A method and apparatus for hydrogen production from hydrocarbons

Abstract

The application discloses a method and device for hydrogen production from hydrocarbon, which comprises the following steps: introducing a catalyst into an activation zone of a fluidized reactor, and contacting the catalyst with an activation agent to obtain an activated catalyst; preheating a hydrocarbon raw material, atomizing the preheated hydrocarbon raw material by using an atomizing gas, and introducing the atomized hydrocarbon raw material from the bottom of a dehydrogenation zone of the fluidized reactor to contact the activated catalyst and perform a dehydrogenation reaction; introducing water vapor from a gas inlet at the lower part of a gasification zone of the fluidized reactor to perform a gasification reaction with carbon-containing catalyst and reaction oil gas from the dehydrogenation zone; and separating reaction products and carbon deposition catalyst to obtain hydrogen-containing generated gas and carbon deposition catalyst. The application has high adaptability to raw materials, and hydrogen production from hydrocarbon is performed by using a fluidized bed process, so that the hydrogen production rate is high.

Description

TECHNICAL FIELD

[0001] The present application relates to a method and device for hydrogen production from hydrocarbons. BACKGROUND

[0002] With the development trend of crude oil becoming heavier and poorer, and the further improvement of clean fuel production standards, the demand for hydrogen is also increasing. At present, China's hydrogen energy market is in its early stages of development, with an average annual demand of about 22 million tons of hydrogen. It is estimated that by 2060, China's average annual demand for hydrogen will reach about 130 million tons.

[0003] Currently, hydrogen is mainly derived from natural gas steam reforming, which is the most mature and commonly used method. However, due to the strong endothermic and reversible nature of the reaction process, natural gas steam reforming for hydrogen production has the disadvantages of high reaction temperature, high energy consumption, complex process, short catalyst life, etc. At the same time, due to the scarcity of crude oil resources in China and the heavy nature of crude oil, in addition to producing high-value gasoline and diesel products, a certain amount of catalytic diesel and other inferior oil products, as well as low-value hydrocarbon products such as dry gas and liquefied gas, are also produced in the refining industry. Therefore, developing a hydrogen production technology that can achieve high-value utilization of inferior oil or low-value hydrocarbons not only creates benefits for refineries, but also meets the basic needs of China's national conditions.

[0004] CN103183318B discloses a mobile bed light hydrocarbon steam conversion method and device for hydrogen production. Light hydrocarbons and water vapor are preheated and then enter the reactor from the bottom, where they are in countercurrent contact with reforming catalysts and adsorption catalysts of different particle sizes. Light hydrocarbons and water vapor react under steam reforming conditions to produce H2, CO, and CO2. The light hydrocarbon feedstock in this method is still mainly methane, and the reforming catalyst is not effectively regenerated.

[0005] CN103373706B discloses a method and device for hydrogen production by methane reforming. Hydrogen and methane are first contacted with a composite catalyst for catalyst reduction. Methane and water vapor enter the reforming reactor from the middle. The separated product gas enters the subsequent separation and purification device to obtain high-purity hydrogen, eliminating the need for a reduction device after regeneration of the composite catalyst.

[0006] From the above-mentioned patents, it can be seen that in the existing methods for hydrogen production using fluidized reactors, the reaction feedstock is mainly light hydrocarbons such as methane, natural gas, and oil field gas. The reaction feedstock is relatively single, and there are problems such as long process flow, complex process, and harsh production conditions. Therefore, in order to meet the market demand for hydrogen, expand the source range of hydrocarbon feedstock, and reduce the energy consumption of the reaction process, it is necessary to develop a production device and method for hydrogen production from hydrocarbon fluidized beds. This not only increases the production of chemical raw material hydrogen, but also realizes the high-value utilization of inferior oil or low-value hydrocarbons in refineries, achieving the effect of energy saving and emission reduction. SUMMARY

[0007] It is one of the objectives of the present application to provide a method for producing hydrogen from hydrocarbons.

[0008] It is another objective of the present application to provide a device for producing hydrogen from hydrocarbons.

[0009] The method for producing hydrogen from hydrocarbons provided by the present application comprises the following steps:

[0010] (1) introducing catalyst into the activation zone of the fluidized reactor to contact with activation agent to obtain activated catalyst;

[0011] (2) preheated hydrocarbon raw material is atomized by atomizing gas and introduced from the bottom of the dehydrogenation zone of the fluidized reactor to contact with activated catalyst to perform dehydrogenation reaction;

[0012] (3) water vapor is introduced from the gas inlet at the lower part of the gasification zone of the fluidized reactor to perform gasification reaction with carbon-containing catalyst and reaction oil gas from the dehydrogenation zone, and the reaction product is separated from the carbon deposition catalyst to obtain hydrogen-containing product gas and carbon deposition catalyst.

[0013] The activation agent is selected from one or more of C1-C2 hydrocarbons, C1-C4 alcohols, H2 and CO. The weight ratio of the activation agent to the hydrocarbon raw material is 0.001-1, preferably 0.001-0.5.

[0014] The hydrocarbons are selected from one or a mixture of more than one of petroleum hydrocarbons, mineral oil and synthetic oil, the petroleum hydrocarbons are gaseous hydrocarbons, gasoline, diesel, vacuum wax oil, atmospheric residue, vacuum wax oil mixed with part of the vacuum residue or hydrocarbon oil obtained by secondary processing, the mineral oil is selected from one or a mixture of more than one of coal liquefaction oil, oil sand oil and shale oil, and the synthetic oil is distillate oil obtained by F-T synthesis of coal, natural gas or asphalt. The hydrocarbon oil obtained by secondary processing is selected from one or several of coking gasoline, catalytic diesel, hydrogenated diesel, coking diesel, coking wax oil, deasphalted oil, and furfural refined raffinate oil.

[0015] The atomizing gas is nitrogen.

[0016] The reaction conditions of the dehydrogenation zone are as follows: the reaction temperature is 580-780°C, preferably 580-750°C, the reaction time is 0.1-3 seconds, preferably 0.3-2.8 seconds, and the weight ratio of catalyst to hydrocarbon raw material is 4-40, preferably 15-30.

[0017] The reaction conditions of the gasification zone are as follows: the reaction temperature is 550-760°C, preferably 550-730°C, the gasification time is 0.5-8 seconds, preferably 0.8-7 seconds, and the weight ratio of water vapor to hydrocarbon raw material is 0.01-5.

[0018] The catalyst comprises 5-65% of natural mineral, 10-60% of oxide, 20-60% of large pore zeolite and 0.1-30% of metal active component, based on the dry weight of the catalyst.

[0019] The natural mineral in the catalyst is selected from one or more of kaolin, halloysite, montmorillonite, diatomite, boehmite, sepiolite, halloysite, hydrotalcite, bentonite and rectorite, and the content of the natural mineral is 5-65% by weight, preferably 15-60% by weight, based on the dry weight.

[0020] The oxide is one or more of silicon oxide, aluminum oxide, zirconium oxide, titanium oxide and amorphous silicon aluminum, and the content of the oxide is 10-60% by weight, preferably 10-30% by weight, more preferably 12-28% by weight, based on the total weight of the catalyst.

[0021] The large pore zeolite is one or more of rare earth Y, rare earth hydrogen Y, ultra-stable Y and high-silicon Y.

[0022] The content of the metal active component is 0.1-30% by weight, preferably 5-30% by weight, more preferably 8-25% by weight. The metal active component is one or more of compounds of transition metal elements, preferably one or more of nickel, cobalt, iron, tungsten, molybdenum, manganese, copper, zirconium and chromium.

[0023] The coked catalyst obtained in step (3) is directly returned to the activation zone of the fluidized reactor after stripping, or / and is returned to the activation zone of the fluidized reactor after regeneration with oxygen-containing gas. The regeneration operation conditions are: temperature of 600-800°C; gas superficial linear velocity of 0.2-1.2 m / s, and average residence time of the coked catalyst of 1-10 minutes. The oxygen-containing gas is oxygen or / and air.

[0024] The fluidized reactor is selected from one or a combination of more than one of riser reactor, isokinetic fluidized bed reactor, equal-diameter fluidized bed reactor, upward conveying line and downward conveying line.

[0025] The hydrogen-containing generated gas is purified to obtain hydrogen.

[0026] The device for hydrogen production from hydrocarbons provided by the present application comprises a fluidized reactor, a settler, a stripper, a separation and purification unit and an optional regenerator, wherein the fluidized reactor comprises an activation zone, a dehydrogenation zone and a gasification zone, the fluidized reactor is communicated with the settler, the upper part of the settler is connected with the separation and purification unit through a pipeline, the lower part of the settler is communicated with the stripper, the stripper is directly connected with the activation zone of the fluidized reactor through a catalyst pipeline, or / and the stripper is connected with the activation zone of the fluidized reactor through the regenerator and the catalyst pipeline.

[0027] Compared with the prior art, the present application has the following advantages:

[0028] 1. The catalyst used in the reaction device and method is activated in the activation zone of the fluidized bed reactor after oxidation regeneration, and a separate catalyst reduction device is not needed.

[0029] 2. The reaction device and method have strong raw material applicability, and are suitable not only for light hydrocarbons such as gaseous hydrocarbons, straight-run naphtha, straight-run diesel and the like, but also for hydrocarbon oils obtained by secondary processing, such as one or more of coking gasoline, catalytic diesel, hydrogenated diesel, coking diesel, coking wax oil, deasphalted oil and furfural refined raffinate oil.

[0030] 3. The hydrocarbon raw material and water vapor are fed in different reaction zones to first undergo a dehydrogenation reaction and then undergo a gasification reaction, which not only strengthens the contact and reaction between the hydrocarbon raw material and the catalyst, but also can improve the hydrogen production rate and reduce the difficulty of the separation process. BRIEF DESCRIPTION OF DRAWINGS

[0031] The accompanying drawings are included to provide a further understanding of the present application and are incorporated in and constitute a part of the specification, illustrate embodiments of the present application and serve to explain the present application and do not limit the present application. In the drawings:

[0032] Figure 1 、 Figure 2 The accompanying drawings are included to provide a further understanding of the present application and are incorporated in and constitute a part of the specification, illustrate embodiments of the present application and serve to explain the present application and do not limit the present application. In the drawings:

[0033] DETAILED DESCRIPTION

[0034] 1 riser reactor 2 regenerator 3 settler

[0035] 4 separation and purification unit 5 heater 10 activation zone

[0036] 11 dehydrogenation zone 12 gasification zone 13 stripping section

[0037] 14 cyclone separator 15 gas collecting chamber 21 air distributor

[0038] 22 cyclone separator 23 spent catalyst transfer chute 24 regenerated catalyst transfer chute

[0039] 31-42 pipelines DETAILED DESCRIPTION

[0040] The specific embodiments of the present application will be described in detail below with reference to the accompanying drawings.

[0041] Figure 1 For one specific embodiment of the process of the present application, catalyst regeneration is included. The process is as follows:

[0042] The activator 31 enters the activation zone 10 from the bottom of the riser reactor 1, contacts the regenerated catalyst from the regenerated catalyst transfer chute 24, activates and fluidizes the catalyst, and flows upward along the riser with the regenerated catalyst; the hydrocarbon feedstock 33, after being preheated by the heater 5 with the atomizing gas 32, enters the bottom of the dehydrogenation zone 11, contacts the activated catalyst, and reacts to produce reaction oil gas and carbon-laden catalyst, which enter the gasification zone 12; the water vapor 34 enters from the gas inlet at the bottom of the gasification zone 12, contacts the reaction oil gas and carbon-laden catalyst from the dehydrogenation zone 11, and reacts to produce hydrogen-containing product gas and carbon-laden catalyst, which enter the cyclone separator 14 in the settler 3, and are separated to produce hydrogen-containing product gas and carbon-laden catalyst, respectively; the hydrogen-containing product gas enters the gas collection chamber 15, and is separated into hydrogen gas 41, carbon dioxide 39, carbon monoxide 40, and hydrocarbon product 42 in the separation and purification unit 4; the hydrocarbon product 42 can be further converted in part or in whole by recycling.

[0043] The carbon-laden catalyst fines return to the settler 3 from the cyclone separator 14, and the carbon-laden catalyst in the settler 3 flows to the stripping section 13, contacts the stripping water vapor from the pipeline 35, and the reaction products stripped from the carbon-laden catalyst are separated by the cyclone separator 14 and enter the gas collection chamber 15. The stripped carbon-laden catalyst is adjusted by the spent catalyst slide valve, enters the regenerator 2 through the spent catalyst transfer chute 23, and contacts the air from the pipeline 36 after passing through the air distributor 21, which burns off the coke on the carbon-laden catalyst in the regenerator 2 to regenerate the deactivated carbon-laden catalyst; the regenerated flue gas is separated by the cyclone separator 22, enters the subsequent energy recovery system through the upper gas flue 37; and the regenerated catalyst is adjusted by the regenerated catalyst slide valve of the regenerated catalyst transfer chute 24, and returns to the activation zone 10 of the riser reactor 1.

[0044] Figure 2 For another embodiment of the present application, unlike the above embodiment, Figure 1 the stripped catalyst directly returns to the activation zone 10 of the riser reactor through the catalyst transfer chute, and the required heat is supplied externally.

[0045] The following examples will further illustrate the present application but do not limit the present application.

[0046] The catalyst used in the examples is a modified commercial catalyst, the commercial catalyst is CDOS, and the modification method is as follows:

[0047] The commercial catalyst CDOS was dried at 200°C for 3 hours and then calcined at 650°C for 8 hours to obtain a catalyst intermediate. 200 g of the catalyst intermediate was immersed in a 300 g solution containing nickel nitrate (20% by weight) at room temperature and normal pressure for 12 hours. The immersed catalyst intermediate was filtered at room temperature and 80 kPa to obtain a filter cake of 180 g, which was dried and calcined at 550°C for 2 hours to obtain catalyst A with a Ni content of 10.71% by weight.

[0048] The hydrocarbon feedstock used in Examples 1-4 and Comparative Example is catalytic diesel, the properties of which are listed in Table 1. The hydrocarbon feedstock used in Example 5 is n-butane.

[0049] Example 1

[0050] This example is used to illustrate the effect of the method for fluidized bed hydrogen production provided by the present application.

[0051] According to the flow shown in Figure 1 , a test was carried out on a pilot plant riser reactor using catalytic diesel and catalyst A, and H2 was used as the activating agent to activate the regenerated catalyst. The catalytic diesel was atomized by N2 and then entered the reactor from the bottom of the dehydrogenation zone to contact with the activated catalyst for dehydrogenation reaction to obtain reaction oil gas and carbon-containing catalyst. The reaction oil gas and carbon-containing catalyst entered the gasification zone to react with steam from the bottom of the gasification zone to obtain hydrogen-containing product gas and carbonized catalyst, which entered a closed cyclone separator from the outlet of the reactor to separate the reaction products from the carbonized catalyst. The reaction products were separated into cracked gas and liquid according to the distillation range in the separation system, and the carbonized catalyst was stripped and then entered the regenerator for regeneration using air. The reaction process operating conditions, hydrogen production rate and cracked gas product distribution are shown in Table 2.

[0052] Example 2

[0053] According to the flow shown in Figure 1 , a test was carried out on a pilot plant riser reactor using methanol as the activating agent to activate the regenerated catalyst. The reaction process operating conditions, hydrogen production rate and cracked gas product distribution are shown in Table 3.

[0054] Example 3

[0055] According to the flow shown in Figure 1The process shown was tested on a medium-sized riser reactor. CO was used as an activator to activate the regenerated catalyst. The operating conditions, hydrogen yield, and distribution of cracked gas products are shown in Table 3.

[0056] Example 4

[0057] according to Figure 1 The process shown was tested on a medium-sized riser reactor. CH4 was used as an activator to activate the regenerated catalyst. The operating conditions, hydrogen yield, and distribution of cracked gas products are shown in Table 3.

[0058] Example 5

[0059] according to Figure 2 The process described was tested on a medium-sized riser reactor. n-Butane was used as the feedstock, catalyst A was used, and H2 was used as the activator to activate the catalyst from the stripping zone. Catalyst diesel, after being atomized by N2, entered the reactor from the bottom of the dehydrogenation zone, contacting the activated catalyst to undergo a dehydrogenation reaction, yielding reactant oil and carbonized catalyst. The reactant oil and carbonized catalyst then entered the gasification zone, reacting with water vapor from the bottom of the gasification zone. The resulting hydrogen-containing product gas and catalyst entered a closed cyclone separator from the reactor outlet to separate the reaction products from the catalyst. The reaction products were separated into cracked gas and liquid according to their distillation range in the separation system. The catalyst, after stripping, was directly returned to the activation zone. The heat required for the reaction was supplied by electric heating. The operating conditions, hydrogen yield, and distribution of cracked gas products are shown in Table 3.

[0060] Comparative Example 1

[0061] The experiment was conducted on a medium-sized riser reactor using catalytic diesel as feedstock and catalyst A. Unlike Example 1, steam was used for catalyst activation and fluidization. The reaction process operating conditions, the yield of hydrogen obtained, and the gas phase composition are shown in Table 2.

[0062] Comparative Example 2

[0063] according to Figure 1 The process shown was tested on a medium-sized riser reactor using catalytic diesel as feedstock and an unmodified industrial catalyst, commercially available under the brand name CDOS. The remaining reaction conditions were the same as in Example 1. The operating conditions, hydrogen yield, and gas phase composition of the reaction process are shown in Table 2.

[0064] Comparative Example 3

[0065] according to Figure 1The distillation range, the test was carried out on a pilot plant riser reactor, using catalytic diesel as raw material, using catalyst A, different from example 1, the gasification zone was not connected with water vapor, the operation conditions of the reaction process, the hydrogen yield and the gas phase composition are shown in table 2.

[0066] From the results of the above examples and comparative examples, it can be seen that the method and catalyst provided by the application can significantly improve the hydrogen yield.

[0067] Table 1 Catalytic diesel properties

[0068] Density (20°C) / (kg / m3 3 )]]> 961.1 Viscosity (50°C) / (mm 2 / sec) 2.128 Elemental mass composition / wt% C 90.99 H 9.01 S 0.473 N 0.068 Distillation range / °C IBP 173 10％ 229 50％ 259 90％ 323 95％ 336 EBP 358

[0069] Table 2

[0070]

[0071] Table 3

[0072]

Claims

1. A method for producing hydrogen from hydrocarbons, comprising the following steps: (1) The catalyst is introduced into the activation zone of the fluidized bed reactor and contacted with the activator to obtain the activated catalyst. The activator is selected from one or more of C1-C2 hydrocarbons, C1-C4 alcohols, H2 and CO. The weight ratio of the activator to the hydrocarbon feedstock is 0.001-1. Based on the dry weight of the catalyst, the catalyst includes 5%~65% natural minerals, 10%~60% oxides, 20%~60% macroporous zeolite and 0.1%~30% metal active components. (2) After the preheated hydrocarbon feedstock is atomized by atomizing gas, it is introduced from the bottom of the dehydrogenation zone of the fluidized bed reactor and comes into contact with the activated catalyst to carry out the dehydrogenation reaction. The reaction conditions in the dehydrogenation zone are: reaction temperature of 580-780℃, reaction time of 0.1-3 seconds, and the weight ratio of catalyst to hydrocarbon feedstock of 4-40. (3) Water vapor is introduced from the lower gas inlet of the gasification zone of the fluidized bed reactor and reacts with the carbonized catalyst and reaction oil gas from the dehydrogenation zone. The reaction products and the carbonized catalyst are separated to obtain hydrogen-containing generated gas and carbonized catalyst. The reaction conditions in the gasification zone are: reaction temperature 550-760℃, gasification time 0.5-8 seconds, and the weight ratio of water vapor to hydrocarbon feedstock is 0.01-5.

2. The method according to claim 1, characterized in that, The weight ratio of the activator to the hydrocarbon raw material is 0.001-0.

5.

3. The method according to claim 1, characterized in that, The hydrocarbons are selected from catalytic diesel oil and n-butane.

4. The method according to claim 1, characterized in that, The atomizing gas is nitrogen.

5. The method according to claim 1, characterized in that, The dehydrogenation reaction conditions are as follows: reaction temperature of 580-750℃, reaction time of 0.3-2.8 seconds, and the weight ratio of catalyst to hydrocarbon feedstock of 15-30.

6. The method according to claim 1, characterized in that, The reaction conditions in the vaporization zone are: reaction temperature 550-730℃, vaporization time 0.8-7 seconds.

7. The method according to claim 1, characterized in that, The content of the metal active component is 5% to 30% by weight, and the metal active component is selected from one or more compounds of transition metal elements.

8. The method according to claim 7, characterized in that, The content of the metal active component is 8% to 25% by weight.

9. The method according to claim 7, characterized in that, The active metal component is selected from one or more of nickel, cobalt, iron, tungsten, molybdenum, manganese, copper, zirconium, and chromium.

10. The method according to claim 1, characterized in that, The carbonized catalyst obtained in step (3) is stripped and then directly returned to the activation zone of the fluidized bed reactor, or / and regenerated by oxygen-containing gas and then returned to the activation zone of the fluidized bed reactor.

11. The method according to claim 10, characterized in that, The regeneration operation conditions are: temperature of 600-800℃; apparent gas linear velocity of 0.2-1.2 m / s; and average residence time of the coked catalyst of 1-10 minutes.

12. The method according to claim 10, characterized in that, The oxygen-containing gas is oxygen and / or air.

13. The method according to claim 1, characterized in that, The fluidized bed reactor is selected from one or more of the following: riser reactor, constant linear velocity fluidized bed reactor, constant diameter fluidized bed reactor, upward conveyor line, and downward conveyor line.

14. The method according to claim 1, characterized in that, The hydrogen-containing generated gas is purified to obtain hydrogen gas.

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