CN116004277A - Method for simultaneously preparing hydrogen and low-carbon olefin by using hydrocarbon raw material - Google Patents

Abstract

A method for preparing hydrogen and low-carbon olefin simultaneously by using hydrocarbon raw materials comprises the steps of feeding the hydrocarbon raw materials into a fluidization reactor, enabling the hydrocarbon raw materials to contact with a catalyst containing zeolite and metal oxide to generate hydrocarbon bond cleavage reaction to generate hydrogen and carbon bond cleavage reaction to generate low-carbon olefin, and separating reaction oil gas and a catalyst with carbon obtained by the reaction; further separating the reaction oil and gas into products comprising hydrogen, carbon monoxide, carbon dioxide, lower olefins and other hydrocarbons; the obtained catalyst with carbon is sent to a regenerator for carbon burning reaction, and the catalyst with carbon is sent to a regenerator as regenerated catalyst after being burnt and regeneratedThe original device, the regenerated catalyst after reduction enters the fluidization reactor for recycling after steam stripping; the regenerated flue gas enters a separation unit to be separated to obtain carbon monoxide and carbon dioxide. The method of the invention converts the low-value petroleum hydrocarbon raw material into hydrogen and low-carbon olefin, realizes the high-efficiency utilization of limited petroleum resources, and simultaneously realizes CO ₂ Enrichment is helpful for carbon recycling.

Description

A method for simultaneously producing hydrogen and light olefins by using hydrocarbon raw materials

technical field

The invention relates to a method for producing hydrogen and light olefins by using hydrocarbon raw materials as raw materials.

Background technique

Hydrogen is hailed as the core of the future world energy architecture, and is also considered to be the cleanest fuel. However, if the hydrogen comes from fossil fuels, its production process is not "clean". At present, more than 96% of commercial hydrogen is produced from fossil fuels, and a large amount of carbon dioxide will be emitted during the hydrogen production process. This type of hydrogen is also called "grey hydrogen". Water resources, including seawater, are the largest "hydrogen mines" on the earth, and hydrogen production by electrolysis of water is considered to be an effective method for producing hydrogen. However, whether renewable energy electrolysis of water to produce green hydrogen can actually be solved and applied to large-scale carbon reduction, three major problems need to be overcome: large-scale electrolysis of water, low energy consumption and high stability.

Using fossil energy as a resource to realize the production of "blue hydrogen" is an effective technical means to solve hydrogen energy. The relatively mature technology routes for the existing fossil energy hydrogen production methods include reforming hydrogen production using fossil energy such as coal and natural gas, and pyrolysis reforming hydrogen production of chemical raw materials represented by the hydrogen production technology of alcohol cracking. At present, hydrogen production from natural gas reforming and high-temperature cracking in China are mainly used in large-scale hydrogen production industries. The raw material gas in the process of hydrogen production from natural gas is also fuel gas and does not need to be transported, but the investment in hydrogen production from natural gas is relatively high, which is suitable for large-scale industrial production. Coal gasification hydrogen production is the first choice for industrial large-scale hydrogen production, and it is also the mainstream method of fossil energy hydrogen production in my country. The hydrogen production process converts coal into synthesis gas (CO, CH ₄ , H ₂ , CO ₂ , N _{2 ,} etc.) Raw materials for various products such as liquid fuels and natural gas, widely used in petrochemical, steel and other fields. The coal-to-hydrogen technology route is mature and efficient, and can be produced stably on a large scale. However, the power consumption of coal-to-hydrogen fuel is higher than that of natural gas to produce hydrogen, and the requirements for system steam and electricity are high, and enterprises need supporting boilers. In addition, environmental protection issues are prominent. Existing urban refineries have strict environmental requirements, and there are many constraints on coal transportation, which also limit the application of this technology in modern refineries.

The national standard for oil quality requirements has been raised (reduced sulfur content), and the market demand for light oil products has increased. These factors have led to the wider application of hydrogenation processes, which have been the main drivers of the surge in refinery hydrogen demand. According to statistics, the annual growth rate of global refinery hydrogen demand exceeds 4%. The importance of hydrogen production in the refining process has gradually become prominent. A 10Mt/a refinery equipped with a residual oil hydrogenation unit consumes about 1% of the crude oil processed, while a refinery without residual oil hydrogenation consumes about 0.7%, which comes from the hydrogen inside the refinery The supply will be difficult to meet the future demand for hydrogen growth, so more flexible and feasible hydrogen supply strategies need to be explored. The current oil refining capacity is overcapacity, and the transformation of oil refining enterprises from "fuel-based" to "chemical-based" is already the general trend. If flexible hydrogen production technology can be developed and low-carbon olefins are produced, it will undoubtedly have good economic and social benefits.

Contents of the invention

The purpose of this invention is to provide a kind of method that adopts hydrocarbon feedstock to produce hydrogen and light olefins simultaneously.

The method for producing hydrogen and light olefins simultaneously by using hydrocarbon raw materials provided by the invention comprises the following steps:

Send the hydrocarbon raw material into the fluidized reactor, and contact with the catalyst containing zeolite and metal oxide to generate carbon-hydrogen bond breaking reaction to generate hydrogen and carbon-carbon bond breaking reaction to generate low-carbon olefins, and obtain hydrogen and low-carbon olefins Reaction of oil gas and carbon catalyst;

Separation of the reaction oil gas containing hydrogen and light olefins from the carbon catalyst;

The separated reaction oil gas is further separated into products including hydrogen, carbon monoxide, carbon dioxide, light olefins and other hydrocarbons;

The obtained charcoal catalyst is sent to the regenerator for charcoal burning reaction. After the charcoal catalyst is regenerated by burning, it is sent to the reducer as a regenerated catalyst, and the regenerated flue gas enters the separation unit to separate and obtain carbon monoxide and carbon dioxide;

The regenerated catalyst enters the reducer and contacts with the reducing agent to undergo a reduction reaction. The reduced regenerated catalyst enters the fluidized reactor after being stripped.

The hydrocarbon raw material is selected from one or more mixtures of petroleum hydrocarbons, mineral oil and synthetic oil, and the petroleum hydrocarbons are gaseous hydrocarbons, gasoline, diesel oil, vacuum wax oil, atmospheric residue, Vacuum wax oil is mixed with part of vacuum residue or hydrocarbon oil obtained from secondary processing; mineral oil is selected from one or more mixtures of coal liquefied oil, oil sand oil and shale oil; synthetic oil is coal , Natural gas or bitumen distillate obtained by Fischer-Tropsch synthesis. The hydrocarbon oil obtained by secondary processing is selected from one or more of coker gasoline, catalytic diesel oil, hydrogenated diesel oil, coker diesel oil, coker wax oil, deasphalted oil, and furfural refined raffinate oil.

Based on the dry basis weight of the catalyst, the catalyst comprises 5%-65% of natural minerals, 10%-60% of oxides, 10%-60% of zeolites and 0.1%-30% of metal active groups point. The zeolites include medium-pore zeolites and optional large-pore zeolites, the medium-pore zeolites are ZSM series zeolites and/or ZRP zeolites, and the large-pore zeolites are selected from rare earth Y, rare earth hydrogen Y, ultrastable Y and high One or more of silicon Y. The mesoporous zeolite accounts for 5-100% by weight of the total weight of the zeolite, preferably 20-50% by weight. The content of the metal active component is 0.1%-30% by weight, preferably 0.5-20% by weight. The metal active component is selected from one or more compounds of transition metal elements, preferably one or more of nickel, cobalt, iron, tungsten, molybdenum, manganese, copper, zirconium and chromium.

The reaction temperature of the fluidized reactor is 450-800°C, preferably 550-700°C, the reaction time is 0.1-10 seconds, preferably 1-8 seconds, the weight ratio of catalyst to hydrocarbon feedstock is 5-100, preferably 20-50, steam The weight ratio to hydrocarbon feedstock is 0.1-20, preferably 1-10.

The fluidized reactor is selected from one or more combinations of riser reactor, fast bed and dense phase fluidized bed. The fluidized reactor comprises a pre-lift section and at least one fluidized reactor in the reaction zone from bottom to top, preferably 2-8, more preferably 2-3 in the reaction zone.

The concentration of oxygen in the oxygen-containing gas at the bottom of the regenerator is 22-100% by volume, preferably 25-80% by volume.

The regeneration operating conditions are: temperature is 550-700°C, preferably 600-650°C; gas superficial linear velocity is 0.2-1.2 m/s, preferably 0.4-0.8 m/s, and the average residence time of the deactivated catalyst is 1-10 minutes The clock is preferably 2-6 minutes.

The volume ratio of CO/CO ₂ in the regenerated flue gas is 0.2-2.0, preferably 0.8-1.5.

The operating conditions of the reducer are as follows: the temperature is 550-700°C; the apparent linear velocity of the gas is 0.5-3 m/s.

The reducing agent is selected from one or more of small molecular alkanes, preferably one or more of methane, ethane, propane, n-butane and isobutane.

The separated gas obtained from the regeneration flue gas and separation unit is rich in carbon monoxide, which can be used as a raw material for water gas shift to further produce hydrogen and carbon dioxide; it can also be sent to a carbon monoxide boiler to recover the waste heat of flue gas to produce high-quality steam.

The present invention uses hydrocarbon raw materials as raw materials, which not only reduces the cost of raw materials for hydrogen production from natural gas, but also helps alleviate the tense supply situation in my country's natural gas market, which has strategic significance for the stable development of my country's energy structure.

The present invention uses a fluidized reactor to produce hydrogen, and the catalyst circulates between the reactor and the regenerator, which not only realizes the regeneration of the deactivated catalyst, but also transfers a large amount of heat for the reaction, which greatly reduces the cost of the hydrogen production process. Energy consumption is required, process economy is achieved.

The present invention preferably adopts the low-temperature incomplete regeneration technology, and the CO/ _CO2 ratio in the regenerated flue gas is high, which can provide cheap raw material gas for the water-gas shift process and realize the optimal utilization of resources; waste heat can also be recovered from the flue gas to produce Hot water or high-quality steam is supplied to other devices, which can act as a utility island, realize rational utilization of energy in the refinery, and improve process economy.

The oxygen-containing gas used for the regeneration of the deactivated catalyst in the hydrogen production process of the present invention is preferably an oxygen-rich gas, which greatly increases the concentration of _CO in the flue gas, and can realize large-scale production of _CO , and then through capture, utilization, and storage The technology reduces carbon emissions and enables the production of blue hydrogen.

The invention reduces the catalyst before the hydrogen production reaction, reduces the high-valence state metal oxide to the low-valence state metal oxide, improves the dehydrogenation activity of the catalyst, and improves the hydrogen selectivity.

In the present invention, a catalytic conversion method is adopted to convert hydrocarbon raw materials into hydrogen and by-produce light olefins, and the yield of hydrogen is high. The invention not only realizes the high-value utilization of hydrocarbon raw materials and meets the market demand for low-carbon olefins, but also enriches the _CO2 produced in the process, which not only brings hydrogen energy, but also facilitates carbon capture. It can bring greater economic and social benefits to the petrochemical industry.

Description of drawings

Fig. 1 is the process flow diagram of the specific embodiment of the simultaneous production of hydrogen and light olefins provided by the present invention by using hydrocarbon feedstock.

Detailed ways

A method for simultaneously producing hydrogen and light olefins by using a hydrocarbon feedstock, the method comprising the following steps:

Send the hydrocarbon raw material into the fluidized reactor, and contact with the catalyst containing zeolite and metal oxide to generate carbon-hydrogen bond breaking reaction to generate hydrogen and carbon-carbon bond breaking reaction to generate low-carbon olefins, and obtain hydrogen and low-carbon olefins Reaction of oil gas and carbon catalyst;

Separation of the reaction oil gas containing hydrogen and light olefins from the carbon catalyst;

The separated reaction oil gas is further separated into products including hydrogen, carbon monoxide, carbon dioxide, light olefins and other hydrocarbons;

The obtained charcoal catalyst is sent to the regenerator for charcoal burning reaction. After the charcoal catalyst is regenerated by burning, it is sent to the reducer as a regenerated catalyst, and the regenerated flue gas enters the separation unit to separate and obtain carbon monoxide and carbon dioxide;

The regenerated catalyst enters the reducer and contacts with the reducing agent to undergo a reduction reaction. The reduced regenerated catalyst enters the fluidized reactor after being stripped.

The carbon monoxide obtained from the regeneration flue gas and separation unit can be used as the raw material for water gas shift to further produce hydrogen and carbon dioxide; it can also be sent to the carbon monoxide boiler to recover the waste heat of the flue gas to generate high-pressure steam.

The hydrocarbon raw materials are various animal and vegetable oils rich in hydrocarbons, and the hydrocarbon raw materials are selected from one or more mixtures of petroleum hydrocarbons, mineral oils and synthetic oils. Petroleum hydrocarbons are well known to those skilled in the art, for example, they can be petroleum distillates obtained from a primary processing unit, and can be selected from one or more of gasoline, diesel oil, vacuum wax oil, atmospheric residue, vacuum residue Several mixed oils, or distillates obtained from secondary processing units, can be selected from one or more of coker gasoline, catalytic diesel oil, hydrogenated diesel oil, coker diesel oil, coker wax oil, deasphalted oil, and furfural refined raffinate oil. kind. Mineral oil is selected from one or more mixtures of coal liquefied oil, oil sands oil and shale oil. Synthetic oil is distillate oil obtained by Fischer-Tropsch synthesis of coal, natural gas or asphalt.

The catalyst comprises the following components in weight percent:

A) 5%-65% natural minerals,

B) 10%-60% oxide,

C) 10%-60% zeolite, and

D) 0.1%-30% metal active components.

The method provided by the invention can be carried out in various existing fluidized reactors, and said fluidized reactor is selected from one or more combinations of turbulent bed, fast bed and dilute phase transport bed. The fluidized reactor includes a pre-lifting section and at least one reaction zone fluidized reactor from bottom to top. In order to allow the raw material oil to fully react, and according to different target product quality requirements, the reaction zone can be 2-8, preferably 2-3.

The conditions of the catalytic decomposition reaction include: the reaction temperature of the fluidized reactor is 450-800°C, preferably 550-700°C, the reaction time is 0.1-10 seconds, preferably 1-8 seconds, the weight ratio of the catalyst to the hydrocarbon raw material 5-100, preferably 20-50; the weight ratio of steam and hydrocarbon feedstock is 0.1-20, preferably 1-10.

According to the method provided by the present invention, the deactivated catalyst is generally separated from the reaction oil and gas to obtain the deactivated catalyst and reaction oil and gas, and then the obtained reaction oil and gas are separated into hydrogen, CO ₂ , CO, gaseous hydrocarbons and liquid hydrocarbons through a subsequent separation unit. Fractions, gaseous hydrocarbons are further separated by gas separation equipment to obtain hydrocarbon components such as ethylene and propylene, and the methods for separating hydrogen, ethylene and propylene from the reaction product are similar to the conventional technical methods in the art, and the present invention is not limited to this. Not described in detail.

In the method provided by the present invention, the CO obtained in the separation unit can be sent to the water-gas shift unit to react with water vapor for further reaction to obtain hydrogen and enrich CO ₂ . The water-gas shift carried out by the CO and steam adopts the prior art well known to those skilled in the art.

In the method provided by the invention, preferably the deactivated catalyst enters the stripping section under the action of gravity, and the hydrocarbon product adsorbed on the deactivated catalyst is stripped out by water vapor, and the deactivated catalyst after stripping enters the regenerator.

Spent catalyst can be regenerated in conventional regenerators, either a single regenerator or multiple regenerators can be used. During the regeneration process, the oxygen-containing gas is generally introduced from the bottom of the regenerator. After the oxygen-containing gas is introduced into the regenerator, the deactivated catalyst contacts with oxygen and burns for regeneration. The flue gas enters the water gas shift unit. The oxygen-containing gas used for regeneration of the deactivated catalyst is preferably oxygen-enriched gas. The concentration of oxygen in the oxygen-containing gas at the bottom of the regenerator is 22-100% by volume, preferably 25%-80% by volume.

In the method provided by the present invention, preferably low temperature incomplete regeneration, operating conditions are: temperature is 550-700 ° C, preferably 600-650 ° C; gas apparent linear velocity is 0.2-1.2 m/s, preferably 0.4-0.8 m/s, loss The average residence time of the active catalyst is 1-10 minutes, preferably 2-6 minutes.

In the method provided by the invention, the volume ratio of CO/ _CO in the regenerated flue gas is 0.2-2.0, preferably 0.8-1.5.

In the method provided by the present invention, the operating conditions of the reducer of the regenerated catalyst are as follows: the temperature is 550-700°C; the apparent linear velocity of the gas is 0.5-3 m/s.

In the method provided by the invention, the reducing agent of the regenerated catalyst is selected from one or more of small molecular alkanes, preferably one or more of methane, ethane, propane, n-butane and isobutane.

In the method provided by the invention, the stripping gas of the regenerated catalyst is nitrogen.

In the method provided by the invention, the natural minerals in the catalyst are selected from kaolin, halloysite, montmorillonite, diatomaceous earth, attapulgite, sepiolite, halloysite, hydrotalcite, bentonite and One or more in retort soil, the content of natural minerals is 5% by weight-65% by weight on a dry basis, preferably 15% by weight-60% by weight; the oxide is silicon oxide, aluminum oxide One or more of , zirconia, titanium oxide, amorphous silica-alumina, based on the total amount of catalyst, in terms of oxide weight percentage, the content of oxide is 10% by weight-60% by weight, preferably 10% by weight - 30% by weight, more preferably 12% by weight - 28% by weight. The zeolite includes a medium-pore zeolite and an optional large-pore zeolite, and the large-pore zeolite is one or more selected from rare earth Y, rare earth hydrogen Y, ultrastable Y and high silicon Y. The medium-pore zeolite is ZSM series zeolite and/or ZRP zeolite, and the medium-pore zeolite accounts for 5-100% by weight of the total weight of the zeolite, preferably 20-50% by weight.

Based on the weight of the catalyst, the content of the metal active component is 0.1%-30% by weight, preferably 0.5-20% by weight. The metal active component is selected from one or more compounds of transition metal elements, preferably one or more of nickel, cobalt, iron, tungsten, molybdenum, manganese, copper, zirconium and chromium.

In the method provided by the invention, catalyst preparation method adopts the preparation method of conventional catalytic cracking catalyst, and this is the preparation method known to those skilled in the art. Loading metal on catalyst can adopt impregnation method, also can adopt slurry mixing method, preferred impregnation method, these methods are known to those skilled in the art.

The accompanying drawings are used to provide a further understanding of the present invention, and constitute a part of the description, together with the following specific embodiments, are used to explain the present invention, but do not constitute a limitation to the present invention.

As shown in Figure 1, the regenerated catalyst enters the reducer 23, and a reduction reaction occurs in contact with the reducing agent 16, and the regenerated catalyst after the reduction enters the stripper 24 and contacts the nitrogen gas 17 injected at the bottom, and the carbon monoxide and carbon dioxide carried by the catalyst are stripped The regenerated catalyst through reduction and stripping enters the fluidized reactor 1 through the regenerated catalyst of the regenerated inclined pipe 18, and accelerates upwards along the reactor under the effect of the pre-lift medium 22. After the water vapor in 6 is mixed, it is injected into the fluidized reactor 1 to contact with the regenerated catalyst, and the hydrocarbon raw material undergoes a catalytic reaction on the hot catalyst and accelerates upward. After the generated reaction product and the catalyst with carbon are separated by the cyclone separator 7 and collected in the gas collection chamber 8, the reaction product enters the separation unit 3 through the pipeline 9 and is separated into hydrogen 10, carbon monoxide 11, carbon dioxide 12, hydrocarbon products 13, hydrocarbon products Further separation can obtain low-carbon olefins such as ethylene, propylene and butene, low-carbon alkanes and other hydrocarbon products, and low-carbon alkanes can be used as the reducing agent of the present invention, and the remaining other hydrocarbon products can be further refined in part or all. transform. The separated carbon monoxide can be used as raw material for water gas shift, and light olefins can be further separated into ethylene.

After the catalyst with carbon enters the stripping section 4 and contacts the stripping steam 18 for stripping, it enters the regenerator 2 through the inclined pipe 19 and contacts the oxygen-enriched gas 14 that enters through the main air distribution plate 15, and a charcoal burning reaction occurs, and the loss of Activate the coke on the catalyst to regenerate the catalyst with carbon, and the regenerated flue gas enters the separation unit 3 through the separator 20 and the flue gas pipeline 21, and separates to obtain carbon monoxide 11 and carbon dioxide 12. The regenerated catalyst enters into the reducer 23 and the stripper 24 in turn, and the regenerated catalyst after reduction and stripping enters the fluidized reactor through the inclined pipe 18 for recycling.

The following examples will further illustrate the present invention, but do not thereby limit the present invention.

The raw materials used in the examples and comparative examples are all catalytic diesel oils, the properties of which are shown in Table 1. The used catalyst industrial catalyst of comparative example, trade mark is DMMC-1, and property is as shown in table 2.

The catalyst preparation method used in the embodiment is briefly described as follows:

1) Beat 75.4 kg of kaolin (solid content 71.6% by weight) with 250 kg of decationized water, then add 54.8 kg of pseudo-boehmite (solid content 63% by weight), adjust its pH to 2-4 with hydrochloric acid, and stir Evenly, stand and age at 60-70°C for 1 hour, keep the pH at 2-4, lower the temperature below 60°C, add 41.5 kg of aluminum sol (the content of Al ₂ O ₃ is 21.7% by weight), and stir for 40 minutes, A mixed slurry was obtained.

2) ZRP-1 (dry basis is 22 kg) and DASY zeolite (dry basis is 22.5 kg) are added to the obtained mixed slurry, stirred evenly, spray-dried and formed, and ammonium dihydrogen phosphate solution (phosphorus content is 1% by weight) ) washing, washing away free Na+, and roasting to obtain a molecular sieve catalyst sample.

3) Dissolve 3 kg of Ni(NO ₃ ) ₂ in 5.5 kg of water to prepare Ni(NO ₃ ) ₂ ·6H ₂ O aqueous solution, and immerse 10 kg of molecular sieve catalyst samples in Ni(NO ₃ ) ₂ ·6H ₂ O aqueous solution , the resulting mixture was dried at 180°C for 4 hours and calcined at 600°C for 2 hours. After repeated impregnation, drying and roasting, the content of Ni supported on the catalyst sample reached 15%, and the catalyst A of the embodiment was obtained.

Example 1

Carry out test according to the flow process of Fig. 1, carry out the catalytic decomposition reaction test of catalyzed diesel oil on the riser reactor, catalyzed diesel oil enters the lower part of riser reactor, contacts with the regenerated catalyst of heat and carries out catalytic decomposition reaction, reaction product and deactivation The catalyst enters the closed cyclone separator from the outlet of the reactor, the reaction product and the deactivated catalyst are separated rapidly, and the reaction product is separated into cracked gas and liquid according to the distillation range in the separation system.

The deactivated catalyst enters the stripping section under the action of gravity, and the hydrocarbon products adsorbed on the deactivated catalyst are stripped by water vapor, and the deactivated catalyst after stripping enters the regenerator, and is regenerated by contacting with oxygen-rich air; regeneration The final catalyst is sequentially contacted with methane to undergo a reduction reaction, and then stripped by nitrogen and returned to the riser reactor for recycling. The operating conditions and product distribution are listed in Table 3.

From the results in Table 3, it can be seen that the yield of hydrogen is as high as 18.33%, the yield of ethylene is 2.38%, the yield of propylene is 6.05%, the yield of CO in the reaction product is _28.53 %, and the yield of CO is 25.64%. The CO concentration in the regeneration flue gas was 7.60 vol%, and the _CO2 concentration was 25.33 vol%.

Comparative example 1

On the medium-sized device of riser, carry out test, catalytic diesel oil raw material is identical with embodiment 1, and catalyzer is DMMC-1.

Catalyzed diesel oil enters the riser reactor and the hot DMMC-1 catalyst contacts and undergoes catalytic decomposition reaction. The reaction product and deactivated catalyst enter the closed cyclone separator from the reactor outlet, and the reaction product and deactivated catalyst are separated rapidly. The separation system separates products such as gas and liquid according to the distillation range.

The deactivated catalyst enters the stripping section under the action of gravity, and the hydrocarbon products adsorbed on the deactivated catalyst are stripped by water vapor, and the deactivated catalyst after stripping enters the regenerator and is regenerated by contacting with air; the regenerated catalyst Return to the riser reactor and recycle again; Operating conditions and product distribution are listed in Table 3.

As can be seen from the results in Table 3, the yield of hydrogen was 1.42%, the yield of ethylene was 3.69%, the yield of propylene was 9.52%, and the yields of CO _and CO in the reaction products were relatively low. No CO was detected in the regeneration flue gas, the _CO2 concentration was 15.56 vol%, and the _O2 concentration was 3.50 vol%.

The preferred embodiments of the present invention have been described in detail above, but the present invention is not limited to the specific details in the above embodiments. Within the scope of the technical concept of the present invention, various simple modifications can be made to the technical solutions of the present invention. These simple Modifications all belong to the protection scope of the present invention.

In addition, it should be noted that the various specific technical features described in the above specific embodiments can be combined in any suitable manner if there is no contradiction. In order to avoid unnecessary repetition, the present invention includes various possible The combination method will not be explained separately.

In addition, various combinations of different embodiments of the present invention can also be combined arbitrarily, as long as they do not violate the idea of the present invention, they should also be regarded as the content of the invention of the present invention.

Table 1

project catalytic diesel <![CDATA[Density(20℃)/(kg/m³)]]> 948.9 <![CDATA[kinematic viscosity (20°C)/(mm²/sec)]]> 3.754 Total acid value/(mg KOH/g) <0.05 Freezing point/℃ -27 Closed flash point/℃ 66 Distillation range/℃ the initial boiling point 197.9 5% 220.4 10% 228.4 50% 262.5 70% 283.1 90% 308.6 95% 319.5 end point 329.9

Table 2

the Example 1 Comparative example 1 the Catalyst A Industrial catalyst DMMC-1 physical properties the the <![CDATA[Specific Surface Area, m²/gram]]> 118 106 <![CDATA[Specific surface area of molecular sieve, m²/gram]]> 30 42 <![CDATA[pore volume, cm³/gram]]> 0.126 0.13 Sieve composition, wt% the the 0～40 microns 22.7 30.9 0～80 microns 64.4 75.0 0～105 microns 87.1 89.1 0～149 microns 97.9 98.4 Average particle size/micron 55.0 55.6 Microreactivity, % 50 65 Metal content, % the the Ni 14.5 0.06

table 3

the Example 1 Comparative example 1 Riser reaction conditions the the Reaction temperature, °C 675 675 Response time, seconds 1.8 1.8 Catalyst to catalytic diesel weight ratio 20 20 Water to Oil Weight Ratio 0.25 0.25 regeneration condition the the regeneration temperature, ℃ 645 680 Oxygen concentration in regeneration air, volume % 50 twenty one Gas superficial linear velocity, m/s 0.8 0.8 Average dwell time, minutes 4.5 4.5 Reduction condition the the Reduction temperature, ℃ 650 / Weight ratio of reductant to catalyst circulation 0.005 / Product distribution, weight % the the CO 28.53 0.25 <![CDATA[CO₂]]> 25.64 1.01 <![CDATA[H₂]]> 18.33 1.42 Vinyl 2.38 3.69 Propylene 6.05 9.52 Other gaseous hydrocarbons 11.09 30.47 liquid hydrocarbon 2.67 43.53 coke 5.31 10.11 total 100.00 100.00 Composition of regeneration flue gas, volume % the the CO 7.60 / <![CDATA[CO₂]]> 25.33 15.56 <![CDATA[N₂]]> 67.05 80.94 <![CDATA[O₂]]> 0.02 3.5

Claims (17)

1. A method for producing hydrogen and light olefins simultaneously by using a hydrocarbon feedstock, comprising the following steps: Send the hydrocarbon raw material into the fluidized reactor, and contact with the catalyst containing zeolite and metal oxide to generate carbon-hydrogen bond cleavage reaction to generate hydrogen and carbon-carbon bond cleavage reaction to generate low-carbon olefins to obtain a reaction containing hydrogen and low-carbon olefins Oil and gas and carbon catalysts; Separation of the reaction oil gas containing hydrogen and light olefins from the carbon catalyst; The separated reaction oil gas is further separated into hydrogen, carbon monoxide, carbon dioxide, light olefins and other hydrocarbon products; Send the obtained carbon-coated catalyst to the regenerator for charcoal-burning reaction. After the carbon-coated catalyst is regenerated by charring, it is sent to the reducer as a regenerated catalyst, and the regenerated flue gas enters the separation unit to separate and obtain carbon monoxide and carbon dioxide; The regenerated catalyst enters the reducer and contacts with the reducing agent to undergo a reduction reaction, and the reduced regenerated catalyst enters the fluidized reactor after being stripped. 2. The method according to claim 1, characterized in that the operating conditions of the reducer are as follows: the temperature is 550-700° C.; the apparent linear velocity of the gas is 0.5-3 m/s. 3. The method according to claim 1, characterized in that, the reductant is selected from one or more of small molecular alkanes, preferably one of methane, ethane, propane, n-butane and isobutane species or several. 4. The method according to claim 1, characterized in that the reaction temperature of the fluidized reactor is 450-800° C., the reaction time is 0.1-10 seconds, the weight ratio of catalyst to hydrocarbon feedstock is 5-100, water The weight ratio of steam to hydrocarbon feedstock is 0.1-20. 5. The method according to claim 1, wherein the reaction conditions are: the reaction temperature is 550-700°C, the reaction time is 1-8 seconds, the weight ratio of the catalyst to the hydrocarbon feedstock is 20-50; The weight ratio to hydrocarbon feedstock is 1-10. 6. The method according to claim 1, characterized in that the fluidized reactor is selected from one or more combinations of riser reactors, fast beds and dense-phase fluidized beds. 7. The method according to claim 6, characterized in that, the fluidized reactor comprises the fluidized reactor of a pre-lift section and at least one reaction zone successively from bottom to top, and the preferred 2-8 in the said reaction zone , more preferably 2-3. 8. The method according to claim 1, characterized in that the concentration of oxygen in the oxygen-containing gas at the bottom of the regenerator is 22-100% by volume. 9. The method according to claim 1, characterized in that the concentration of oxygen in the oxygen-containing gas at the bottom of the regenerator is 25-80% by volume. 10. The method according to claim 1, characterized in that, the regeneration operating conditions are as follows: the temperature is 550-700°C; the apparent linear velocity of the gas is 0.2-1.2 m/s, and the average residence time of the deactivated catalyst is 1 -10 minutes. 11. The method according to claim 1, characterized in that, the regeneration operating conditions are as follows: the temperature is 600-650°C; the apparent linear velocity of the gas is 0.4-0.8 m/s, and the average residence time of the deactivated catalyst is 2 -6 minutes. 12. The method according to claim 1, characterized in that the CO/CO ₂ volume ratio in the regenerated flue gas is 0.2-2.0, preferably 0.8-1.5. 13. The method according to claim 1, wherein the hydrocarbon feedstock is selected from one or more mixtures of petroleum hydrocarbons, mineral oils and synthetic oils, and the petroleum hydrocarbons are gaseous hydrocarbons , gasoline, diesel, vacuum gas oil, atmospheric residue, vacuum gas oil mixed with part of vacuum residue or hydrocarbon oil obtained from secondary processing; mineral oil is selected from coal liquefied oil, oil sands oil and shale oil One or a mixture of more than one of them; synthetic oil is a distillate obtained by Fischer-Tropsch synthesis of coal, natural gas or asphalt. 14. The method according to claim 13, the hydrocarbon oil obtained by said secondary processing is selected from coker gasoline, catalytic diesel oil, hydrogenated diesel oil, coker diesel oil, coker wax oil, deasphalt oil, furfural refining raffinate one or several. 15. The method according to claim 1, characterized in that, based on the dry basis weight of the catalyst, the catalyst comprises 5%-65% of natural minerals, 10%-60% of oxides, 10 %-60% zeolite and 0.1%-30% metal active component, said zeolite includes medium-pore zeolite and optional large-pore zeolite, said medium-pore zeolite is ZSM series zeolite and/or ZRP zeolite, said large-pore zeolite The pore zeolite is one or more selected from rare earth Y, rare earth hydrogen Y, ultrastable Y and high silicon Y. 16. The method according to claim 15, characterized in that the mesoporous zeolite accounts for 5-100% by weight, preferably 20-50% by weight, of the total weight of the zeolite. 17. The method according to claim 15, characterized in that the content of the metal active component is 0.1%-30% by weight, preferably 0.5-20% by weight. The metal active component is selected from one or more compounds of transition metal elements, preferably one or more of nickel, cobalt, iron, tungsten, molybdenum, manganese, copper, zirconium and chromium.

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