Continuous activation apparatus and activation method for forming pure-phase epsilon / epsilon' iron carbide catalyst for fischer-tropsch synthesis

Abstract

The application discloses a continuous activation device and method for forming a pure-phase epsilon / epsilon' iron carbide catalyst for Fischer-Tropsch synthesis, and the continuous activation device comprises a catalyst feeding tank, a first fluidized bed reactor, a first storage tank, a second fluidized bed reactor, a first synthetic gas pipe, a second storage tank, a third fluidized bed reactor, a second synthetic gas pipe and a catalyst finished product tank; and the continuous activation of the pure-phase epsilon / epsilon' iron carbide catalyst for Fischer-Tropsch synthesis is realized so as to be industrialized.

Description

Technical Field

[0001] This invention relates to the field of Fischer-Tropsch synthesis catalyst technology, specifically to a continuous activation apparatus and activation method for forming a pure-phase ε / ε' iron carbide catalyst for Fischer-Tropsch synthesis. Background Technology

[0002] my country's energy structure is characterized by abundant coal, scarce oil, and limited natural gas. Therefore, converting raw coal into high-quality liquid fuels and high-value-added chemicals has always been a focus of attention. Coal indirect liquefaction (CLD) is a process that uses coal as raw material, transforming it into clean syngas through gasification and purification, and then converting it into large-molecule hydrocarbons and alcohols via Fischer-Tropsch synthesis. It is an important pathway to achieve the rational, efficient, and clean utilization of coal.

[0003] The Fischer-Tropsch synthesis reaction is the core part of the indirect coal liquefaction process. It is the process of synthesizing hydrocarbon products from syngas (H2+CO) in the presence of metal catalysts such as iron, cobalt, and ruthenium.

[0004] New technologies have led to the development of pure-phase iron carbide catalysts for Fischer-Tropsch synthesis. For example, CN112569987A and CN112569993A both describe methods for preparing pure-phase iron carbide for Fischer-Tropsch synthesis. The 2018 article "Synthesis of stable and low-CO2selective ε-iron carbide Fischer-Tropsch catalysts," published in the journal *Science Advances*, also introduced the first synthesis of pure-phase ε(′)-Fe2C with near-zero CO2 selectivity. Catalysts formed with pure-phase iron carbide as the main component exhibit significantly higher performance than existing catalysts formed with various components such as elemental iron, magnetite, and multiple types of iron carbide. Their most significant advantage is the ability to significantly reduce the CO2 content, a byproduct of Fischer-Tropsch synthesis, making them highly suitable for the efficient production of oil and wax products in modern coal chemical Fischer-Tropsch synthesis large-scale industries.

[0005] For Fischer-Tropsch synthesis iron-based catalysts used in industrial applications, due to requirements in catalyst preparation and molding processes, storage and transportation, and catalyst feeding processes, the iron is in an oxidized state when it leaves the factory. It needs to be activated by reducing gases such as hydrogen, carbon monoxide, and low-carbon olefins before the Fischer-Tropsch synthesis reaction to form a stable active phase for the reaction. Therefore, for industrial Fischer-Tropsch synthesis catalysts aiming to use pure-phase iron carbide as the active component, the activation process is the process of converting the iron oxide component, which plays a major catalytic role, into the pure-phase iron carbide component. Thus, the activation process has a crucial impact on the Fischer-Tropsch synthesis performance of iron-based industrial catalysts. Currently, there are patent reports on methods for forming pure-phase iron carbide, but the process is relatively complex and difficult to apply industrially; while existing industrial activation methods for Fischer-Tropsch iron-based catalysts cannot form pure-phase iron carbide. CN112569982A, CN112569993A, and CN112569987A disclose methods for preparing catalysts containing precipitated pure-phase ε / ε' iron carbide, supported pure-phase ε / ε' iron carbide, and pure-phase ε / ε' iron carbide catalysts or iron carbide-containing composite catalysts, respectively. These methods incorporate an activation process into the catalyst preparation process. The main conditions involve first reducing the prepared iron oxide or iron oxide-containing precursor under hydrogen conditions, followed by a low-temperature pretreatment reaction using syngas, and finally preparing pure-phase ε / ε' iron carbide under syngas conditions. However, these technologies belong to the laboratory synthesis stage and do not provide industrial-scale activation equipment or activation process flows. Industrially produced catalysts are in an oxidized state upon delivery and must be activated into an active iron phase capable of Fischer-Tropsch synthesis using a separate industrial device before being transferred to a Fischer-Tropsch synthesis reactor for the Fischer-Tropsch synthesis reaction. If the above-mentioned activation process for forming pure phase iron carbide were to be applied industrially, the existing equipment and processes would not be able to meet the requirements due to the complexity of the procedure and process, and new reactor configurations and process flows would need to be designed.

[0006] CN112569988A, CN112569990A, and CN112569995A are compositions containing two pure phases of iron carbide formed by combining precipitated ε / ε' iron carbide, supported ε / ε' iron carbide, and ε / ε' iron carbide with θ iron carbide, respectively.

[0007] CN112569975A discloses a method for first forming pure phase χ iron carbide, pure phase θ iron carbide, and pure phase ε / ε iron carbide, respectively. ’ Iron carbide, and then a multiphase iron carbide catalyst.

[0008] CN103551207A discloses a fixed fluidized bed or gas-solid bubbling bed Fischer-Tropsch catalyst reduction and activation system and process, which includes a cyclone separator to recover catalyst entrained in the gas, a catalyst metering scale to control the mass of catalyst entering the Fischer-Tropsch reactor, and a gas mixer to regulate the reducing gas atmosphere. The system operates under the following conditions: temperature 200–480℃, pressure 0.1–5.0 MPa, hydrogen-to-carbon ratio of reducing gas 0.5–30, inlet linear velocity of reducing gas 0.05–0.9 m / s, catalyst concentration in the dense phase zone of the fluidized bed 5–50%, and reduction and activation time 2–48 hours. This technology is suitable for the activation of general catalysts, i.e., the formation of mixed-phase active components.

[0009] CN100404137C discloses an industrial reduction method for granular iron-based Fischer-Tropsch synthesis catalysts. The reduction process is divided into a reduction stage and a conditioning stage. The reduction stage corresponds to the heating stage, and the conditioning stage corresponds to the exothermic reaction stage. The operating temperature is 260-450℃, the pressure is 1.5-5.0MPa, and the inlet linear velocity is 0.15-0.7m / s. However, this method is not specifically for the formation of pure-phase iron carbide.

[0010] CN104549559B discloses an activation method for an iron-based catalyst. This method divides the activation process into three main steps. The aim is to first convert iron oxide into a Fischer-Tropsch-active iron phase, and then use low-carbon olefins to further activate the catalyst, achieving the desired Fischer-Tropsch synthesis to produce olefins. The activation process consists of three steps: Step 1: At 0.01-5.0 MPa, an inert gas and / or hydrogen are used for the reaction, and the reactor temperature is raised to 230-480℃, held for 3-60 h. Step 2: While maintaining constant temperature and pressure, an inert gas and / or syngas are used for the reaction, held for 6-100 h. Step 3: The pressure is increased to 0.2-10 MPa, and an inert gas and / or a mixed gas of multi-carbon olefins are used for the reaction, held for 1-10 h, ending the reduction. However, this catalyst activation method primarily aims to produce low-carbon olefins as the reaction product, and the activation process selection is not specifically for the formation of pure-phase iron carbide.

[0011] CN 107149948A discloses an activation method for a gas-solid fluidized bed Fischer-Tropsch synthesis catalyst. To control the moisture content in the reactor, a dehydration tank is installed before the reactor. The activation heating process is divided into different stages: primarily maintaining a constant temperature of 120℃ for 2-8 hours, and then maintaining a constant temperature of 220-230℃ for 3-10 hours to ensure reaction stability. The reaction pressure is 0.5-2 MPa, the empty tower gas velocity is 0.04-0.12 m / s, and the final constant temperature is 260-280℃. The reducing gases are hydrogen and CO in a molar ratio of 40-200:1. However, this method is not specifically for the formation of pure-phase iron carbide; the activated catalyst prepared may contain iron oxide, elemental iron, and heterogeneous iron carbide simultaneously. Summary of the Invention

[0012] The purpose of this invention is to provide a continuous activation apparatus and method for forming a pure-phase ε / ε' iron carbide catalyst for Fischer-Tropsch synthesis, so as to achieve continuous activation of the pure-phase ε / ε' iron carbide catalyst for Fischer-Tropsch synthesis for industrial production.

[0013] To achieve one aspect of the above-mentioned objectives, the technical solution adopted by the present invention is as follows:

[0014] A continuous activation apparatus for forming a pure-phase ε / ε' iron carbide catalyst for Fischer-Tropsch synthesis, wherein the continuous activation apparatus comprises:

[0015] A catalyst feeder is configured to supply the Fischer-Tropsch iron-based catalyst to be activated into a hydrogen activation reactor;

[0016] The first fluidized bed reactor is configured to use hydrogen from a hydrogen source as a reducing gas to reduce iron oxide in the Fischer-Tropsch iron-based catalyst from the catalyst feed tank, and to discharge the reduced catalyst from the bottom and the reduced tail gas from the top.

[0017] The first storage tank is provided with a first hydrogen injection port connected to a hydrogen source, a reduction catalyst inlet connected to the first fluidized bed reactor, and a reduction catalyst outlet connected to the second fluidized bed reactor.

[0018] The second fluidized bed reactor is configured to use the first syngas from the first syngas pipe to perform carbonization pretreatment on the reduction catalyst from the first storage tank, and discharge the pretreated catalyst from the bottom and the pretreated tail gas from the top.

[0019] A first syngas pipe is connected at one end to the second fluidized bed reactor and at the other end to the third fluidized bed reactor. The first syngas pipe is equipped with a first dehydration tank for dehydrating the carbonization tail gas from the third fluidized bed reactor and sending it as the first syngas into the second fluidized bed. Optionally, the first syngas pipe is also equipped with a make-up gas pipe connected to a hydrogen source and / or a CO source for adjusting the molar ratio of hydrogen to CO in the first syngas.

[0020] The second storage tank is provided with a second hydrogen injection port connected to a hydrogen source, a pretreatment catalyst inlet connected to the second fluidized bed reactor, and a pretreatment catalyst outlet connected to the third fluidized bed reactor.

[0021] The third fluidized bed reactor is configured to use the second syngas from the second syngas pipe to carbonize the pretreatment catalyst from the second storage tank, and to discharge the pure phase ε / ε' iron carbide catalyst from the bottom and the carbonization tail gas from the top.

[0022] The second syngas pipe is connected at one end to the third fluidized bed reactor and at the other end to a hydrogen source and a CO source, and is used to receive hydrogen and CO as the second syngas and send them into the third fluidized bed.

[0023] The catalyst product tank is used to receive pure-phase ε / ε' iron carbide catalyst product from the third fluidized bed reactor.

[0024] According to the continuous activation apparatus of the present invention, preferably, the continuous activation apparatus further includes a pretreatment exhaust gas recycling unit, the pretreatment exhaust gas recycling unit comprising:

[0025] A first cooler is configured to cool the pretreatment exhaust gas from the second fluidized bed reactor to condense the water and oil therein;

[0026] The first gas-liquid separator is configured to perform gas-liquid separation on the pre-treatment exhaust gas cooled by the first cooler to remove condensate and oil therein.

[0027] The first compressor is configured to pressurize the pre-treatment exhaust gas after the condensate and oil have been separated by the first gas-liquid separator and send it into the second syngas pipe; and

[0028] The vent pipe is used to discharge a portion of the pretreatment exhaust gas after the condensate and oil have been separated by the first gas-liquid separator as vent gas.

[0029] According to the continuous activation apparatus of the present invention, preferably, the continuous activation apparatus further includes a reducing gas preheating unit, the reducing gas preheating unit comprising:

[0030] A first heat exchanger is configured to exchange heat and raise the temperature of the reducing gas entering the first fluidized bed reactor with the reducing tail gas discharged from the top of the first fluidized bed reactor; and

[0031] The heater is configured to further heat the reducing gas from the first heat exchanger and feed it into the first fluidized bed reactor.

[0032] According to the continuous activation apparatus of the present invention, preferably, the continuous activation apparatus further includes a reduction tail gas recycling unit, the reduction tail gas recycling unit comprising:

[0033] The second heat exchanger is configured to further exchange heat and cool down the reduction tail gas from the first heat exchanger with the second syngas to be introduced into the third fluidized bed reactor.

[0034] The second cooler is configured to cool the reduction exhaust gas from the second heat exchanger so that the water therein condenses.

[0035] The second gas-liquid separator is configured to perform gas-liquid separation on the reduction tail gas cooled by the second cooler to remove the condensate therein;

[0036] The second compressor is configured to pressurize the reduction exhaust gas after the condensate has been separated by the second gas-liquid separator; and

[0037] The second dehydration tank is configured to remove moisture from the reduction exhaust gas from the second compressor and reuse the dehydrated reduction exhaust gas as reduction gas.

[0038] According to the continuous activation apparatus of the present invention, preferably, the continuous activation apparatus further includes a third heat exchanger and an optional inlet gas detector, wherein the third heat exchanger is used to exchange heat and cool the carbonized tail gas to be entered into the first dehydration tank with the second syngas to be entered into the second heat exchanger; the inlet gas detector is used to detect the molar ratio of hydrogen to CO in the carbonized tail gas.

[0039] According to the continuous activation apparatus of the present invention, preferably, the first storage tank and the second storage tank are further provided with heat exchange components for regulating temperature;

[0040] The reactor bodies of the first, second, and third fluidized bed reactors include an upper straight section and a lower straight section connected by a tapering section, an upper end cap disposed at the upper end of the upper straight section, and a lower end cap disposed at the lower end of the lower straight section. A gas distributor is provided at the lower part of the lower straight section, thereby forming a fluidized zone above the gas distributor within the lower straight section. A gas-solid separator is provided at the upper part of the upper straight section, thereby forming a separation zone below the gas-solid separator within the upper straight section. An air inlet is provided below the gas distributor, and an exhaust outlet is provided above the gas-solid separator. A catalyst outlet is provided at the lower part of the fluidized zone, and a catalyst feed inlet is also provided above the catalyst outlet in the fluidized zone. Preferably, a heat exchange component is also provided in the fluidized zone for temperature regulation.

[0041] In another aspect of achieving the above-mentioned objective, the present invention also provides an activation method for forming a Fischer-Tropsch synthesis pure-phase ε / ε' iron carbide catalyst using the aforementioned continuous activation apparatus.

[0042] According to the activation method of the present invention, preferably, in the first fluidized bed reactor, the reactor pressure is 0.01-1.5 MPa, the temperature is 300-510 °C, and the gas hourly space velocity is 600-20000 Nm. 3 / h / t, the reduction reaction time is 0.5-10h;

[0043] According to the activation method of the present invention, preferably, in the second fluidized bed reactor, the hydrogen to CO molar ratio of the first synthesis gas introduced is 1.2-2.8:1, the reactor pressure is 0.005-0.7 MPa, the temperature is 90-180°C, and the gas hourly space velocity is 300-8000 Nm. 3 / h / t, pretreatment time is 15-90min;

[0044] According to the activation method of the present invention, preferably, in the third fluidized bed reactor, the hydrogen to CO molar ratio of the introduced second synthesis gas is 1-3:1, the reactor pressure is 0.01-1.0 MPa, the temperature is 200-280 °C, and the gas hourly space velocity is 500-20000 Nm. 3 / h / t, carbonization treatment time is 1.5-10h.

[0045] According to the activation method of the present invention, preferably, in the first and second fluidized bed reactors, the apparent gas velocity is controlled at 0.5-1.1 m / s, more preferably 0.6-0.8 m / s; according to the activation method of the present invention, preferably, in the third fluidized bed reactor, the apparent gas velocity is controlled at 0.03-0.4 m / s, more preferably 0.15-0.30 m / s.

[0046] According to the activation method of the present invention, preferably, when the catalyst is transferred in the first and second storage tanks, the catalyst in the storage tank is first pressurized with hydrogen, then the catalyst is controlled to be close to the temperature of the reactor to be transferred using the heat exchange components in the storage tank, then the pressure of the storage tank is kept higher than that of the reactor to be transferred, and then the required amount of catalyst is transported to the target reactor by gas delivery using the pressure difference.

[0047] Compared with the prior art, the present invention has the following advantages:

[0048] 1) Through the design of the activation reaction device and process flow, pure phase ε / ε' iron carbide can be produced in industrial plants, enabling the application of advanced pure phase iron carbide catalysts in industrial plants.

[0049] 2) The three reactors are distributed at different locations in the process. The heat carried by the gas exiting the highest temperature reactor can be used as the heat source for another reactor, which can improve the energy utilization efficiency of the unit and achieve the effects of saving costs, saving energy and reducing carbon emissions.

[0050] 3) This invention uses the tail gas from the third fluidized bed reactor as the inlet gas for the second fluidized bed reactor, making full use of the presence of byproduct gases such as low-carbon olefins generated in the third fluidized bed reactor. On the one hand, this helps to increase the gas density entering the second fluidized bed reactor, thereby stabilizing the fluidization state of the second fluidized bed reactor and preventing problems such as channeling and dead zones. This avoids the problem of excessively high local temperature of the catalyst, resulting in carbon deposition and sintering, which would reduce the catalyst activation effect. On the other hand, the appropriate amount of low-carbon olefins and other byproduct gases in the carbonization reaction tail gas have relatively high reactivity. Under the relatively mild reaction conditions in the pretreatment reactor, this is conducive to the uniform and stable formation of the carbonized phase layer on the catalyst surface and the refinement of the active phase grains to prevent aggregation. This ensures a good foundation for the pretreatment to facilitate the formation of pure-phase iron carbide catalyst in the subsequent carbonization reaction and is beneficial to the long-term stability of the catalyst in Fischer-Tropsch synthesis.

[0051] 4) The apparent gas velocity in the turbulent fluidized state of the first and second fluidized bed reactors is relatively high, which can promptly remove the water generated in the reaction and prevent the water from damaging the subsequent carbonization reaction; the apparent gas velocity in the bubbling fluidized state of the carbonization reaction is relatively low, which can greatly reduce catalyst wear and help reduce the reactor volume.

[0052] 5) By coordinating the three reactors, the gas-solid fluidization activation process can be transformed from the current intermittent operation to a continuous operation, eliminating the waiting time required for the reactor heating, cooling, and replacement processes of the existing single reactor, reducing the reactor idle rate, and improving production efficiency. Attached Figure Description

[0053] Figure 1 A flowchart illustrating one embodiment of the continuous activation apparatus of the present invention;

[0054] The markings in the diagram are explained as follows: 1-First heat exchanger; 2-Heater; 3-First fluidized bed reactor, also known as hydrogen reduction reactor; 4-Second heat exchanger; 5-Second cooler; 6-Second gas-liquid separator; 7-Second compressor; 8-Second dehydration tank; 9-Catalyst feed tank; 10-Third heat exchanger; 11-Third fluidized bed reactor, also known as carbonization reactor; 12-First dehydration tank; 13-Inlet gas detector; 14-Second fluidized bed reactor, also known as pretreatment reactor; 15-First cooler; 16-First gas-liquid separator; 17-First compressor; 18-Catalyst finished product tank; 19-First storage tank; 20-Second storage tank; 21-First syngas pipe; 22-Second syngas pipe; 23-Relaxation gas pipe. Detailed Implementation

[0055] The present invention will be further described below with reference to the embodiments and accompanying drawings. However, the present invention is not limited to the listed embodiments, but should also include equivalent improvements and modifications to the technical solutions defined in the appended claims of the present invention.

[0056] The catalyst raw material for the activation process of the present invention is an iron-based Fischer-Tropsch synthesis catalyst in which iron is in the oxidized state. For example, catalyst precursors prepared in CN112569982A, CN112569993A or CN112569987A can be used, in which the Fe component is in the oxidized state. The average particle size of the finished oxidized catalyst particles used as raw material is preferably 70-80 micrometers, which belongs to the Geldart A class particle range in the field of gas-solid fluidization. Preferably, the catalyst wear of the raw material is less than 4 g / h.

[0057] In one implementation, such as Figure 1 As shown, the continuous activation apparatus of the present invention includes:

[0058] Catalyst feed tank 9 is configured to provide the Fischer-Tropsch iron-based catalyst to be activated into the first fluidized bed reactor as a raw material;

[0059] The first fluidized bed reactor 3 is configured to use hydrogen from a hydrogen source as a reducing gas to reduce iron oxide in the Fischer-Tropsch iron-based catalyst from the catalyst feed tank, and discharge the reduced catalyst from the bottom and the reduced tail gas from the top.

[0060] The first storage tank 19 is provided with a first hydrogen injection port connected to a hydrogen source, a reduction catalyst inlet connected to the first fluidized bed reactor 3, and a reduction catalyst outlet connected to the second fluidized bed reactor 14, so as to receive the reduction catalyst from the first fluidized bed reactor 3 and send it to the second fluidized bed reactor 14 for processing when needed.

[0061] The second fluidized bed reactor 14 is configured to use the first syngas from the first syngas pipe 21 to perform carbonization pretreatment on the reduction catalyst from the first storage tank 19, and discharge the pretreated catalyst from the bottom and the pretreated tail gas from the top.

[0062] A first syngas pipe 21 is connected at one end to the second fluidized bed reactor 14 and at the other end to the third fluidized bed reactor 11. The first syngas pipe 21 is provided with a first dehydration tank 12, such as a dehydration tank filled with desiccant as is well known in the art, for dehydrating the carbonized tail gas from the third fluidized bed reactor 11 and sending it as the first syngas into the second fluidized bed 14. Optionally, the first syngas pipe 21 is also provided with a make-up gas pipe connected to a hydrogen source and / or a CO source for adjusting the molar ratio of hydrogen to CO in the first syngas.

[0063] The second storage tank 20 is provided with a second hydrogen injection port connected to a hydrogen source, a pretreatment catalyst inlet connected to the second fluidized bed reactor 14, and a pretreatment catalyst outlet connected to the third fluidized bed reactor 11, so as to receive the pretreatment catalyst from the second fluidized bed reactor 14 and send it to the third fluidized bed reactor 11 for processing when needed.

[0064] The third fluidized bed reactor 11 is configured to use the second syngas from the second syngas pipe 22 to carbonize the pretreatment catalyst from the second storage tank 20, and to discharge the pure phase ε / ε' iron carbide catalyst from the bottom and the carbonization tail gas from the top.

[0065] The second syngas pipe 22 is connected at one end to the third fluidized bed reactor 11 and at the other end to a hydrogen source and a CO source, and is used to receive hydrogen and CO as the second syngas and send them into the third fluidized bed 11.

[0066] Catalyst product tank 18 is used to receive pure phase ε / ε' iron carbide catalyst product from the third fluidized bed reactor 11.

[0067] In one embodiment, the continuous activation apparatus further includes a pretreatment tail gas recycling unit, which includes a first cooler 15, a first gas-liquid separator 16, a first compressor 17, and a venting pipe 23, wherein: the first cooler 15 is configured to cool the pretreatment tail gas from the second fluidized bed reactor 14 to condense the water and oil therein; studies have found that in the second and third fluidized bed reactors, due to the reduction or carbonization of the catalyst and the presence of hydrogen and CO, the pretreatment tail gas contains product water generated by reduction, carbonization, and Fischer-Tropsch synthesis, as well as small amounts of carbon dioxide, methane, and C2+. Alkane / olefins, etc., are cooled by the first cooler 15, which condenses water and some alkanes (oil), so that the remaining tail gas mainly contains hydrogen, CO and a small amount of by-product alkanes, olefins, carbon dioxide, etc.; the first gas-liquid separator 16 is configured to perform gas-liquid separation on the pre-treatment tail gas cooled by the first cooler 15 to remove condensate and oil; the first compressor 17 is configured to pressurize the pre-treatment tail gas after the condensate and oil have been separated by the first gas-liquid separator 16 and send it into the second synthesis gas pipe 22; the vent pipe 23 is used to discharge part of the pre-treatment tail gas after the condensate and oil have been separated by the first gas-liquid separator 16 as vent gas.

[0068] It is understood that by discharging part of the pretreatment tail gas, excessive accumulation of methane and CO2 as byproducts in the second syngas can be prevented, thereby avoiding adverse effects on the carbonization reaction in the third fluidized bed reactor and ensuring the acquisition of pure-phase ε / ε' iron carbide catalyst. For this purpose, preferably, the volume percentage of hydrogen and CO in the second syngas entering the third fluidized bed reactor is not less than 98%. However, it should be noted that the present invention does not pursue a 100% volume percentage of hydrogen and CO in the second syngas; for example, it can be 98.5%, 99%, or 99.5%.

[0069] In one embodiment, the continuous activation device further includes a reducing gas preheating unit, which includes a first heat exchanger 1 and a heater 2, wherein: the first heat exchanger 1 is configured to exchange heat and raise the temperature of the reducing gas to be introduced into the first fluidized bed reactor 3 with the reducing tail gas discharged from the top of the first fluidized bed reactor 3; the heater 2 is configured to further heat and raise the temperature of the reducing gas from the first heat exchanger 1 and send it into the first fluidized bed reactor 3.

[0070] In one embodiment, the continuous activation device further includes a reduction tail gas reuse unit, which includes a second heat exchanger 4, a second cooler 5, a second gas-liquid separator 6, a second compressor 7, and a second dehydration tank 8, wherein: the second heat exchanger 4 is configured to further exchange heat and cool the reduction tail gas from the first heat exchanger 1 with the second synthesis gas to be introduced into the third fluidized bed reactor 11; the second cooler 5 is configured to cool the reduction tail gas from the second heat exchanger 4 to condense the water therein; the second gas-liquid separator 6 is configured to perform gas-liquid separation on the reduction tail gas cooled by the second cooler 5 to remove the condensate therein; the second compressor 7 is configured to pressurize the reduction tail gas after the condensate has been separated by the second gas-liquid separator 6; and the second dehydration tank 8 is configured to remove moisture from the reduction tail gas from the second compressor 7 and reuse the dehydrated reduction tail gas as reducing gas.

[0071] In one embodiment, the continuous activation device further includes a third heat exchanger 10 and an optional inlet gas detector 13, wherein the third heat exchanger 10 is used to exchange heat and cool the carbonized tail gas to be entered into the first dehydration tank 12 with the second synthesis gas to be entered into the second heat exchanger 4; the inlet gas detector 13 is used to detect the molar ratio of hydrogen to CO in the carbonized tail gas for adjustment.

[0072] In one embodiment, the first and second storage tanks are further provided with heat exchange components, such as heat exchange coils or tubes, for temperature regulation.

[0073] In one embodiment, the reactor bodies of the first, second, and third fluidized bed reactors include an upper straight section and a lower straight section connected by a tapering section, an upper end cap disposed at the upper end of the upper straight section, and a lower end cap disposed at the lower end of the lower straight section; a gas distributor is provided at the lower part of the lower straight section, thereby forming a fluidized zone above the gas distributor in the lower straight section; a gas-solid separator, such as a cyclone separator or a filter, is provided at the upper part of the upper straight section to reduce tail gas entrainment, thereby forming a separation zone below the gas-solid separator in the upper straight section; an air inlet is provided below the gas distributor, and an exhaust port is provided above the gas-solid separator; a catalyst outlet is provided at the lower part of the fluidized zone, and a catalyst feed inlet is also provided above the catalyst outlet in the fluidized zone; preferably, a heat exchange component is also provided in the fluidized zone for temperature regulation.

[0074] Compared to laboratory preparation, the large-scale production of the industrial-scale equipment of this invention requires significantly different and even higher standards for the sufficient contact between the catalyst and the reactant gas. In this invention, a certain gas velocity needs to be maintained to ensure that the material in the fluidized zone reaches a gas-solid fluidized state. In one embodiment, the first fluidized bed reactor 3 and the second fluidized bed reactor 14 employ a gas-solid turbulent fluidized state reaction with an apparent gas velocity of 0.5-1.1 m / s, preferably 0.6-0.8 m / s, such as 0.7 m / s, so that products such as water and carbon dioxide can be discharged in time to avoid adverse effects on the reaction. In another embodiment, the third fluidized bed reactor 11 employs a bubbling fluidized state reaction with an apparent gas velocity of 0.03-0.4 m / s, preferably 0.15-0.30 m / s, such as 0.2 or 0.25 m / s, to better facilitate carbonization. Those skilled in the art can adjust parameters such as reactor diameter and height to meet the gas velocity requirements while ensuring the reactor gas hourly space velocity, which will not be elaborated here.

[0075] In one implementation, during operation, the gas circulation of the reduction tail gas reuse unit and the pretreatment tail gas reuse unit is first controlled to operate stably. Fresh hydrogen is mixed with the circulating gas (also called reduction tail gas) from the second dehydration tank 8 and used as reducing gas. After passing through the first heat exchanger 1 and the heater 2 in sequence, it reaches a predetermined temperature and enters the first fluidized bed reactor 3. The gas exiting the reactor is used as circulating gas. After passing through the first heat exchanger 1, the second heat exchanger 4, the second cooler 5, the second gas-liquid separator 6, the second compressor 7, and the second dehydration tank 8 in sequence, it is mixed with fresh hydrogen to become reducing gas. This cycle continues.

[0076] In one embodiment, during operation, the fresh gas from the pretreatment and carbonization reaction is hydrogen and carbon monoxide. It is mixed with the pretreatment tail gas after the separation of condensate and oil in the circulation and used as the second synthesis gas. It is first preheated by the third heat exchanger 10, and then heated by the second heat exchanger 4 before entering the third fluidized bed reactor 11. Of course, based on temperature regulation considerations, a portion of the gas from the third heat exchanger 10 can directly enter the third fluidized bed reactor without being heated by the second heat exchanger 4. The amount of this portion can be adjusted appropriately according to the temperature requirements. After carbonization treatment, the carbonization tail gas is discharged and enters the second fluidized bed reactor 14 after passing through the third heat exchanger 10 and the first dehydration tank 12. The pretreatment tail gas exiting the second fluidized bed reactor 14 enters the first cooler 15, and then enters the first gas-liquid separator 16. A portion of the gas exiting the first gas-liquid separator 16 is discharged as purge gas, and the remaining portion is used as circulating gas to enter the first compressor 17 for pressurization. It is then mixed with the fresh gas to form the second synthesis gas, which is used as the feed gas to the third fluidized bed reactor 11. This cycle continues. Studies have found that, particularly in the second fluidized bed reactor 14, using the tail gas from the third fluidized bed reactor as the inlet gas for the second fluidized bed reactor fully utilizes the presence of byproduct gases such as low-carbon olefins generated in the third fluidized bed reactor (for example, the volume percentage of (H2+CO) in the first synthesis gas can be 94%-97%, such as 95% or 96%). On the one hand, this helps to increase the gas density entering the second fluidized bed reactor, thereby stabilizing the fluidization state of the second fluidized bed reactor and preventing problems such as channeling and dead zones. This avoids the catalyst from being overheated locally, resulting in carbon buildup and sintering, which would reduce the catalyst activation effect. On the other hand, the appropriate amount of low-carbon olefins and other byproduct gases in the carbonization reaction tail gas, with their relatively high reactivity, under the relatively mild reaction conditions in the pretreatment reactor, is conducive to the uniform and stable formation of the carbonized phase layer on the catalyst surface and the refinement of the active phase grains to prevent aggregation. This ensures a good foundation for the pretreatment to facilitate the formation of pure-phase iron carbide catalyst in the subsequent carbonization reaction and is beneficial to the long-term stability of the catalyst in Fischer-Tropsch synthesis.

[0077] In one embodiment, hydrogen is used as the reducing gas in the first fluidized bed reactor. The reactor pressure is 0.01-1.5 MPa, such as 0.08 MPa, 0.1 MPa, 0.5 MPa, or 1 MPa, and the temperature is 300-510°C, such as 340°C, 360°C, 400°C, or 450°C. The gas hourly space velocity is controlled to be 600-20000 Nm depending on the amount of catalyst added. 3 / h / t, for example, 1000 Nm 3 / h / t, 5000Nm 3 / h / t、10000Nm 3 / h / t or 15000Nm 3 / h / t.

[0078] In one embodiment, the second and third fluidized bed reactors are in a syngas atmosphere dominated by (H2+CO). The hydrogen-to-carbon ratio (i.e., the molar ratio of hydrogen to CO, also known as the hydrogen-to-carbon molar ratio) of the second syngas entering the third fluidized bed reactor is 1-3:1, for example, 2:1, 2.5:1, or 2.8:1. The reactor pressure is 0.01-1.0 MPa, for example, 0.08 MPa, 0.1 MPa, 0.5 MPa, or 1 MPa, and the temperature is 200-280°C, for example, 220°C, 240°C, 250°C, or 260°C. The gas hourly space velocity is controlled at 500-20000 Nm based on the expected catalyst addition. 3 / h / t, for example, 1000 Nm 3 / h / t, 5000Nm 3 / h / t、10000Nm 3 / h / t or 15000Nm 3 / h / t; The hydrogen-to-carbon ratio in the first synthesis gas entering the second fluidized bed reactor is 1.2-2.8:1, for example, 2:1, 2.5:1, or 2.8:1; the reactor pressure is 0.005-0.7 MPa, for example, 0.08 MPa, 0.1 MPa, 0.4 MPa, or 0.6 MPa; the temperature is 90-180℃, for example, 100℃, 120℃, 140℃, or 160℃; and the gas hourly space velocity is controlled at 300-8000 Nm according to the expected catalyst addition. 3 / h / t For example, 800Nm 3 / h / t、1000Nm 3 / h / t, 4000Nm 3 / h / t or 6000Nm 3 / h / t.

[0079] In one embodiment, the catalyst in the catalyst feeding tank 9 is preheated to within ±40°C of the temperature of the first fluidized bed reactor and added to the first fluidized bed reactor by gas delivery. The reaction takes place in this reactor for 0.5-10 hours, such as 1, 3, 5 and 8 hours.

[0080] In one embodiment, the catalyst outlet of the first fluidized bed reactor 3 is connected to a first storage tank 19, and one end of the outlet of the first storage tank is connected to a second fluidized bed reactor 14. When the first storage tank 19 is not receiving catalyst, it maintains a hydrogen atmosphere and the pressure is lower than that of the first fluidized bed reactor 3.

[0081] In one embodiment, after the catalyst has reacted completely in the first fluidized bed reactor 3, it is transported to the first storage tank 19 by gas conveying using the pressure difference. Then, the catalyst is continuously fed into the first fluidized bed reactor 3 from the catalyst feeding tank 9 to carry out the second reduction reaction. The first fluidized bed reactor 3 is operated continuously in this way.

[0082] In one embodiment, before the catalyst is fed into the second fluidized bed reactor 14, hydrogen is used to pressurize the first storage tank 19. At the same time, the temperature of the catalyst is adjusted to a range close to the reaction temperature of the second fluidized bed, for example, ±10°C, using the heat exchange coil in the first storage tank 19. Then, the pressure of the first storage tank is kept higher than that of the second fluidized bed reactor 14. Using the pressure difference, the amount of catalyst required for the pretreatment reaction is transported to the second fluidized bed reactor 14 by gas delivery. The reaction is carried out in this reactor for 15-90 minutes, for example, 45, 60 or 75 minutes.

[0083] In one embodiment, the catalyst outlet of the second fluidized bed reactor 14 is connected to a second storage tank 20, and one end of the outlet of the second storage tank 20 is connected to a third fluidized bed reactor 11. When the second storage tank 20 is not receiving catalyst, it maintains a hydrogen atmosphere and the pressure is lower than that of the second fluidized bed reactor 14.

[0084] In one embodiment, after the catalyst has reacted completely in the second fluidized bed reactor 14, it is transported to the second storage tank 20 by gas conveying using the pressure difference. Then, the catalyst continues to be fed from the first storage tank 19 into the second fluidized bed reactor 14 for a second pretreatment reaction. The second fluidized bed reactor 14 operates continuously in this manner.

[0085] In one embodiment, the catalyst outlet of the second storage tank 20 is connected to the third fluidized bed reactor 11. Hydrogen is used to pressurize the catalyst in the second storage tank 20, and the catalyst is adjusted to a temperature close to that of the third fluidized bed reactor, for example, ±40°C or ±10°C, using the heat exchange coil inside the tank. Then, the pressure of the second storage tank 20 is kept higher than that of the third fluidized bed reactor 11. Using the pressure difference, the amount of catalyst required for the carbonization reaction is transported to the third fluidized bed reactor 11 by gas delivery, and the reaction is carried out in this reactor for 1.5-10 hours, for example, 2, 4, 6 and 8 hours.

[0086] In one embodiment, after the carbonization reaction is completed, the activated pure phase ε / ε' iron carbide catalyst is unloaded into the catalyst product tank 18 and stored under a nitrogen atmosphere, or directly added to the Fischer-Tropsch synthesis reactor for catalytic Fischer-Tropsch synthesis reaction.

[0087] In this invention, a certain gas velocity needs to be maintained to ensure that the material in the fluidized zone reaches a gas-solid fluidized state. In one embodiment, the first fluidized bed reactor 3 and the second fluidized bed reactor 14 adopt a gas-solid turbulent fluidized state reaction, with an apparent gas velocity of 0.5-1.1 m / s, preferably 0.6-0.8 m / s, such as 0.7 m / s; in another embodiment, the third fluidized bed reactor 11 adopts a bubbling fluidized state reaction, with an apparent gas velocity of 0.03-0.4 m / s, preferably 0.15-0.30 m / s, such as 0.2 or 0.25 m / s, to better facilitate the carbonization effect. Those skilled in the art can adjust parameters such as reactor diameter and height to meet the gas velocity requirements while ensuring the gas hourly space velocity in the reactor, which will not be elaborated here.

[0088] In one embodiment, the design capacity of the second fluidized bed reactor is equal to the capacity of the first fluidized bed reactor × (pretreatment reaction time / hydrogen reduction reaction time); the design capacity of the third fluidized bed reactor is equal to the capacity of the second fluidized bed reactor × (carbonization reaction time / pretreatment reaction time), so that the reactors can work together more effectively and optimize continuous production.

[0089] Example 1

[0090] According to step (1) of Example 1 of CN112569982A, the catalyst with an average particle size of 70-80 micrometers was used as the raw material catalyst (the same below). The catalyst was heated to 360°C in the catalyst feeding tank for later use.

[0091] like Figure 1 As shown, the atmosphere of the first fluidized bed reactor was first replaced with a hydrogen atmosphere, and the pressure was controlled at 0.1 MPa, the reactor temperature at 400℃, and the apparent gas velocity at 0.6 m / s. 1000 kg of catalyst was added to the reactor through a catalyst feeder, and the reaction was carried out at a constant temperature for 3 hours. The first fluidized bed reactor has an inner diameter of 2.6 m, a height of 5 m, and a designed capacity of 1000 kg.

[0092] Before the hydrogen reduction reaction ends, the pressure of the first storage tank is controlled at 0.05 MPa. After the hydrogen reduction reaction ends, 1000 kg of catalyst in the first fluidized bed reactor is transferred to the first storage tank.

[0093] After the transfer is completed, 1000 kg of catalyst is transferred from the catalyst feed tank to the first fluidized bed reactor for another reaction. After the reaction is completed in 3 hours, it is transferred to the first storage tank in the same way, and so on.

[0094] After the catalyst enters the first storage tank, it is cooled to 190°C through a heat exchange coil, while hydrogen is used to pressurize the first storage tank to 0.8 MPa.

[0095] During the above process, the pretreatment temperature was controlled at 180℃ and the pressure at 0.5MPa, and the system was kept in stable circulation. Then, 500kg of catalyst in the first storage tank was transferred to the second fluidized bed reactor for pretreatment reaction for 1.5h. The fluidized bed reactor had a fluidization zone diameter of 0.6m, a height of 6m, an apparent gas velocity of 0.6m / s, and a hydrogen-to-carbon molar ratio of 2:1.

[0096] Before the pretreatment reaction ends, the pressure in the second catalyst storage tank is controlled at 0.2 MPa. After the pretreatment reaction is completed, the catalyst is transferred into the second storage tank.

[0097] After the transfer is completed, the second fluidized bed reactor continues to transfer 500 kg of catalyst from the first storage tank to carry out the reaction again, and so on.

[0098] After the catalyst enters the second storage tank, it is heated to 220°C through a heat exchange coil and pressurized to 1.3 MPa using hydrogen.

[0099] During the above process, the temperature of the third fluidized bed reactor was controlled at 260℃, and the pressure was controlled at 1.0MPa, ensuring stable circulation of the system. At this time, 1000kg of catalyst was transferred from the second storage tank to the third fluidized bed reactor for a carbonization reaction of 3 hours. The fluidized zone of the third fluidized bed reactor had a diameter of 1.5m and a height of 6m, an apparent gas velocity of 0.2m / s, and a hydrogen-to-carbon molar ratio of 3:1.

[0100] During the carbonization reaction, the catalyst product tank was replaced with a nitrogen atmosphere, and the pressure was controlled at 0.7 MPa. After the carbonization reaction was completed, the activated catalyst was transferred to the catalyst product tank, and Mössbauer spectroscopy was performed using a Transmission 57Fe, 57Co(Rh) source sinusoidal velocity spectrometer. The obtained iron carbide was 100% pure active phase ε / ε' iron carbide, as shown in Table 1.

[0101] Following the steps above, after one batch of catalyst is reacted in the three reactors, the next batch of catalyst is added, forming a continuous activation process.

[0102] Example 2

[0103] like Figure 1 As shown, the catalyst is heated to 360°C in the feeding tank for later use.

[0104] First, the atmosphere in the first fluidized bed reactor was replaced with a hydrogen atmosphere, and the pressure was controlled at 0.1 MPa, the reactor temperature at 400℃, and the apparent gas velocity at 1 m / s. 1000 kg of catalyst was added to the reactor through a catalyst feeder, and the reaction was carried out at a constant temperature for 2 hours. The first fluidized bed reactor has an inner diameter of 2.6 m, a height of 5 m, and a designed capacity of 1000 kg.

[0105] Before the hydrogen reduction reaction ends, the pressure of the first storage tank is controlled at 0.05 MPa. After the hydrogen reduction reaction ends, 1000 kg of catalyst in the first fluidized bed reactor is transferred to the first storage tank.

[0106] After the transfer is completed, 1000 kg of catalyst is transferred from the catalyst feed tank to the first fluidized bed reactor for another reaction. After the reaction is completed in 2 hours, it is transferred to the first storage tank in the same way, and so on.

[0107] After the catalyst enters the first storage tank, it is cooled to 190°C through a heat exchange coil, while hydrogen is used to pressurize the first storage tank to 0.8 MPa.

[0108] During the above process, the pretreatment temperature was controlled at 180℃ and the pressure at 0.5MPa, and the system was kept running stably. Then, 500kg of catalyst was transferred from the first storage tank to the second fluidized bed reactor for pretreatment reaction for 1 hour. The fluidized bed reactor had a fluidization zone diameter of 0.6m, a height of 6m, an apparent gas velocity of 1m / s, and a hydrogen-to-carbon molar ratio of 2.5:1.

[0109] Before the pretreatment reaction ends, the pressure in the second catalyst storage tank is controlled at 0.2 MPa. After the pretreatment reaction is completed, the catalyst is transferred into the second storage tank.

[0110] After the transfer is completed, the second fluidized bed reactor continues to transfer 500 kg of catalyst from the first storage tank to carry out the reaction again, and so on.

[0111] After the catalyst enters the second storage tank, it is heated to 220°C through a heat exchange coil and pressurized to 1.3 MPa using hydrogen.

[0112] During the above process, the temperature of the third fluidized bed reactor was controlled at 260℃ and the pressure at 1.0MPa, and the system operated stably in a circulating manner. At this time, 1000kg of catalyst was transferred from the second storage tank to the third fluidized bed reactor for a carbonization reaction of 2 hours. The fluidized zone of the third fluidized bed reactor had a diameter of 1.5m and a height of 6m, an apparent gas velocity of 0.35m / s, and a hydrogen-to-carbon molar ratio of 2.8:1.

[0113] During the carbonization reaction, the catalyst product tank was replaced with a nitrogen atmosphere, and the pressure was controlled at 0.7 MPa. After the carbonization reaction was completed, the catalyst was transferred to the catalyst product tank and Mössbauer spectrometry was performed using a Transmission 57Fe, 57Co(Rh) source sinusoidal velocity spectrometer (see Table 1). The obtained iron carbide was 100% pure active phase ε / ε' iron carbide, as shown in Table 1.

[0114] Following the steps above, after one batch of catalyst is reacted in the three reactors, the next batch of catalyst is added, forming a continuous activation process.

[0115] Example 3

[0116] like Figure 1 As shown, the catalyst is heated to 390°C in the feeding tank for later use.

[0117] First, the atmosphere in the first fluidized bed reactor was replaced with a hydrogen atmosphere, and the pressure was controlled at 0.1 MPa, the reactor temperature at 420℃, and the apparent gas velocity at 0.5 m / s. 1000 kg of catalyst was added to the reactor through a catalyst feeder, and the reaction was carried out at a constant temperature for 2 hours. The first fluidized bed reactor has an inner diameter of 2.6 m, a height of 5 m, and a designed capacity of 1000 kg.

[0118] Before the hydrogen reduction reaction ends, the pressure of the first storage tank is controlled at 0.05 MPa. After the hydrogen reduction reaction ends, 1000 kg of catalyst in the first fluidized bed reactor is transferred to the first storage tank.

[0119] After the transfer is completed, 1000 kg of catalyst is transferred from the catalyst feed tank to the first fluidized bed reactor for another reaction. After the reaction is completed in 2 hours, it is transferred to the first storage tank in the same way, and so on.

[0120] After the catalyst enters the first storage tank, it is cooled to 190°C through a heat exchange coil, while hydrogen is used to pressurize the first storage tank to 0.8 MPa.

[0121] During the above process, the pretreatment temperature was controlled at 180℃ and the pressure at 0.5MPa, and the system was operated stably in a circulating manner. Then, 500kg of catalyst was transferred from the first storage tank to the second fluidized bed reactor for pretreatment reaction for 1 hour. The fluidized bed reactor had a fluidization zone diameter of 0.5m, a height of 6m, an apparent gas velocity of 0.5m / s, and a hydrogen-to-carbon molar ratio of 2.5:1.

[0122] Before the pretreatment reaction ends, the pressure in the second catalyst storage tank is controlled at 0.2 MPa. After the pretreatment reaction is completed, the catalyst is transferred into the second storage tank.

[0123] After the transfer is completed, the second fluidized bed reactor continues to transfer 500 kg of catalyst from the first storage tank to carry out the reaction again, and so on.

[0124] After the catalyst enters the second storage tank, it is heated to 250°C through a heat exchange coil and pressurized to 1.3 MPa using hydrogen.

[0125] During the above process, the temperature of the third fluidized bed reactor was controlled at 250℃ and the pressure at 1.0MPa, and the system operated stably in a circulating manner. At this time, 1000kg of catalyst was transferred from the second storage tank to the third fluidized bed reactor for a carbonization reaction of 2 hours. The fluidized zone of the third fluidized bed reactor had a diameter of 1.5m and a height of 6m, an apparent gas velocity of 0.1m / s, and a hydrogen-to-carbon molar ratio of 2.8:1.

[0126] During the carbonization reaction, the catalyst product tank was replaced with a nitrogen atmosphere, and the pressure was controlled at 0.7 MPa. After the carbonization reaction was completed, the catalyst was transferred to the catalyst product tank and Mössbauer spectrometry was performed using a Transmission 57Fe, 57Co(Rh) source sinusoidal velocity spectrometer (see Table 1). The obtained iron carbide was 100% pure active phase ε / ε' iron carbide.

[0127] Following the steps above, after one batch of catalyst is reacted in the three reactors, the next batch of catalyst is added, forming a continuous activation process as shown in Table 1.

[0128] Comparative Example 1

[0129] Compared to Example 1, the apparent gas velocity in the reactor was adjusted as follows:

[0130] like Figure 1 As shown, the catalyst is heated to 360°C in the feeding tank for later use.

[0131] First, the atmosphere in the first fluidized bed reactor was replaced with a hydrogen atmosphere, and the pressure was controlled at 0.1 MPa, the reactor temperature at 400℃, and the apparent gas velocity at 0.3 m / s. 1000 kg of catalyst was added to the reactor through a catalyst feeder, and the reaction was carried out at a constant temperature for 3 hours. The first fluidized bed reactor has an inner diameter of 2.6 m, a height of 5 m, and a designed capacity of 1000 kg.

[0132] Before the hydrogen reduction reaction ends, the pressure in the first storage tank is controlled at 0.05 MPa. After the hydrogen reduction reaction ends, 1000 kg of catalyst is transferred to the first storage tank.

[0133] After the transfer is completed, 1000 kg of catalyst is transferred from the catalyst feed tank to the first fluidized bed reactor for another reaction. After the reaction is completed in 3 hours, it is transferred to the first storage tank in the same way, and so on.

[0134] After the catalyst enters the first storage tank, it is cooled to 190°C through a heat exchange coil, while hydrogen is used to pressurize the first storage tank to 0.8 MPa.

[0135] During the above process, the pretreatment temperature was controlled at 180℃ and the pressure at 0.5MPa, and the system was kept in stable circulation. Then, 500kg of catalyst in the first storage tank was transferred to the second fluidized bed reactor for pretreatment reaction for 1.5h. The fluidized bed reactor had a fluidization zone diameter of 0.6m, a height of 6m, an apparent gas velocity of 0.3m / s, and a hydrogen-to-carbon molar ratio of 2:1.

[0136] Before the pretreatment reaction ends, the pressure in the second catalyst storage tank is controlled at 0.2 MPa. After the pretreatment reaction is completed, the catalyst is transferred into the second storage tank.

[0137] After the transfer is completed, the second fluidized bed reactor continues to transfer 500 kg of catalyst from the first storage tank to carry out the reaction again, and so on.

[0138] After the catalyst enters the second storage tank, it is heated to 220°C through a heat exchange coil and pressurized to 1.3 MPa using hydrogen.

[0139] During the above process, the temperature of the third fluidized bed reactor was controlled at 260℃, and the pressure was controlled at 1.0MPa, ensuring stable circulation of the system. At this time, 1000kg of catalyst was transferred from the second storage tank to the third fluidized bed reactor for a carbonization reaction of 3 hours. The fluidized zone of the third fluidized bed reactor had a diameter of 1.5m and a height of 6m, an apparent gas velocity of 0.02m / s, and a hydrogen-to-carbon molar ratio of 3:1.

[0140] During the carbonization reaction, the catalyst product tank was replaced with a nitrogen atmosphere, and the pressure was controlled at 0.7 MPa. After the carbonization reaction was completed, the catalyst was transferred to the catalyst product tank and Mössbauer spectrometry was performed using a Transmission 57Fe, 57Co(Rh) source sinusoidal velocity spectrometer (see Table 1). The obtained iron carbide contained 85% of the active phase ε / ε' iron carbide.

[0141] Following the steps above, after one batch of catalyst is reacted in the three reactors, the next batch of catalyst is added, forming a continuous activation process.

[0142] Comparative Example 2

[0143] Compared with Example 2, the difference is that the first syngas (pure fresh gas, without carbonized tail gas) with a (H2+CO) ratio of 100% is exchanged with the carbonized tail gas and the temperature is adjusted before being sent to the second fluidized bed reactor. The carbonized tail gas is exchanged with the first syngas to cool down and then sent to the first cooler for further cooling before being sent to the first gas-liquid separator 16 to remove condensate and oil. The rest is the same.

[0144] The finished catalyst from the catalyst product tank was analyzed using a Mössbauer spectrometer (Transmission 57Fe, 57Co(Rh) source sinusoidal velocity spectrometer) (see Table 1). The obtained iron carbide contained 100% of the active phase ε / ε' iron carbide. Samples from the finished catalyst product tank after the entire system reaction was completed were subjected to Fischer-Tropsch synthesis for 500 hours to evaluate the long-term stability of the catalyst activity. The evaluation conditions were as follows:

[0145] The catalytic performance of the carbonized products from Example 2 and Comparative Example 2 was evaluated in a fixed-bed continuous reactor. The catalyst loading was 10.0 g. Evaluation conditions: T = 248 °C, P = 2.45 MPa, H2:CO = 1.8:1, (H2+CO) total = 42000 mL / h / g-Fe (standard state flow rate, relative to Fe element).

[0146] The results are shown in Table 2.

[0147] Table 1

[0148]

[0149]

[0150] Table 2

[0151]

[0152] All devices or components involved in this invention can be existing processing facilities, devices, or components with corresponding functions in the art, and will not be described in detail. Unless otherwise specified, all matters are understood or known to those skilled in the art based on their prior knowledge, and will not be described in detail. To highlight the concept of this invention, many necessary equipment for industrial applications, such as pumps, valves, and control components, are omitted from the figures.

[0153] Obviously, the above embodiments of the present invention are merely examples for clearly illustrating the present invention, and are not intended to limit the implementation of the present invention. Those skilled in the art can make other variations or modifications based on the above description. It is impossible to exhaustively list all embodiments here. All obvious variations or modifications derived from the technical solutions of the present invention are within the spirit and scope of the present invention.

Claims

1. A continuous activation apparatus for forming a pure-phase ε / ε' iron carbide catalyst for Fischer-Tropsch synthesis, characterized in that, The continuous activation device includes: A catalyst feeder is configured to supply the Fischer-Tropsch iron-based catalyst to be activated into a hydrogen activation reactor; The first fluidized bed reactor is configured to use hydrogen from a hydrogen source as a reducing gas to reduce iron oxide in the Fischer-Tropsch iron-based catalyst from the catalyst feed tank, and to discharge the reduced catalyst from the bottom and the reduced tail gas from the top. The first storage tank is provided with a first hydrogen injection port connected to a hydrogen source, a reduction catalyst inlet connected to the first fluidized bed reactor, and a reduction catalyst outlet connected to the second fluidized bed reactor. The second fluidized bed reactor is configured to use the first syngas from the first syngas pipe to perform carbonization pretreatment on the reduction catalyst from the first storage tank, and discharge the pretreated catalyst from the bottom and the pretreated tail gas from the top. The first syngas pipe is connected at one end to the second fluidized bed reactor and at the other end to the third fluidized bed reactor. The first syngas pipe is equipped with a first dehydration tank, which is used to dehydrate the carbonization tail gas of the third fluidized bed reactor and send it as the first syngas into the second fluidized bed reactor. The second storage tank is provided with a second hydrogen injection port connected to a hydrogen source, a pretreatment catalyst inlet connected to the second fluidized bed reactor, and a pretreatment catalyst outlet connected to the third fluidized bed reactor. The third fluidized bed reactor is configured to use the second syngas from the second syngas pipe to carbonize the pretreatment catalyst from the second storage tank, and to discharge the pure phase ε / ε' iron carbide catalyst from the bottom and the carbonization tail gas from the top. The second syngas pipe is connected at one end to the third fluidized bed reactor and at the other end to a hydrogen source and a CO source, and is used to receive hydrogen and CO as the second syngas and send them into the third fluidized bed reactor. The catalyst product tank is used to receive pure-phase ε / ε' iron carbide catalyst product from the third fluidized bed reactor.

2. The continuous activation device according to claim 1, characterized in that, The first synthesis gas pipe is also provided with a make-up gas pipe connected to a hydrogen source and / or a CO source, which is used to adjust the molar ratio of hydrogen to CO in the first synthesis gas.

3. The continuous activation apparatus according to claim 1, characterized in that, The continuous activation device further includes a pretreatment exhaust gas reuse unit, which includes: A first cooler is configured to cool the pretreatment exhaust gas from the second fluidized bed reactor to condense the water and oil therein; The first gas-liquid separator is configured to perform gas-liquid separation on the pre-treatment exhaust gas cooled by the first cooler to remove condensate and oil therein. The first compressor is configured to pressurize the pre-treatment exhaust gas after the condensate and oil have been separated by the first gas-liquid separator and send it into the second syngas pipe; and The vent pipe is used to discharge a portion of the pretreatment exhaust gas after the condensate and oil have been separated by the first gas-liquid separator as vent gas.

4. The continuous activation apparatus according to any one of claims 1-3, characterized in that, The continuous activation device further includes a reducing gas preheating unit, which comprises: A first heat exchanger is configured to exchange heat and raise the temperature of the reducing gas entering the first fluidized bed reactor with the reducing tail gas discharged from the top of the first fluidized bed reactor; and The heater is configured to further heat the reducing gas from the first heat exchanger and feed it into the first fluidized bed reactor.

5. The continuous activation apparatus according to claim 4, characterized in that, The continuous activation device further includes a reduction exhaust gas reuse unit, which includes: The second heat exchanger is configured to further exchange heat and cool down the reduction tail gas from the first heat exchanger with the second syngas to be introduced into the third fluidized bed reactor. The second cooler is configured to cool the reduction exhaust gas from the second heat exchanger so that the water therein condenses. The second gas-liquid separator is configured to perform gas-liquid separation on the reduction tail gas cooled by the second cooler to remove the condensate therein; The second compressor is configured to pressurize the reduction exhaust gas after the condensate has been separated by the second gas-liquid separator; and The second dehydration tank is configured to remove moisture from the reduction exhaust gas from the second compressor and reuse the dehydrated reduction exhaust gas as reduction gas.

6. The continuous activation apparatus according to any one of claims 1-3 and 5, characterized in that, The continuous activation device further includes a third heat exchanger, wherein the third heat exchanger is used to exchange heat and cool the carbonized tail gas to be entered into the first dehydration tank with the second synthesis gas to be entered into the second heat exchanger.

7. The continuous activation apparatus according to claim 4, characterized in that, The continuous activation device further includes a third heat exchanger and an inlet gas detector. The third heat exchanger is used to exchange heat and cool the carbonized tail gas to be introduced into the first dehydration tank with the second synthesis gas to be introduced into the second heat exchanger. The inlet gas detector is used to detect the molar ratio of hydrogen to CO in the carbonized tail gas.

8. The continuous activation apparatus according to any one of claims 1-3, 5 and 7, characterized in that, The first and second storage tanks are also equipped with heat exchange components for temperature regulation; The reactor bodies of the first, second, and third fluidized bed reactors include an upper straight section and a lower straight section connected by a tapering section, an upper end cap disposed at the upper end of the upper straight section, and a lower end cap disposed at the lower end of the lower straight section. A gas distributor is provided at the lower part of the lower straight section, thereby forming a fluidized zone above the gas distributor in the lower straight section. A gas-solid separator is provided at the upper part of the upper straight section, thereby forming a separation zone below the gas-solid separator in the upper straight section. An air inlet is provided below the gas distributor, and an exhaust outlet is provided above the gas-solid separator. A catalyst outlet is provided at the lower part of the fluidized zone, and a catalyst feed inlet is also provided above the catalyst outlet in the fluidized zone.

9. The continuous activation apparatus according to claim 6, characterized in that, The first and second storage tanks are also equipped with heat exchange components for temperature regulation; The reactor bodies of the first, second, and third fluidized bed reactors include an upper cylindrical section and a lower cylindrical section connected by a tapering section, an upper end cap disposed at the upper end of the upper cylindrical section, and a lower end cap disposed at the lower end of the lower cylindrical section. A gas distributor is provided at the lower part of the lower cylindrical section, thereby forming a fluidized zone above the gas distributor within the lower cylindrical section. A gas-solid separator is provided at the upper part of the upper cylindrical section, thereby forming a separation zone below the gas-solid separator within the upper cylindrical section. An air inlet is provided below the gas distributor, and an exhaust outlet is provided above the gas-solid separator. A catalyst outlet is provided at the lower part of the fluidized zone, and a catalyst feed inlet is also provided above the catalyst outlet in the fluidized zone. The fluidization zone is also equipped with a heat exchange component for temperature regulation.

10. An activation method for forming a Fischer-Tropsch synthesis pure-phase ε / ε' iron carbide catalyst using a continuous activation apparatus according to any one of claims 1-9.

11. The activation method according to claim 10, characterized in that, The first fluidized bed reactor has a reactor pressure of 0.01-1.5 MPa, a temperature of 300-510℃, a gas hourly space velocity of 600-20000 Nm 3 / h / t, and a reduction reaction time of 0.5-10 h. In the second fluidized bed reactor, the hydrogen and CO molar ratio of the introduced first synthesis gas is 1.2-2.8:1, the reactor pressure is 0.005-0.7 MPa, the temperature is 90-180℃, the gas hourly space velocity is 300-8000 Nm 3 / h / t, and the pretreatment time is 15-90 min. In the third fluidized bed reactor, the hydrogen and CO molar ratio of the introduced second synthesis gas is 1-3:1, the reactor pressure is 0.01-1.0 MPa, the temperature is 200-280℃, the gas hourly space velocity is 500-20000 Nm 3 / h / t, and the carbonization treatment time is 1.5-10 h.

12. The activation method according to claim 10 or 11, characterized in that, In the first and second fluidized bed reactors, the apparent gas velocity is controlled at 0.5-1.1 m / s; in the third fluidized bed reactor, the apparent gas velocity is controlled at 0.03-0.4 m / s.

13. The activation method according to claim 12, characterized in that, In the first and second fluidized bed reactors, the apparent gas velocity is controlled at 0.6-0.8 m / s; In the third fluidized bed reactor, the apparent gas velocity is controlled at 0.15-0.30 m / s.

14. The activation method according to any one of claims 10-11 and 13, characterized in that, During the transfer of catalyst between the first and second storage tanks, hydrogen is first used to pressurize the catalyst in the storage tank. Then, the heat exchange components in the storage tank are used to control the catalyst to approach the temperature of the reactor to be transferred to. After that, the pressure in the storage tank is kept higher than that in the reactor to be transferred to. Then, the required amount of catalyst is delivered to the target reactor by gas delivery using the pressure difference.

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