% wt. and water in an amount of 0% wt. up to 50% by weight, and the total content of methanol, ethanol and water may exceed or be equal to 10% by weight.

In a preferred embodiment, the raw material of the bed reactor according to this method can consist of hydrogen in an amount of from 0% by weight. up to 20% wt., methane in amounts from 0% wt. up to 50% wt., ethane in amounts from 0% wt. up to 50% wt., ethylene in amounts from 0% wt. up to 20% wt., propane in amounts from 0% wt. up to 50% wt., propylene in amounts from 0% wt. up to 20% wt., butane in amounts from 0% wt. up to 90% wt., butene in amounts from 0% wt. up to 90% wt., pentane in an amount from 0% wt. up to 90% wt., pentene in an amount from 0% wt. up to 90% wt., hexane in an amount from 0% wt. up to 90% wt., hexene in amounts from 0% wt. up to 90% wt., methanol in an amount from 0% wt. up to 50% wt., ethanol in an amount from 0% wt. up to 50% wt. and water in an amount of 0% wt. up to 50% by weight, and the total content of methanol, ethanol and water may exceed or be equal to 10% by weight.

In a preferred embodiment, the oxygen-containing compound according to this method may be one of the group consisting of methanol, dimethyl ether, and mixtures of methanol and dimethyl ether.

In a preferred embodiment, the waste fluidizing gas from the spent catalyst zone of this process may be one of the group consisting of nitrogen, steam, and a mixture of nitrogen and steam.

In a preferred embodiment, the regeneration gas according to this method may contain from 0% wt. up to 100% wt. air, from 0% wt. up to 50% wt. oxygen, from 0% wt. up to 50% wt. nitrogen and from 0% wt. up to 50% wt. water vapor.

In a preferred embodiment, the active catalyst component may be a SAPO-34 molecular sieve.

In a preferred embodiment, the coke content of the regenerated catalyst may be less than or equal to 3 wt%.

In a preferred embodiment, the coke content of the coke-controlled catalyst can range from 4 wt.%. up to 9% wt., and the quartile deviation of the distribution of coke content can be less than 1% wt.; The coke particles may include polymethylbenzene and polymethylnaphthalene, and the total mass of polymethylbenzene and polymethylnaphthalene may be more than 70% by weight. from the total mass of coke or may be equal to the specified amount; and the mass of coke particles with a molecular weight greater than 184 may be less than 25% by weight. of the total mass of coke or may be equal to the specified amount.

In a preferred embodiment, the coke content of the spent catalyst can range from 9% wt. up to 13% wt. and more preferably, the coke content in the spent catalyst can range from 10% wt. up to 12% wt..

In a preferred embodiment, the operating conditions for carrying out the process in the riser reactor (1-1) can be as follows: apparent linear gas velocity: from 3 m/s to 10 m/s; reaction temperature: from 400 °C to 700 °C; reaction pressure: from 100 kPa to 500 kPa; and layer density: from 10 kg/ ^m3 to 150 kg/ ^m3 .

In a preferred embodiment, the operating conditions for carrying out the process in the reaction zone of the bed reactor (1-2) can be as follows: apparent linear gas velocity: from 0.1 m/s to 1.0 m/s; reaction temperature: from 300 °C to 650 °C; reaction pressure: from 100 kPa to 500 kPa; and layer density: from 150 kg/ ^m3 to 800 kg/ ^m3 .

In a preferred embodiment, the operating conditions for the process in the reaction zone of the methanol conversion reactor (2) may be as follows: apparent linear gas velocity: from 0.5 m/s to 7.0 m/s; reaction temperature: from 350 °C to 550 °C; reaction pressure: from 100 kPa to 500 kPa; and layer density: from 100 kg/ ^m3 to 500 kg/ ^m3 .

In a preferred embodiment, the operating conditions for the process in the spent catalyst zone of the methanol conversion reactor (2) may be as follows: apparent linear gas velocity: from 0.1 m/s to 1.0 m/s; reaction temperature: from 350 °C to 550 °C; reaction pressure: from 100 kPa to 500 kPa; and layer density: from 200 kg/ ^m3 to 800 kg/ ^m3 .

In a preferred embodiment, the operating conditions for the process in the regenerator (3) can be as follows: apparent linear gas velocity: from 0.5 m/s to 2.0 m/s; regeneration temperature: from 600 °C to 750 °C; regeneration pressure: from 100 kPa to 500 kPa; and layer density: from 150 kg/ ^m3 to 700 kg/ ^m3 .

According to the method according to this application, the gaseous product can consist of ethylene in an amount of 40% by weight. up to 55% wt., propylene in an amount from 37% wt. up to 53% wt., hydrocarbon compounds C ₄ -C ₆ in an amount of 4% wt. or less and other components in an amount of 4% wt. or less; in this case, other components can be methane, ethane, propane, hydrogen, CO, CO ₂ , etc., and the overall selectivity of ethylene and propylene in the gaseous product can be from 92% wt. up to 96% wt.

Example 1

In this example, the device shown in FIG. 1, in which there is no perforated plate in the bed reactor.

In this example, the elevator reactor feedstock is a mixture of 6 wt.% butane, 81% wt. butene, 2% wt. methanol and 11% wt. water; The raw material of the bed reactor is a mixture of 6% wt. butane, 81% wt. butene, 2% wt. methanol and 11% wt. water; the oxygen-containing compound is methanol; the fluidizing gas in the spent catalyst zone is nitrogen; the regeneration gas is air; the active component of the catalyst is a SAPO-34 molecular sieve; the coke content in the regenerated catalyst is approximately 3 wt.%; The coke content of the coke controlled catalyst is approximately 6 wt.%, with the total weight of polymethylbenzene and polymethylnaphthalene being approximately 82 wt.%. of the total mass of coke, the mass of coke particles with a molecular weight of more than 184 is approximately 5% wt. of the total mass of coke, and the quartile deviation of the distribution of coke content in a catalyst with a controlled coke content is approximately 0.9% wt.; The coke content in the spent catalyst is approximately 12% by weight; the operating conditions for the process in the lift reactor (1-1) are as follows: apparent linear gas velocity: 10.0 m/s; reaction temperature: 700 °C; reaction pressure: 100 kPa; and layer density: 10 kg/m ³ ; The process operating conditions in the reaction zone of the bed reactor (1-2) are as follows: apparent linear gas velocity: approximately 0.3 m/s, reaction temperature: approximately 650 °C, reaction pressure: approximately 100 kPa and bed density: approximately 600 kg/m ³ ; The process operating conditions in the reaction zone of the methanol conversion reactor (2) are as follows: apparent linear gas velocity: approximately 7.0 m/s, reaction temperature: approximately 550 °C, reaction pressure: approximately 100 kPa and bed density: approximately 100 kg /m ³ ; The process operating conditions in the spent catalyst zone of the methanol conversion reactor (2) are as follows: apparent linear gas velocity: approximately 1.0 m/s, reaction temperature: approximately 550 °C, reaction pressure: approximately 100 kPa and bed density: approximately 200 kg/m ³ ; and the operating conditions for the process in the regenerator (3) are as follows: apparent linear gas velocity: approximately 0.5 m/s, regeneration temperature: approximately 750 °C, regeneration pressure: approximately 100 kPa and bed density: approximately 700 kg/m ³ .

In this example, the average hourly oxygenate feed rate to the methanol conversion reactor is approximately 20 h ^-1 ; the gaseous product consists of ethylene in an amount of 55% wt., propylene in an amount of 37% wt., hydrocarbon compounds C ₄ -C ₆ in an amount of 4% wt. and other components in an amount of 4% by weight, while other components include methane, ethane, propane, hydrogen, CO, CO ₂ , etc.; and the specific product consumption is 2.60 tons of methanol per ton of low-carbon olefins.

Example 2

In this example, the device shown in FIG. 1, in which 4 perforated plates having a porosity of 30% are installed in a bed reactor.

In this example, the riser reactor feedstock is a mixture of 22% wt. methane, 24% wt. ethane, 3% wt. ethylene, 28% wt. propane, 4% wt. propylene, 7% wt. hydrogen and 12% wt. water; The raw material of the bed reactor is a mixture of 22% wt. methane, 24% wt. ethane, 3% wt. ethylene, 28% wt. propane, 4% wt. propylene, 7% wt. hydrogen and 12% wt. water; the oxygen-containing compound is a mixture of 82% wt. methanol and 18% wt. dimethyl ether; the fluidizing gas in the spent catalyst zone is water vapor; regeneration gas is a mixture of 50% wt. air and 50% wt. water vapor; the active component of the catalyst is a SAPO-34 molecular sieve; the coke content in the regenerated catalyst is approximately 1% by weight; The coke content of the coke controlled catalyst is approximately 4 wt.%, with the total weight of polymethylbenzene and polymethylnaphthalene being approximately 76 wt.%. of the total mass of coke, the mass of coke particles with a molecular weight of more than 184 is approximately 15% wt. of the total mass of coke, and the quartile deviation of the distribution of coke content in a catalyst with a controlled coke content is approximately 0.5 wt.%. The coke content in the spent catalyst is approximately 9% by weight; the operating conditions for the process in the lift reactor (1-1) are as follows: apparent linear gas velocity: 3.0 m/s; reaction temperature: 400 °C; reaction pressure: 500 kPa; and layer density: 150 kg/m ³ ; The process operating conditions in the reaction zone of the bed reactor (1-2) are as follows: apparent linear gas velocity: approximately 0.1 m/s, reaction temperature: approximately 300 °C, reaction pressure: approximately 500 kPa and bed density: approximately 800 kg/m ³ ; The process operating conditions in the reaction zone of the methanol conversion reactor (2) are as follows: apparent linear gas velocity: approximately 0.5 m/s, reaction temperature: approximately 350 °C, reaction pressure: approximately 500 kPa and bed density: approximately 500 kg /m ³ ; The process operating conditions in the spent catalyst zone of the methanol conversion reactor (2) are as follows: apparent linear gas velocity: approximately 0.1 m/s, reaction temperature: approximately 350 °C, reaction pressure: approximately 500 kPa and bed density: approximately 800 kg/m ³ ; and the operating conditions for the process in the regenerator (3) are as follows: apparent linear gas velocity: approximately 2.0 m/s, regeneration temperature: approximately 600 °C, regeneration pressure: approximately 500 kPa and bed density: approximately 150 kg/m ³ .

In this example, the average hourly oxygenate feed rate to the methanol conversion reactor is approximately 5 h ^-1 ; the gaseous product consists of ethylene in an amount of 40% wt., propylene in an amount of 53% wt., hydrocarbon compounds C ₄ -C ₆ in an amount of 4% wt. and other components in an amount of 3% by weight, while other components include methane, ethane, propane, hydrogen, CO, CO ₂ , etc.; and the specific product consumption is 2.58 tons of methanol per ton of low-carbon olefins.

Example 3

In this example, the device shown in FIG. 1, in which 4 perforated plates having a porosity of 1% are installed in a bed reactor.

In this example, the raw material of the riser reactor is a mixture of 1% wt. propane, 1% wt. propylene, 3% wt. butane, 51% wt. butene, 3% wt. pentane, 22% wt. pentene, 1% wt. hexane, 7% wt. hexene, 2% wt. methanol and 9% wt. water; The raw material of the bed reactor is a mixture of 1% wt. propane, 1% wt. propylene, 3% wt. butane, 51% wt. butene, 3% wt. pentane, 22% wt. pentene, 1% wt. hexane, 7% wt. hexene, 2% wt. methanol and 9% wt. water; the oxygen-containing compound consists of dimethyl ether; The fluidizing gas in the spent catalyst zone is a mixture of 5% wt. nitrogen and 95% wt. water vapor; regeneration gas is a mixture of 50% wt. air and 50% wt. oxygen; the active component of the catalyst is a SAPO-34 molecular sieve; the coke content in the regenerated catalyst is approximately 2% by weight; The coke content of the coke controlled catalyst is approximately 5 wt.%, with the total weight of polymethylbenzene and polymethylnaphthalene being approximately 79 wt.%. of the total mass of coke, the mass of coke particles with a molecular weight of more than 184 is approximately 10% wt. of the total mass of coke, and the quartile deviation of the distribution of coke content in a catalyst with a controlled coke content is approximately 0.5 wt.%. the coke content in the spent catalyst is approximately 11% by weight; the operating conditions for the process in the lift reactor (1-1) are as follows: apparent linear gas velocity: 6.0 m/s; reaction temperature: 600 °C; reaction pressure: 300 kPa; and layer density: 80 kg/m ³ ; The process operating conditions in the reaction zone of the bed reactor (1-2) are as follows: apparent linear gas velocity: approximately 0.4 m/s, reaction temperature: approximately 550 °C, reaction pressure: approximately 300 kPa and bed density: approximately 500 kg/ ^m3 ; The process operating conditions in the reaction zone of the methanol conversion reactor (2) are as follows: apparent linear gas velocity: approximately 3.0 m/s, reaction temperature: approximately 450 °C, reaction pressure: approximately 300 kPa and bed density: approximately 230 kg /m ³ ; The operating conditions for the process in the spent catalyst zone of the methanol conversion reactor (2) are as follows: apparent linear gas velocity: approximately 0.2 m/s, reaction temperature: approximately 450 °C, reaction pressure: approximately 300 kPa and bed density: approximately 600 kg/m ³ ; and the operating conditions for the process in the regenerator (3) are as follows: apparent linear gas velocity: approximately 1.0 m/s, regeneration temperature: approximately 680 °C, regeneration pressure: approximately 300 kPa and bed density: approximately 360 kg/m ³ .

In this example, the average hourly oxygenate feed rate to the methanol conversion reactor is approximately 15 h ^-1 ; the gaseous product consists of ethylene in an amount of 50% wt., propylene in an amount of 45% wt., hydrocarbon compounds C ₄ -C ₆ in an amount of 3% wt. and other components in an amount of 2% by weight, while other components include methane, ethane, propane, hydrogen, CO, CO ₂ , etc.; and the specific product consumption is 2.53 tons of methanol per ton of low-carbon olefins.

Example 4

In this example, the device shown in FIG. 1, in which 9 perforated plates having a porosity of 5% are installed in a bed reactor.

In this example, the riser reactor feedstock is a mixture of 5 wt.% butane, 72% wt. butene, 8% wt. methanol and 15% wt. water; The raw material of the bed reactor is a mixture of 5% wt. butane, 72% wt. butene, 8% wt. methanol and 15% wt. water; the oxygen-containing compound is methanol; The fluidizing gas in the spent catalyst zone is a mixture of 73% wt. nitrogen and 27% wt. water vapor; regeneration gas is a mixture of 50% wt. air and 50% wt. nitrogen; the active component of the catalyst is a SAPO-34 molecular sieve; the coke content in the regenerated catalyst is approximately 2% by weight; The coke content of the coke controlled catalyst is approximately 9 wt.%, with the total weight of polymethylbenzene and polymethylnaphthalene being approximately 71 wt.%. of the total mass of coke, the mass of coke particles with a molecular weight of more than 184 is approximately 25% wt. of the total mass of coke, and the quartile deviation of the distribution of coke content in a catalyst with a controlled coke content is approximately 0.2% by weight; The coke content in the spent catalyst is approximately 13% by weight; the operating conditions for the process in the elevator reactor (1-1) are as follows: apparent linear gas velocity: 4.0 m/s; reaction temperature: 550 °C; reaction pressure: 200 kPa; and layer density: 120 kg/m ³ ; The process operating conditions in the reaction zone of the bed reactor (1-2) are as follows: apparent linear gas velocity: approximately 1.0 m/s, reaction temperature: approximately 500 °C, reaction pressure: approximately 200 kPa and bed density: approximately 150 kg/ ^m3 ; The process operating conditions in the reaction zone of the methanol conversion reactor (2) are as follows: apparent linear gas velocity: approximately 4.0 m/s, reaction temperature: approximately 500 °C, reaction pressure: approximately 200 kPa and bed density: approximately 160 kg /m ³ ; The process operating conditions in the spent catalyst zone of the methanol conversion reactor (2) are as follows: apparent linear gas velocity: approximately 0.5 m/s, reaction temperature: approximately 500 °C, reaction pressure: approximately 200 kPa and bed density: approximately 300 kg/m ³ ; and the operating conditions for the process in the regenerator (3) are as follows: apparent linear gas velocity: approximately 1.5 m/s, regeneration temperature: approximately 700 °C, regeneration pressure: approximately 200 kPa and bed density: approximately 280 kg/m ³ .

In this example, the average hourly oxygenate feed rate to the methanol conversion reactor is approximately 11 h ^-1 ; the gaseous product consists of ethylene in an amount of 52% wt., propylene in an amount of 44% wt., hydrocarbon compounds C ₄ -C ₆ in an amount of 2% wt. and other components in an amount of 2% by weight, while other components include methane, ethane, propane, hydrogen, CO, CO ₂ , etc.; and the specific product consumption is 2.50 tons of methanol per ton of low-carbon olefins.

Example 5

This example is a comparative example and differs from example 4 in that the coke control reaction is not used to modify the DMTO catalyst on-line; and the raw material supplied to the riser reactor and bed reactor is nitrogen, which is an inert gas and does not change the properties of the regenerated catalyst in the lift reactor and bed reactor, i.e., the catalyst entering the reaction zone II of the conversion reactor methanol, is a regenerated catalyst.

In this example, the gaseous product consists of ethylene in an amount of 46% wt., propylene in an amount of 37% wt., C ₄ -C ₆ hydrocarbon compounds in an amount of 12% wt. and other components in an amount of 5% by weight, while other components include methane, ethane, propane, hydrogen, CO, CO ₂ , etc.; and the specific product consumption is 2.90 tons of methanol per ton of low-carbon olefins.

This comparative example shows that on-line modification of a DMTO catalyst through a coke control reaction can significantly improve catalyst performance and reduce specific product consumption.

The above examples are just a few examples of the present application and are not intended to limit it in any way. Although the present application has been described above with preferred examples, the present application is not limited to them. Certain changes or modifications that may be suggested by any technical person skilled in the art using the technical content disclosed above without departing from the scope of the technical solutions of this application are equivalent to equivalent implementation cases and are included in the specified scope of technical solutions.

Claims (98)

1. Reactor (1) for controlling coke content, which contains a riser reactor (1-1) and a bed reactor (1-2); wherein the bed reactor contains a shell (1-3) of the bed reactor, and the shell of the bed reactor covers the reaction zone I, the transition zone and the gas-solid separation zone I from bottom to top; in the inner lower part of the reaction zone I there is a distributor (1-4) of the reactor with a bed; outside reaction zone I there is a pipe (1-9) for supplying catalyst with a controlled coke content; the upper section of the lift reactor passes through the lower part of the bed reactor, which is inserted into the bed reactor in the axial direction; the outlet end of the elevator reactor is located in the transition zone, and in reaction zone I, a plurality of perforated plates (1-5) are installed, sequentially located in the axial direction at the periphery of the elevator reactor. 2. The reactor for controlling coke content according to claim 1, in which the outlet end of the lift reactor is located above the plurality of perforated plates; a bed reactor distributor is located below the plurality of perforated plates; Preferably, the plurality of perforated plates have a porosity of from 1% to 30%. 3. The reactor for controlling coke content according to claim 1, in which the gas-solid separation zone I is equipped with a separator for separating gas and solid matter of the bed reactor and a chamber (1-7) for collecting gas of the bed reactor; a gas separator outlet communicates with the bed reactor gas collection chamber for separating gas and solids of the bed reactor; above the plurality of perforated plates is an outlet for a catalyst separator for separating gas and solids of the bed reactor; a gas collection chamber of the bed reactor is in communication with a product gas supply pipe (1-8) for coke control located outside the bed reactor; preferably, the bed reactor distributor is configured to supply raw material of the bed reactor; And the raw material of the bed reactor contains hydrogen in an amount of 0% wt. up to 20% wt., methane in an amount from 0% wt. up to 50% wt., ethane in amounts from 0% wt. up to 50% wt., ethylene in an amount from 0% wt. up to 20% wt., propane in amounts from 0% wt. up to 50% wt., propylene in amounts from 0% wt. up to 20% wt., butane in amounts from 0% wt. up to 90% wt., butene in an amount from 0% wt. up to 90% wt., pentane in an amount from 0% wt. up to 90% wt., pentene in an amount from 0% wt. up to 90% wt., hexane in an amount from 0% wt. up to 90% wt., hexene in an amount from 0% wt. up to 90% wt., methanol in an amount from 0% wt. up to 50% wt., ethanol in an amount from 0% wt. up to 50% wt. and water in an amount of 0% wt. up to 50% by weight, and the total content of methanol, ethanol and water is greater than or equal to 10% by weight; preferably, the elevator reactor is configured to supply catalyst and raw material of the elevator reactor; And The raw material of the elevator reactor contains hydrogen in an amount of 0% wt. up to 20% wt., methane in an amount from 0% wt. up to 50% wt., ethane in amounts from 0% wt. up to 50% wt., ethylene in an amount from 0% wt. up to 20% wt., propane in amounts from 0% wt. up to 50% wt., propylene in amounts from 0% wt. up to 20% wt., butane in amounts from 0% wt. up to 90% wt., butene in an amount from 0% wt. up to 90% wt., pentane in an amount from 0% wt. up to 90% wt., pentene in an amount from 0% wt. up to 90% wt., hexane in an amount from 0% wt. up to 90% wt., hexene in an amount from 0% wt. up to 90% wt., methanol in an amount from 0% wt. up to 50% wt., ethanol in an amount from 0% wt. up to 50% wt. and water in an amount of 0% wt. up to 50 wt.%, and the total content of methanol, ethanol and water is greater than or equal to 10 wt.%. 4. A device for producing low-carbon olefins from an oxygen-containing compound, the device comprising a methanol conversion reactor (2) and a reactor for controlling coke content according to claim 1. 5. The device according to claim 4, in which the methanol conversion reactor contains a methanol conversion reactor shell (2-1) and a supply pipe (2-3); the shell of the methanol conversion reactor includes a lower shell and an upper shell; the lower shell encloses reaction zone II, and the distributor (2-2) of the methanol conversion reactor is located in the inner lower part of reaction zone II; the feed pipe is located above reaction zone II; the feed pipe has one closed end and the other end in communication with reaction zone II; the upper shell is located on the periphery of the supply pipe; the upper shell and the wall of the supply pipe form a cavity; said cavity is divided from bottom to top into a spent catalyst zone and a gas-solid separation zone II; the spent catalyst zone is equipped with a gas distributor (2-6) for the spent catalyst zone; preferably, gas-solid separation zone II is provided with a first gas-solid separation unit (2-4) of the methanol conversion reactor; an upper portion of the supply pipe is connected to an inlet of the first gas-solid separation unit of the methanol conversion reactor; the catalyst outlet of the first gas-solid separation unit of the methanol conversion reactor is located in the spent catalyst zone; a gas outlet of the first gas-solid separation unit of the methanol conversion reactor is in communication with a gas collection chamber (2-5) of the methanol conversion reactor; A a gas collection chamber of the methanol conversion reactor is connected to a pipe (2-9) for supplying a gaseous product; preferably, gas-solid separation zone II is also provided with a second gas-solid separation unit (2-8) of the methanol conversion reactor; in gas-solid separation zone II, a gas inlet of a second gas-solid separation unit of the methanol conversion reactor is located; in the spent catalyst zone, a catalyst outlet of the second gas-solid separation unit of the methanol conversion reactor is located; And a gas outlet of a second unit for separating gas and solids of the methanol conversion reactor is connected to the gas collection chamber of the methanol conversion reactor; preferably, the spent catalyst zone gas distributor is located below the first gas-solid separation unit of the methanol conversion reactor and the second gas-solid separation unit of the methanol conversion reactor; A in the spent catalyst zone, a cooler (2-7) of the methanol conversion reactor is additionally provided; preferably, a pipe (2-10) for circulating the spent catalyst and an inclined pipe (2-12) for the spent catalyst are also installed outside the spent catalyst zone; the spent catalyst circulation pipe is configured to connect the spent catalyst zone and reaction zone II; A the inclined pipe for the spent catalyst is configured to discharge the spent catalyst; preferably, the gas-solid separation zone II communicates with the gas collection chamber of the bed reactor via a product gas supply pipe to control the coke content; A reaction zone II communicates with reaction zone I through a pipe for supplying catalyst with a controlled coke content. 6. The device according to claim 4, additionally containing a regenerator (3); and the regenerator is connected to the spent catalyst inclined pipe, and the spent catalyst is supplied to the regenerator; the regenerator is connected to the lift reactor and the regenerated catalyst is supplied to the reactor to control the coke content; a regenerator distributor (3-2) is installed in the inner lower part of the regenerator; preferably, a stripping column (3-6) of the regenerator is additionally installed in the lower part of the regenerator; located inside the regenerator is the upper section of the regenerator stripping column, and above the regenerator distributor is the inlet of the upper section of the regenerator stripping column; The lower section of the regenerator stripping column is located outside the regenerator, and the outlet of the lower section of the regenerator stripping column is connected to the elevator reactor; And the regenerator stripping column is additionally equipped with a regenerator cooler (3-7); preferably, the regenerator is connected to the spent catalyst inclined pipe through a spent catalyst supply pipe (2-15) and a stripper column (2-13) of the methanol conversion reactor; A the regenerator is connected to the inlet of the riser reactor by means of a stripping column of the regenerator and an inclined pipe (3-9) for the regenerated catalyst; preferably, the regenerator is additionally equipped with a block (3-3) for separating gas and solid matter of the regenerator and a chamber (3-4) for collecting regenerator gas; above the regenerator distributor is an outlet for a catalyst unit for separating gas and solids of the regenerator; connected to the gas collection chamber of the regenerator is the gas outlet of the unit for separating gas and solid matter of the regenerator; A the regenerator gas collection chamber is connected to a flue gas supply pipe (3-5) located outside the regenerator. 7. A method for producing low-carbon olefins from an oxygen-containing compound, including a method for modifying a catalyst used in the processing of dimethyl ether/methanol to olefins (DMTO) on-line using a coke control reactor according to claim 1, including supplying the raw material of the riser reactor and the catalyst to the transition zone from the riser reactor and supplying the raw material of the bed reactor to the reaction zone I; And allowing the catalyst to contact and react with the riser reactor feedstock and the bed reactor feedstock to form a coke controllable catalyst and a coke controllable product gas, wherein the catalyst is a DMTO catalyst; and the coke controlled catalyst is a modified DMTO catalyst. 8. The method according to claim 7, according to which the active component of the catalyst is a molecular sieve SAPO-34; preferably the catalyst is a regenerated catalyst; and the coke content of the regenerated catalyst is less than or equal to 3% by weight; Preferably, the coke content of the coke-controlled catalyst is between 4 wt%. up to 9% by weight, and the quartile deviation of the distribution of coke content in the catalyst with controlled coke content is less than 1% by weight; preferably, the coke particles in the coke controlled catalyst contain polymethylbenzene and polymethylnaphthalene; the total mass of polymethylbenzene and polymethylnaphthalene is more than 70% wt. from the total mass of coke or may be equal to the specified amount; the mass of coke particles with a molecular weight of more than 184 is less than 25% wt. from the total mass of coke or may be equal to the specified amount; and the total mass of coke refers to the total mass of coke particles; preferably, the operating conditions for carrying out the process in the riser reactor are as follows: the apparent linear gas velocity is from 3 m/s to 10 m/s; the reaction temperature is from 400°C to 700°C; the reaction pressure is from 100 kPa to 500 kPa; and the layer density ranges from 10 kg/m ³ to 150 kg/m ³ ; Preferably, the operating conditions for the process in the reaction zone I of the bed reactor are as follows: the apparent linear gas velocity is from 0.1 m/s to 1.0 m/s; the reaction temperature is from 300°C to 650°C; the reaction pressure is from 100 kPa to 500 kPa; and the density of the layer is from 150 kg/m ³ to 800 kg/m ³ . 9. The method according to claim 7, further comprising: supplying a gaseous product generated by controlling the coke content into the gas-solid separation zone of the methanol conversion reactor; preferably, the method further includes: supplying catalyst with a controlled coke content to reaction zone II of the methanol conversion reactor; preferably, in reaction zone II, the oxygen-containing feedstock contacts and reacts with a controlled coke content catalyst to form stream A with low carbon olefins and spent catalyst; preferably, stream A is separated into gas-phase stream B and solid-phase stream C after performing gas-solid separation in gas-solid separation zone II of the methanol conversion reactor; gas-phase stream B enters the gas collection chamber of the methanol conversion reactor; solid-phase flow C enters the spent catalyst zone; And gas-phase stream B contains low-carbon olefins, and solid-phase stream C contains spent catalyst; preferably, fluidizing gas from the spent catalyst zone is supplied to the spent catalyst zone; fluidizing gas from the spent catalyst zone and coke control product gas are mixed, the mixture containing a portion of the spent catalyst, to produce stream D; stream D is divided into gas-phase stream E and solid-phase stream F after gas-solid separation; gas-phase stream E enters the gas collection chamber of the methanol conversion reactor; solid-phase flow F enters the spent catalyst zone; gas phase stream E is a gas mixture of fluidizing gas from the spent catalyst zone and product gas for coke control; A solid phase stream F represents spent catalyst; Preferably, gas-phase stream B and gas-phase stream E are mixed in a gas collection chamber of the methanol shift reactor to produce a product gas, and the product gas is supplied to the downstream operating section through a product gas supply pipe. 10. The method according to claim 9, according to which part of the spent catalyst from the spent catalyst zone is returned to the lower part of the reaction zone II through a pipe for circulating the spent catalyst; the remainder of the spent catalyst is discharged through an inclined waste catalyst pipe; preferably, the spent catalyst discharged through the spent catalyst inclined pipe is supplied to the regenerator; A the regeneration gas is supplied to the regenerator to interact and react with the spent catalyst to produce a stream G with flue gas and regenerated catalyst; preferably, stream G is separated into gas and solid; the separated flue gas enters the gas collecting chamber of the regenerator, and then enters the downstream flue gas treatment system through the flue gas supply pipe; A the separated regenerated catalyst is stripped and cooled, and then fed to the reactor to control the coke content; preferably, the oxygen-containing compound contains methanol and/or dimethyl ether (DME); preferably, the coke content in the spent catalyst is from 9% wt. up to 13% wt.; preferably, the fluidizing gas from the spent catalyst zone contains nitrogen and/or water vapor; preferably, the regeneration gas contains from 0% wt. up to 100% wt. air, from 0% wt. up to 50% wt. oxygen, from 0% wt. up to 50% wt. nitrogen and from 0% wt. up to 50% wt. water vapor; And the values of the content of air, oxygen, nitrogen and water vapor are not equal to zero at the same time; Preferably, the operating conditions for the process in the reaction zone II of the methanol conversion reactor are as follows: the apparent linear gas velocity is from 0.5 m/s to 7.0 m/s; the reaction temperature is from 350°C to 550°C; the reaction pressure is from 100 kPa to 500 kPa; and the layer density ranges from 100 kg/m ³ to 500 kg/m ³ ; Preferably, the operating conditions for the process in the spent catalyst zone of the methanol conversion reactor are as follows: the apparent linear gas velocity is from 0.1 m/s to 1.0 m/s; the reaction temperature is from 350°C to 550°C; the reaction pressure is from 100 kPa to 500 kPa; and the layer density ranges from 200 kg/m ³ to 800 kg/m ³ ; preferably, the operating conditions for the process in the regenerator are as follows: the apparent linear gas velocity is from 0.5 m/s to 2.0 m/s; the reaction temperature is from 600°C to 750°C; the reaction pressure is from 100 kPa to 500 kPa; and the density of the layer is from 150 kg/m ³ to 700 kg/m ³ .