A sulfur recovery process for acid gas

Abstract

The present invention belongs to the technical field of sulfur recovery, and specifically relates to a sulfur recovery process of acidic gas, wherein the equipment of the sulfur production unit includes a combustion furnace, a waste heat boiler, a primary condenser, an air flow mixer, a primary conversion reactor, a secondary condenser, a first reheater, a secondary conversion reactor and a tertiary condenser along the process gas flow direction; the outlet process gas flow temperature of the primary condenser is controlled to be 150-170°C; a low-temperature oxidation reactor filled with a low-temperature oxidation catalyst is arranged between the air flow mixer and the primary conversion reactor, or a low-temperature oxidation catalyst layer is arranged at the air flow inlet reaction section of the catalyst bed of the primary conversion reactor; the process gas flow discharged from the primary condenser is mixed with air or oxygen-enriched air in the air flow mixer, and then treated by the low-temperature oxidation reactor or the low-temperature oxidation catalyst layer, and the gas flow temperature is increased to 210-260°C. In this process, the control of the combustion furnace and the primary conversion reactor is simplified, and the operating conditions and operation effects of the sulfur production unit are easier to ensure.

Description

Technical Field

[0001] The invention belongs to the technical field of sulfur recovery, and in particular relates to a sulfur recovery process of acid gas. Background Art

[0002] The sulfur recovery device for acid gas includes a sulfur production unit and a tail gas treatment unit. At present, the sulfur production unit of the mainstream technology has basic equipment along the acid gas flow direction, usually including a combustion furnace, a waste heat boiler, a primary condenser, a two-stage or three-stage conversion reactor (Claus reactor), which converts more than 90% of the hydrogen sulfide contained in the acid gas into sulfur for recovery. The first reheater and the second condenser are set before and after the first conversion reactor, and the second reheater and the third condenser are set before and after the second conversion reactor. If there is a three-stage conversion reactor, the third reheater and the fourth condenser are set before and after it; the liquid sulfur captured by each condenser flows into the liquid sulfur tank for storage; each stage of the conversion reactor operates above the dew point temperature of elemental sulfur, and the main reaction is that H2S and SO2 generate elemental sulfur and release heat.

[0003] In production practice, the operating conditions of the primary conversion reactor have a great impact on the overall operation effect of the sulfur recovery unit, especially on whether the exhaust air of the tail gas treatment unit can meet the emission standards. In addition to further generating elemental sulfur, the primary conversion reactor also uses the higher bed temperature conditions caused by the large adiabatic reaction temperature rise to hydrolyze COS and CS2 generated during the acid gas combustion process to basically convert them into H2S. The total content of H2S and SO2 in the outlet process gas of the primary condenser is usually 8-15v%, and the total content of COS and CS2 is often as high as 0.5-0.8v%. In the primary conversion reactor, the better catalyst loading mode is to load the upper part with a deoxygenation protection type sulfur recovery catalyst with oxygen leakage processing performance (such as a sulfur recovery catalyst containing Fe2O3), and the lower part with a better TiO2 catalyst with COS and CS2 hydrolysis performance and anti-sulfation performance. The lower part of the reactor needs to reach a bed temperature of more than 300℃ and around 330℃ when necessary to achieve the conversion rate requirements of COS and CS2. However, the catalyst loading mode of the prior art usually requires a process gas inlet temperature of more than 210°C.

[0004] The operating conditions and operation effects of the primary conversion reactor depend not only on the loaded catalyst and its overall conversion level, but also on the heating method of the first reheater and the temperature level that the outlet process gas can reach. The function of the first reheater is to increase the temperature of the outlet process gas of the primary condenser from 145-160°C to the temperature of 210-260°C usually required at the inlet of the primary conversion reactor. The reheating method is generally high-temperature hot blending, online heating furnace or process gas heat exchange. Among them, the high-temperature hot blending method is controlled by the high-temperature blending valve set on the top of the combustion furnace, so that the required flow rate and the post-combustion gas flow of the acid gas with a temperature above 1100°C (unseparated elemental sulfur generated by the combustion of acid gas) flow to the first reheater through the pipeline and mix with the process gas from the primary condenser; the online heating furnace method is to use fuel gas (such as light hydrocarbons) to match the air combustion, and the obtained high-temperature gas flow is mixed with the process gas from the primary condenser; the process gas heat exchange method is to use the outlet process gas (temperature of about 300°C) of the primary conversion reactor as the heat source, and heat the process gas from the primary condenser through indirect heat exchange.

[0005] Since the operating conditions and effects of the combustion furnace, the first reheater and the primary conversion reactor are closely related, when the flow rate and composition of the acid gas fluctuate greatly or rapidly (which is the norm in sulfur recovery units), it is difficult to operate the three devices. This is especially obvious after the sulfur recovery unit has been in operation for a long time and the performance of key control components has decayed. Recent studies have found that if the first reheater can be omitted, the top layer of the primary conversion reactor (airflow enters from the top and exits from the bottom) is filled or equipped with a suitable catalyst, and the outlet process gas (temperature 145-160°C) of the primary condenser directly enters the primary conversion reactor, the control of the combustion furnace and the primary conversion reactor may be simplified, and the operating conditions and operating results may be easier to ensure; accordingly, it is believed that the main technical requirements for the top catalyst layer in the primary conversion reactor are that the inlet temperature is 150-160°C, and the temperature of the bottom of the catalyst layer and the exhaust gas flow can be increased to above 210°C, preferably above 220°C, through the reaction temperature rise, especially the adiabatic temperature rise of the O2 contained in the outlet process gas of the primary condenser and the added O2 reacting with H2S, and this effect can be maintained for a long time. However, the catalysts used in the sulfur recovery field, including the catalysts disclosed in the following CN1136046C, CN109126831A, and CN1031194A, cannot achieve these technical effects; among them, the catalysts of CN1136046C and CN109126831A have low catalytic activity for the reaction of H2S and O2 to generate elemental sulfur under the temperature condition of 145-180°C, which cannot meet industrial requirements, and the reaction activity of the catalysts of CN1136046C and CN1031194A is not easy to maintain.

[0006] CN1136046C provides a dual-function sulfur recovery catalyst with deoxidation protection ability and high Claus activity. The catalyst is a common type of primary Claus reactor and can be loaded on the upper part of the primary Claus reactor or loaded on the entire bed. The operating temperature is 200-320°C and the deoxidation ability can reach 20,000ppm (2v%, removed by reaction with H2S to generate elemental sulfur). The catalyst uses activated alumina as a carrier, an iron-containing compound as an active component, and V2O5 as a co-catalyst. The content of the iron compound is 1-15% of the catalyst weight, preferably 3-10%, calculated as Fe2O3, and the content of vanadium is 0.2-2.0% of the catalyst weight, preferably 0.5-1.5% calculated as V2O5. The iron-containing compound includes FeSO4 and / or Fe2(SO4)3 and / or Fe2O3; the specific surface area of ​​the activated alumina carrier is 250-400m 2 / g320~400m is the best 2 / g, pore volume is 0.3-0.6ml / g, preferably 0.4-0.5ml / g, the crystal form of activated alumina is preferably γ and / or ρ type, and may also contain a small amount of pseudo-boehmite. The catalyst can be prepared by a traditional impregnation method. When NH4VO3 is used as a precursor of the co-catalyst, it can form a complex with oxalic acid in an aqueous solution and then loaded on a carrier. The molar ratio of oxalic acid to NH4VO3 is 1-3.

[0007] CN109126831A provides a catalyst for selectively oxidizing H2S to generate elemental sulfur, with silicon dioxide as a carrier and iron oxide as an active component (Fe2O3 content 5-20wt%); the specific surface area of ​​the carrier is 50-150m 2 / g, pore volume 0.4-0.7mL / g, and the most probable pore diameter 15-35nm. This catalyst is a common type of catalyst used in the selective oxidation reactor of the tail gas treatment unit. The operating temperature is 190-230℃. With the added O2, 0.5-1v% of H2S is basically converted into elemental sulfur.

[0008] CN1031194A provides a low-temperature Claus catalyst, the preparation method of which is as follows: drying activated alumina (φ2-8mm) at 100-150°C, impregnating with ferric nitrate or ferrous sulfate solution, drying, treating with ammonium carbonate or ammonia solution, washing with water, drying, and then activating and cooling at 200-350°C to obtain a low-temperature Claus catalyst; wherein the activated alumina pellets used have a double-channel structure, and the crystal form is γ-Al2O3 or / and χ-Al2O3 or / and boehmite, preferably γ-Al2O3; the iron compound loading is preferably 2-8% in terms of Fe2O3. The catalyst is highly oxidized at high oxygen content (0.714-2.86 g / m 3 ) Raw gas and high space velocity (1000~5000h -1) conditions, it has good low-temperature Claus conversion activity, sulfur capacity, anti-sulfation performance and activity stability. The catalyst is used in a low-temperature Claus reactor of a non-mainstream technology for a sulfur production unit. The reactor is usually used to treat process gas with a total content of H2S and SO2 of less than 3v%, preferably less than 1.5v%, and a small flow rate. It is generally used as the final Claus reactor of a small-scale sulfur production unit. Its operation process switches between a reaction conversion mode and a thermal regeneration mode; in the reaction conversion mode stage, the catalyst bed temperature is about 130°C, and it operates below the dew point temperature of elemental sulfur. H2S and SO2 are converted into elemental sulfur and adsorbed in the inner pores of the catalyst in the form of liquid sulfur. The reaction heat needs to be removed to maintain the catalyst bed temperature, and the reactor switches to the thermal regeneration mode when the H2S and SO2 conversion capacity of the reactor is reduced to a certain extent; in the thermal regeneration mode stage, the catalyst bed temperature rises to about 300°C, for example, and a purge gas is introduced to gasify and remove the liquid sulfur in the inner pores of the catalyst, and then the temperature is reduced to switch to the reaction conversion mode.

[0009] Therefore, it is necessary to develop a sulfur recovery process for acid gas, in which the top layer of the catalyst filled in the primary conversion reactor can utilize the reaction temperature rise to increase the gas flow temperature to above 210°C or even above 220°C for a long time. The 145-160°C process gas at the outlet of the primary condenser after the sulfur production unit combustion furnace and the waste heat boiler enters the primary conversion reactor directly or after simple gas distribution without reheating, thereby simplifying the control of the combustion furnace and the primary conversion reactor, and the operating conditions and operating results are easier to guarantee. Summary of the invention

[0010] In order to solve the above technical problems, the present invention provides a sulfur recovery process for acid gas, wherein the equipment of the sulfur making unit comprises a combustion furnace, a waste heat boiler, a primary condenser, an air flow mixer, a primary conversion reactor, a secondary condenser, a first reheater, a secondary conversion reactor and a tertiary condenser along the process gas flow direction; the liquid sulfur captured by the primary to tertiary condensers flows into a liquid sulfur tank for storage, and the outlet process gas flow temperature of the primary condenser is controlled to be 150-170°C; a low-temperature oxidation reactor filled with a low-temperature oxidation catalyst is arranged between the air flow mixer and the primary conversion reactor, or a low-temperature oxidation catalyst layer is arranged at the air flow inlet reaction section (i.e., the upper section of the catalyst bed) of the catalyst bed of the primary conversion reactor; after the process gas flow discharged from the primary condenser is mixed with air or oxygen-enriched air of required flow rate and temperature in the air flow mixer, the gas flow temperature is increased to 210-260°C by the heat released by the reaction in the low-temperature oxidation reactor or the low-temperature oxidation catalyst layer.

[0011] In the composition of the gas flow at the outlet of the gas flow mixer, the volume fractions of H2S, SO2 and O2 (respectively expressed in C H2S , C SO2 , C O2 Indicates, the same below) ratio, when controlled as (CH2S -2C SO2 ):C O2 =2:(0.95-1.3) and C H2S :(C SO2 +C O2 )=2:(0.9-1.1), the effect is better.

[0012] The volume fraction of H2S and SO2 in the process gas stream discharged from the first-stage condenser is preferably controlled to H2S +C SO2 8-16% and C H2S -2C SO2 It is 1.2-2.5v%, which is achieved by controlling the gas distribution conditions of the combustion furnace; at this time, the air or oxygen-enriched air with the required flow rate and preheated to 130-180°C can be added to the air flow mixer.

[0013] The low-temperature oxidation catalyst has a saturated water absorption rate of 0.6-0.7 mL / g and a specific surface area of ​​40-80 m 2 / g, modified alumina containing 10-12wt% of phosphorus oxide as P2O5 as carrier, 5-7wt% of vanadium oxide as V2O5, 4-6wt% of iron oxide as Fe2O3, and 2-4wt% of manganese oxide as MnO; prepared by the following method:

[0014] A. Pseudo-boehmite powder is calcined at 420-450°C for 2-4h, and then crushed to an average particle size (diameter) of 5-10μm. The obtained alumina powder is rolled into balls of Φ6-9mm in a rolling ball granulator with atomized spray and a phosphoric acid aqueous solution of a required concentration; the dried balls are calcined at 1000-1030°C for 2-3h to obtain modified alumina balls of Φ4-6mm; the weight ratio of the alumina powder to the phosphoric acid aqueous solution is 100:(45-50);

[0015] B. Place the modified alumina ball material in step A in a rotating drum, dynamically spray the vanadium oxalate aqueous solution of required concentration and volume, discharge the material after the surface of the ball becomes dry, seal and place for 5-10 hours, and dry at 110-130° C. to obtain a semi-finished ball material loaded with vanadium compound; the spray volume of the vanadium oxalate aqueous solution is 80-95% of the saturated water absorption volume of the modified alumina ball material;

[0016] C. placing the semi-finished ball material loaded with vanadium compound in step B in a rotary drum controlled at 80-90° C., dynamically spraying an aqueous ferrous acetate-manganese acetate solution of required concentration and volume at 80-90° C. until the surface of the ball becomes dry, discharging the material, sealing and standing for 15-30 hours, drying at 110-130° C. under ventilation conditions, and roasting at 400-430° C. under ventilation conditions for 4-6 hours to obtain a low-temperature oxidation catalyst; the spraying volume of the aqueous ferrous acetate-manganese acetate solution is 90-95% of the saturated water absorption volume of the semi-finished ball material loaded with vanadium compound.

[0017] In the sulfur recovery process of acid gas of the present invention, when a low-temperature oxidation reactor is not separately provided, in the catalyst bed of the primary conversion reactor, below the reaction section at the gas flow inlet of the low-temperature oxidation catalyst, a conventional sulfur recovery catalyst with deoxidation protection performance, such as a bifunctional sulfur recovery catalyst of the type described in CN1136046C, having deoxidation protection capability (making O2 react with H2S to generate elemental sulfur and remove it) and Claus activity, can be loaded, and then a sulfur recovery catalyst containing TiO2 can be loaded below the reaction section; below the gas flow inlet of the primary conversion reactor, a conventional sulfur recovery catalyst with COS / CS2 hydrolysis performance and deoxidation protection performance (trifunctional sulfur recovery catalyst) can also be loaded, and then a sulfur recovery catalyst containing TiO2 or an activated alumina sulfur recovery catalyst can be loaded below the reaction section. When a low-temperature oxidation reactor is set up separately, the catalyst bed of the primary conversion reactor is filled with the sulfur recovery catalyst with deoxidation protection performance or the sulfur recovery catalyst with COS / CS2 hydrolysis performance and deoxidation protection performance at the upper part, and the sulfur recovery catalyst containing TiO2 or the activated alumina sulfur recovery catalyst at the middle and lower part. The gas hourly space velocity of the low-temperature oxidation catalyst bed under the above-mentioned use conditions is 500-2000h -1 It has good reaction and heating effect at 1500-2000h -1 The desired reaction and temperature increase effects can be achieved at the same time, which is difficult to achieve with known conventional catalysts.

[0018] The sulfur recovery catalyst with deoxidation protection performance, the sulfur recovery catalyst with COS / CS2 hydrolysis performance and deoxidation protection performance, the sulfur recovery catalyst containing TiO2, and the activated alumina sulfur recovery catalyst are all commonly used catalysts in the field and in the primary conversion reactor, and have high and stable catalytic performance for the Claus reaction of H2S and SO2 to generate elemental sulfur under the corresponding bed reaction conditions. The sulfur recovery catalyst containing TiO2 includes a sulfur recovery catalyst containing TiO2 30-60wt%, the balance being γ and / or ρ type Al2O3, and a sulfur recovery catalyst containing TiO2 80-90wt%, the balance being CaSO4 binder. Among them, the sulfur recovery catalyst containing TiO2, especially the sulfur recovery catalyst containing TiO2 80-90wt% and the rest being CaSO4 binder, has high Claus activity and oxygen leakage resistance. In the bed section where the temperature reaches above 300°C, especially around 330°C (as the Claus reaction proceeds gradually, the reaction heat causes the catalyst and airflow temperature to gradually increase from top to bottom), it has high COS and CS2 hydrolysis performance and can basically hydrolyze and convert COS and CS2.

[0019] The mixing process of the process gas flow discharged from the primary condenser and the air flow and oxygen-enriched air flow can be carried out by a pipeline static mixer or a venturi mixer. The temperature of the air flow, oxygen-enriched air flow or acidic air flow can be 130-180°C. When the temperature is below 60°C, liquid or solid sulfur will be deposited, which will affect the effect of the mixer and the low-temperature oxidation catalyst bed.

[0020] In the production device using the sulfur recovery process of the present invention, during long-term operation, when the reaction temperature rise effect of the low-temperature oxidation catalyst layer in the low-temperature oxidation reactor or the primary conversion reactor drops to a certain level, a volume fraction composition of H2S 12-18%, SO2 3-5%, the remainder being inert gases such as N2, CO2, and water vapor, and an inlet gas flow with a temperature of 220-250° C. can be used at a gas hourly space velocity of 600-1200h -1 Under the appropriate conditions, the performance of the catalyst can be basically restored by 20-30 hours of reduction regeneration treatment.

[0021] Other conditions of the sulfur production unit include: in the catalyst bed of the secondary conversion reactor, the upper part can be filled with the sulfur recovery catalyst with deoxidation protection performance (such as CN1136046C type), and the middle and lower part can be filled with the TiO2-containing sulfur recovery catalyst or activated alumina sulfur recovery catalyst. Each reactor is an adiabatic reactor, and the flow direction of the process gas flow in the catalyst bed is from top to bottom. The reheating method of the first reheater includes an online heating furnace, process gas heat exchange or steam heat exchange, and the outlet process gas flow temperature is controlled to 200-220℃; when the process gas heat exchange method is adopted, the outlet process gas flow (temperature of about 300℃) of the primary conversion reactor is used as the heat source, and the process gas flow from the secondary condenser is heated by indirect and countercurrent heat exchange; the steam heat exchange method is to use the medium-pressure steam from the waste heat boiler and / or the pipeline network as the heat source for indirect heat exchange. The outlet process gas flow temperature of the secondary and tertiary condensers is controlled to be 145-160℃.

[0022] The sulfur recovery process of acid gas of the present invention can adopt various conventional methods in the prior art, including a combination of various suitable methods, as long as it is easy to control, can meet emission requirements and achieve economic indicators.

[0023] The beneficial effects of the acid gas sulfur recovery process of the present invention include: by arranging the gas flow mixer, the low-temperature oxidation reactor or the catalyst bed, the reaction temperature rise of H2S and O2 is utilized to increase the gas flow temperature to 210-260°C, so that the process gas flow at 145-160°C at the outlet of the primary condenser after the combustion furnace of the sulfur-making unit and the waste heat boiler enters the low-temperature oxidation catalyst reactor or the primary conversion reactor without reheating after simple gas distribution, so that the control of the combustion furnace and the primary conversion reactor is simplified, and the operating conditions and operation effects of the sulfur-making unit are easier to ensure as a whole, and the low-temperature oxidation catalyst can basically restore its performance through in-situ reduction regeneration and can have a service life of more than five years; the process has good application prospects, and is not only suitable for new equipment, but also suitable for technical transformation of existing acid gas sulfur recovery equipment.

[0024] From the effects of the following embodiments and comparative examples, it is believed that the main principles of the present invention are as follows, and the above beneficial effects are also the result of the following factors:

[0025] 1. The low temperature oxidation catalyst has a high catalytic activity for the reaction of H2S and O2 to generate elemental sulfur at a temperature of 150-260°C; and a very low catalytic activity for the Claus reaction of H2S and SO2 to generate elemental sulfur at a temperature of 150-210°C; at the same time, the catalyst has a minimum pore diameter of about 40 μm; the combination of these three factors enables the entire low temperature oxidation catalyst bed to operate at a temperature above the dew point of elemental sulfur, thereby substantially avoiding the condensation of liquid sulfur in the pores and surface of the catalyst, thereby ensuring a high temperature of up to 2000 h. -1 It also has good reaction and heating effects under the condition of gas hourly space velocity. The adiabatic temperature rise when every 1v% H2S in the gas flow reacts with 0.5v% O2 to generate elemental sulfur is about 60℃, while the adiabatic temperature rise when every 1v% H2S reacts with 0.5v% SO2 to generate elemental sulfur is less than 30℃ (the amount of elemental sulfur generated is 50% more); under the temperature condition of 150-210℃, if the low-temperature oxidation catalyst has Claus activity, it is not easy to ensure that the temperature of the catalyst particles in the bed is lower than the dew point of the elemental sulfur in the gas flow it treats (from top to bottom in the bed, as elemental sulfur is generated, the concentration and temperature of elemental sulfur increase at the same time). If elemental sulfur condenses in the inner pores of the catalyst, the bed cannot be heated at 1500-2000h. -1 The Aspen Plus software simulation study can verify these situations.

[0026] 2. The reason why the low-temperature oxidation catalyst has basically no Claus activity under the conditions of its use is that, on the one hand, the calcination temperature of the modified alumina balls is high, and they contain more phosphorus oxides, and there are basically no basic centers required for the Claus reaction on the inner and outer surfaces (phosphorus oxides also play a role in maintaining the pore volume of the carrier); on the other hand, the loaded active ingredient has a high content of vanadium oxides and basically no basic centers required for the Claus reaction.

[0027] 3. The reason why the low-temperature oxidation catalyst has a high catalytic activity for the reaction of H2S and O2 to generate elemental sulfur under its use conditions should be due to the loaded vanadium iron manganese composite oxide and its content composition and preparation method that are different from the prior art; in particular, only drying is performed between the two sprayings (step B modified alumina balls sprayed with vanadium oxalate aqueous solution and sealed and allowed to stand for homogenization treatment), and during the placement process after the second spraying (step C sprayed with ferrous acetate-manganese acetate aqueous solution and sealed and allowed to stand for homogenization treatment for 15-30 hours), vanadium oxalate will gradually dissolve and slowly react with ferrous acetate and manganese acetate in the solution, and the generated insoluble intermediates are evenly distributed on the surface of the inner pores of the catalyst, and are converted into the vanadium iron manganese composite oxide with high dispersion and high activity after calcination. When the preparation method is different, the catalytic activity is significantly different.

[0028] 4. The catalysts of the prior art CN1136046C and CN1031194A have low catalytic activity for the reaction of H2S and O2 to generate elemental sulfur when treating the process gas at the inlet of the low-temperature oxidation reactor at a temperature of 145-180°C, and the reaction activity is difficult to maintain. The reason is that the catalyst has a certain Claus, and the dew point of the generated elemental sulfur is easy to exceed the temperature of the catalyst particles and the surrounding airflow, so it gradually condenses into liquid sulfur in the pores of the catalyst. DETAILED DESCRIPTION

[0029] The present invention is described in detail below with reference to the embodiments, but is not intended to limit the present invention.

[0030] Example 1

[0031] The sulfur recovery process of acid gas in this embodiment 1 includes a sulfur production unit and a deep purification unit; the equipment of the sulfur production unit includes a combustion furnace, a waste heat boiler, a primary condenser, an air flow mixer, a low-temperature oxidation reactor, a primary conversion reactor, a secondary condenser, a first reheater, a secondary conversion reactor and a tertiary condenser along the process gas flow direction; the liquid sulfur captured by the primary to tertiary condensers flows into a liquid sulfur tank for storage, and the outlet process gas flow temperature of the primary condenser is controlled to be 150-170°C; wherein, the low-temperature oxidation reactor is filled with a low-temperature oxidation catalyst (scaled production according to the method of Preparation Example 1), the process gas flow discharged from the primary condenser is mixed with air of a required flow rate and a temperature of 130-180°C in the air flow mixer, and then enters the low-temperature oxidation reactor and then enters the primary conversion reactor; in the composition of the gas flow at the outlet of the gas flow mixer, the volume fractions of H2S, SO2, and O2 (respectively expressed in C H2S , C SO2 , C O2 The proportional relationship between them is expressed as (C H2S -2C SO2 ):C O2 =2:(0.95-1.3) and C H2S :(C SO2 +C O2 )=2:(0.9-1.1); H2S and O2 react in the catalyst bed of the low-temperature oxidation reactor to generate gaseous elemental sulfur and release heat, and the outlet gas flow temperature is 210-240°C;

[0032] In the catalyst bed of the primary conversion reactor, the upper part is filled with sulfur recovery catalyst A958 (our company's product, 1 / 3 of the bed height, belonging to the type described in CN1136046C) with deoxidation protection performance, and the lower middle part is filled with sulfur recovery catalyst A988 (our company's product, 2 / 3 of the bed height) containing TiO2 88wt% and the remainder being CaSO4 binder; in the catalyst bed of the secondary conversion reactor, the upper part is filled with sulfur recovery catalyst A958 (our company's product, 1 / 3 of the bed height) with deoxidation protection performance, and the lower middle part is filled with activated alumina sulfur recovery catalyst A918 (our company's product, 2 / 3 of the bed height);

[0033] The standard air flow velocity of the low temperature oxidation reactor is 1200h -1 , the upper limit is 2000h -1 ; The standard airflow velocity of the primary and secondary conversion reactors is 600h -1 , the upper limit is 1000h -1 ;

[0034] The air flow mixer is a Venturi mixer, and the pressure head (standard pressure 0.5MPa) of the air flow at a temperature of 130-180°C is used as the driving force; the outlet process air flow temperature of the secondary and tertiary condensers is controlled to be 145-160°C;

[0035] The first reheater adopts the process gas heat exchange mode, taking the outlet process gas flow (temperature about 300℃) of the first-stage conversion reactor as the heat source, and heats the process gas flow from the second-stage condenser to 210-220℃ through indirect, countercurrent heat exchange, and then enters the second-stage conversion reactor;

[0036] The main equipment of the deep purification unit includes, along the process gas flow direction, a second reheater (natural gas and air are burned for heat transfer and hydrogen is generated), a selective hydrogenation reactor (equipped with the company's selective hydrogenation catalyst A999g), a fourth-stage condenser, a third reheater, an air flow mixer (Venturi mixer, with air flow added to an O2 volume fraction of 0.5-1%), a selective oxidation reactor (equipped with the company's selective oxidation catalyst A888), a fifth-stage condenser, and more than two series-parallel and heat-regenerable adsorption desulfurization towers (equipped with the company's activated carbon desulfurizer Z803g), and the outlet gas flow of the adsorption desulfurization tower is discharged through the chimney; the inlet process gas flow temperature of the selective hydrogenation reactor and the selective oxidation reactor is controlled at 190-200°C, and the standard gas flow space velocity is 600h -1 , the upper limit is 1000h -1; The third reheater uses the outlet process gas flow (temperature of about 250°C) of the selective hydrogenation reactor as a heat source, and heats the process gas flow from the fourth-stage condenser to 190-200°C through indirect, countercurrent heat exchange, and then enters the selective oxidation reactor; the liquid sulfur captured by the fourth-stage and fifth-stage condensers flows into the liquid sulfur tank for storage, the outlet process gas flow temperature of the fourth-stage condenser is controlled to be 145-160°C, and the outlet process gas flow temperature of the fifth-stage condenser is controlled to be 10-50°C; the fifth-stage condenser has two condensation temperature zones of 145-160°C and 10-50°C, which discharge liquid sulfur and sulfur-containing condensed water respectively, the sulfur-containing condensed water is collected and settled, and the bottom sulfur slurry is regularly separated and returned to the 145-160°C condensation temperature zone of the fifth-stage condenser to recover liquid sulfur;

[0037] Each reactor is an adiabatic reactor, and the flow direction of the process gas flow in the catalyst bed is from top to bottom.

[0038] The sulfur recovery process of Example 1 is suitable for treating acid gas with a H2S volume fraction of 40-90%, and a single set of sulfur recovery capacity of 10,000-100,000 tons; at this time, in the process gas flow discharged from the primary condenser, C H2S +C SO2 Can be controlled to 10-16%, C H2S -2C SO2 The preferred control is 1.5-2.2v%, which can be achieved by controlling the gas distribution conditions of the combustion furnace; selecting the outlet process gas flow of the oxidation reactor, it is easy to reach C H2S+SO2 ≤1000mg / m 3 , C COS ≤3mg / m 3 , C CS2 ≤0.5mg / m 3 The level of chimney exhaust air flow can easily reach C H2S ≤1mg / m 3 , C SO2 ≤25 mg / m 3 , C COS+CS2 ≤2mg / m 3 level, meeting the requirements of ultra-low sulfur emission standards.

[0039] Example 2

[0040] The sulfur recovery process of acid gas in Example 2 is basically the same as the sulfur recovery process of acid gas in Example 1, and the main differences include the following aspects:

[0041] (1) In the air flow mixer, an oxygen-enriched air flow with an O2 volume fraction of 30-60% is added and mixed with the process air flow discharged from the primary condenser;

[0042] (2) Among the main equipment of the deep purification unit, the selective oxidation reactor is equipped with the selective oxidation catalyst A888g produced by our company, and its outlet process gas flow can be easily controlled to reach C H2S ≤80mg / m 3 , C SO2 500-1000mg / m 3 , C COS ≤1mg / m 3 , C CS2 ≤0.3mg / m 3 level; the adsorption desulfurization tower is replaced by a circulating liquid absorption tower, and the outlet air flow of the five-stage condenser is treated by the circulating liquid absorption tower and then discharged into the chimney; the circulating liquid absorption tower is atomized and sprayed with calcium-containing water slurry (solid content is less than 5wt%), and the calcium-containing water slurry maintains its absorption capacity by supplementing ultrafine calcium carbonate and maintains the total amount of slurry by introducing part of the condensed water (after sulfur separation) of the five-stage condenser, and the surplus slurry is sent to a sewage device for treatment or to a coal-fired boiler flue gas desulfurization device for utilization.

[0043] The sulfur recovery process of Example 2 is suitable for treating acid gas containing carbon dioxide with a H2S volume fraction of 15-30%, and a single set of sulfur recovery capacity of 10,000-30,000 tons; at this time, in the process gas flow discharged from the primary condenser, C H2S +C SO2 Can be controlled to 8-11%, C H2S -2C SO2 The preferred control is 1.2-1.8v%, which is achieved by controlling the gas distribution conditions of the combustion furnace; the exhaust air flow of the chimney can easily reach C H2S ≤1mg / m 3 , C SO2 ≤50mg / m 3 , C COS+CS2 ≤3mg / m 3 level, in line with low sulfur emission standards.

[0044] Example 3

[0045] The sulfur recovery process of acid gas in Example 3 is basically the same as the sulfur recovery process of acid gas in Example 2, with the main differences that: no low-temperature oxidation reactor is provided between the airflow mixer and the primary conversion reactor, and the upper, middle and lower layers of the primary conversion reactor are respectively filled with the low-temperature oxidation catalyst, sulfur recovery catalyst A958 and sulfur recovery catalyst A988 (with a height ratio of 2:1:2), and the standard airflow space velocity of the entire catalyst bed is 450h-1. -1 , upper limit is 750h -1 .

[0046] The sulfur recovery process of this embodiment 3 has the same application scenario and effect as that of embodiment 2.

[0047] The sulfur recovery process of acid gas described in the above Examples 1-3 is designed on the basis of the low-temperature oxidation catalyst prepared in Example 1 and its evaluation results, combined with the practical experience of the company's sulfur recovery process including the application experience of the other catalysts. Compared with the prior art, Examples 1 and 2 mainly add a low-temperature oxidation reactor, and Example 3 mainly adds a low-temperature oxidation catalyst layer on top of the conventional catalyst layer of the primary converter, and at the same time adopts an air flow mixer for injecting air or oxygen-enriched air, and correspondingly increases the H2S content in the process airflow discharged from the primary condenser (that is, reduces the amount of air or oxygen-enriched air added to the combustion furnace).

[0048] Using Aspen Plus software, the simulation study of the combustion furnace, low temperature oxidation reactor, and primary conversion reactor in the process of Examples 1-3 shows that the control of the combustion furnace and the primary conversion reactor can be greatly simplified, and the operating conditions and operation effects of the sulfur production unit are generally easier to achieve, which can ensure that the service life of the catalyst in the low temperature oxidation reactor and the primary conversion reactor is more than 5 years. Therefore, it has a good application prospect, and is suitable for the technical transformation of the existing acid gas sulfur recovery unit in addition to being suitable for the newly built acid gas sulfur recovery unit.

[0049] Preparation Example 1

[0050] The low temperature oxidation catalyst of this embodiment 1 was prepared by the following steps:

[0051] A. Pseudo-boehmite powder was calcined at 435°C for 3h, and then crushed to an average particle size (diameter) of 7-8μm. The obtained alumina powder was sprayed with a 36.5wt% phosphoric acid aqueous solution (the weight ratio of alumina powder to phosphoric acid aqueous solution was 100:47) in a rolling ball granulator, and rolled into Φ7-8mm balls; the balls were dried at 130°C for 10h, calcined at 1015°C for 2.5h, and cooled to obtain Φ4.7-5.5mm modified alumina balls; the test results of the obtained modified alumina balls included: saturated water absorption rate, i.e. pore volume, 0.66mL / g, specific surface area, 59 m 2 / g, phosphorus oxides as P2O5 10.9wt%;

[0052] B. Place the modified alumina balls (10.0 kg per batch, room temperature) in a rotating drum (with a heating jacket, which can be sealed with a transparent plastic cover), cover it with a transparent plastic cover, and dynamically spray 5.8 L of a 1.32 mol / L vanadium oxalate aqueous solution at room temperature (the spray volume is 88% of the saturated water absorption volume of the modified alumina balls) with a spraying time of 32 min. After spraying the solution, continue to rotate the drum for 20 min until the surface of the balls becomes dry, then discharge the balls, seal them and allow them to stand for 8 h (the cross-section of the balls has the same color), dry them at 120°C for 10 h, discharge them while they are hot, and sample 50 g of the semi-finished balls loaded with vanadium compounds, and immediately use the rest in step C as a spraying batch;

[0053] C. Place the semi-finished ball material loaded with vanadium compound in step B (about 100°C) in a rotating drum (heating sleeve temperature controlled at 85°C), cover it with a transparent plastic cover, and dynamically spray 4.8L of ferrous acetate-manganese acetate aqueous solution containing 1.52mol / L ferrous acetate and 1.07mol / L manganese acetate at 85°C (the spray volume is about 92.5% of the saturated water absorption volume of the semi-finished ball material) with a spraying time of 35min. After spraying the solution, the drum continues to rotate for 15min until the surface of the ball becomes dry, then discharge the material, keep it warm and seal it for 20h, dry it at 120°C under ventilation conditions, and roast it at 415°C under ventilation conditions for 5h to obtain a low-temperature oxidation catalyst.

[0054] The obtained low-temperature oxidation catalyst has a relatively consistent appearance and color of the cross section (including the hemispherical cross section); the sample was tested to contain 5.9wt% of vanadium oxide as V2O5, 5.0wt% of iron oxide as Fe2O3, and 3.1wt% of manganese oxide as MnO; the saturated water absorption rate, i.e., the pore volume, was 0.48mL / g, and the specific surface area was 68 m 2 / g, the most probable pore diameter is 42μm, the bulk density is 1.1kg / L, the crushing strength is 183N / particle, and the abrasion is 0.11wt%.

[0055] Among them, the ball was sampled after the heat-insulated, sealed and static treatment for 20 hours in step C, and the color of the cross-section (including the hemispherical cross-section) was detected to be relatively consistent. The ball was immersed in hot water, sealed and placed in an oven at 85°C for 3 hours. It was detected that the soaking water contained almost no vanadium, iron and manganese, indicating that vanadium oxalate, ferrous acetate and manganese acetate all reacted to generate water-insoluble intermediates during the heat-insulated, sealed and static treatment process described in step C.

[0056] Preparation Example 2

[0057] The low temperature oxidation catalyst of Example 2 was prepared substantially according to the method of steps BC of Preparation Example 1, with the following differences:

[0058] (1) In step B, a 1.58 mol / L vanadyl oxalate aqueous solution was used, and the spray volume remained unchanged at 5.8 L;

[0059] (2) The calcination temperature in step C was changed to 430°C.

[0060] The obtained low-temperature oxidation catalyst was sampled and tested to contain 7.0 wt% of vanadium oxide as V2O5, 4.9 wt% of iron oxide as Fe2O3, and 3.0 wt% of manganese oxide as MnO. The saturated water absorption rate, i.e., the pore volume, was 0.46 mL / g, and the specific surface area was 59 m 2 / g, the most probable pore diameter is 40μm.

[0061] Other effects are basically the same as those in Preparation Example 1.

[0062] Preparation Example 3

[0063] The low temperature oxidation catalyst of Example 3 was prepared substantially according to the method of steps BC of Preparation Example 1, with the following differences:

[0064] (1) In step C, a ferrous acetate-manganese acetate aqueous solution containing 1.75 mol / L ferrous acetate and 0.70 mol / L manganese acetate was used, and the spray volume remained unchanged at 4.8 L;

[0065] (2) The calcination temperature in step C was changed to 400°C.

[0066] The obtained low-temperature oxidation catalyst was sampled and tested to contain 6.0 wt% of vanadium oxide as V2O5, 5.8 wt% of iron oxide as Fe2O3, and 2.0 wt% of manganese oxide as MnO. The saturated water absorption rate, i.e., the pore volume, was 0.45 mL / g, and the specific surface area was 66 m 2 / g, and the most probable pore diameter is 38μm.

[0067] Other effects are basically the same as those in Preparation Example 1.

[0068] Preparation Example 4

[0069] The low temperature oxidation catalyst of Example 4 was prepared substantially according to the method of steps BC of Preparation Example 1, except that:

[0070] (1) In step C, a ferrous acetate-manganese acetate aqueous solution containing 1.33 mol / L ferrous acetate and 1.40 mol / L manganese acetate was used, and the spray volume remained unchanged at 4.8 L.

[0071] The obtained low-temperature oxidation catalyst was sampled and tested to contain 6.0 wt% of vanadium oxide as V2O5, 4.0 wt% of iron oxide as Fe2O3, and 4.0 wt% of manganese oxide as MnO. The saturated water absorption rate, i.e., the pore volume, was 0.44 mL / g, and the specific surface area was 62 m 2 / g, and the most probable pore diameter is 39μm.

[0072] Other effects are basically the same as those in Preparation Example 1.

[0073] Preparation Comparative Example 1

[0074] The catalyst of this comparative example 1 was prepared substantially according to the method of steps BC of preparation example 1, except that:

[0075] (1) In step B, the semi-finished ball material loaded with vanadium compound obtained by drying at 120°C for 10 h is first calcined at 430°C under ventilation conditions for 4 h, cooled to about 100°C and discharged while hot. 50 g of the sample is taken and the rest is immediately used in step C as a spraying batch.

[0076] After the heat preservation and airtight standing treatment for 20 hours in step C, the pellets were sampled, heated water was used to immerse the pellets, sealed and placed in an oven at 85°C for 3 hours, and the soaking water was tested to contain more iron and manganese, indicating that during the heat preservation and airtight standing treatment described in step C, ferrous acetate and manganese acetate could not completely react with vanadium oxide (obtained by roasting vanadium oxalate) to form water-insoluble intermediates. The color of the hemispherical cross section of the obtained catalyst was inconsistent along the radial direction, and the color of the core was slightly lighter than that of the outer edge.

[0077] Preparation Comparative Example 2

[0078] The catalyst of this comparative example 2 was prepared substantially according to the method of steps BC of preparation example 1, except that:

[0079] (1) The calcination process of the catalyst in step C was changed to calcination at 500°C for 4 h under ventilation conditions.

[0080] The obtained catalyst has a consistent appearance and cross-section color; the specific surface area of ​​the sample is 46 m 2 / g.

[0081] Preparation Comparative Example 3

[0082] The catalyst of this comparative example 3 was prepared substantially according to the method of steps BC of preparation example 1, except that:

[0083] (1) The vanadyl oxalate aqueous solution in step B is replaced by an aqueous solution containing 1.32 mol / L ammonium metavanadate and 2.64 mol / L oxalic acid.

[0084] Preparation Comparative Example 4

[0085] The catalyst of this comparative example 4 was prepared substantially according to the method of steps BC of preparation example 1, except that:

[0086] (1) The ferrous acetate-manganese acetate aqueous solution in step C is replaced with an aqueous solution containing 2.59 mol / L ferrous acetate.

[0087] Preparation Comparative Example 5

[0088] The catalyst of this comparative example 5 was prepared substantially according to the method of steps BC of preparation example 1, except that an aqueous solution of ferrous acetate-manganese acetate was sprayed in step B, and an aqueous solution of vanadyl oxalate was sprayed in step C:

[0089] B. Place the modified alumina ball material (10.0 kg) in step A in a drum, control the temperature to 85°C with a heating jacket, cover it with a transparent plastic cover, start the drum, and after the ball material is heated to 85°C, dynamically spray 6.1 L of ferrous acetate-manganese acetate aqueous solution containing 1.20 mol / L ferrous acetate and 0.82 mol / L manganese acetate at 85°C, and spray for 30 minutes. After spraying the solution, continue to rotate the drum for 18 minutes until the surface of the ball material becomes dry, then discharge the material, seal it and keep it warm for 8 hours (the cross-section of the ball material has the same color), dry it at 120°C for 10 hours, and discharge it while it is hot. Sample 50 g of the obtained semi-finished ball material loaded with vanadium compounds, and immediately use all the rest in step C as a spraying batch;

[0090] C. Place the semi-finished ball material (about 100°C) loaded with vanadium compound in step B in a rotating drum (heating jacket to control the temperature at 85°C), cover it tightly with a transparent plastic cover, and dynamically spray 4.5L of vanadium oxalate aqueous solution with a temperature of 85°C and a concentration of 1.70 mol / L for 32 minutes. After spraying the solution, the drum continues to rotate for 26 minutes until the surface of the ball becomes dry, then discharge the material, keep it warm and seal it for 20 hours, dry it at 120°C under ventilation conditions, and roast it at 415°C under ventilation conditions for 5 hours to obtain a low-temperature oxidation catalyst.

[0091] The obtained catalyst has a consistent appearance and color of cross section (including hemispherical cross section); the sampled catalyst contains 5.9wt% of vanadium oxide as V2O5, 5.1wt% of iron oxide as Fe2O3, and 3.0wt% of manganese oxide as MnO; the saturated water absorption rate, i.e., pore volume, is 0.48mL / g, and the specific surface area is 52m 2 / g.

[0092] Among them, the balls were sampled after being kept in a sealed and static state for 20 hours in step C, and the colors of the cross sections (including the hemispherical cross sections) were tested to be relatively consistent. The balls were immersed in hot water, sealed and placed in an oven at 85°C for 3 hours. The soaking water contained almost no iron or manganese, but more than 28% of vanadium was dissolved.

[0093] Catalyst Test Example 1

[0094] The catalysts obtained from the above Preparation Examples 1-4 and Comparative Examples 1-5 were sampled respectively, and the low-temperature oxidation reaction activities were evaluated in the same sulfur recovery catalyst activity evaluation device at the original particle size.

[0095] The evaluation device uses an insulated reaction tube (a vertically mounted cylindrical reaction tube made of polytetrafluoroethylene, with an inner diameter of 60 mm, a wall thickness of 0.55 mm, a height of 600 mm, and a jacket of 25 mm thick insulation cotton), the catalyst loading height is 400 mm (the bed volume is approximately 1.13 L), and the reaction gas flows from top to bottom. The reaction gas flow is first distributed in the Venturi mixer and then enters the reaction tube; the outlet gas flow of the reaction tube passes through the liquid sulfur condenser (heat transfer oil temperature control 150℃) to basically remove the elemental sulfur generated by the reaction, and most of it is pumped back to the Venturi mixer for recycling. The surplus gas is sent to the alkali liquid spray tower to absorb most of the sulfides and cool down before being treated in the activated carbon desulfurization tank and discharged; the power for the air circulation is the oxygen-nitrogen mixed gas flow with a pressure of more than 0.3MPa and preheated to 160-200℃ sprayed into the core tube of the Venturi mixer, and the circulating air flow pipeline in front of the Venturi mixer is distributed with room temperature H2S gas flow, SO2 gas flow and water vapor flow preheated to 160-200℃; the air flow pipeline is a stainless steel pipe with a 20mm thick insulation cotton jacket. The composition and temperature of the air flow at the inlet of the reaction tube are controlled by the amount of each air flow and the preheating temperature of the oxygen-nitrogen mixed air flow and the water vapor flow; before the sulfide-containing air flow is introduced, the catalyst bed and the circulation system are preheated with a circulating nitrogen flow at a temperature of 160-165°C until the temperature of the circulating air flow inlet of the Venturi mixer reaches 150°C.

[0096] The activity of each catalyst was evaluated under two gas flow conditions at the inlet of the reaction tube. Under both gas flow conditions, the evaluation was continued until the temperature of the gas flow at the outlet of the reaction tube reached a stable value or a maximum value, and then the evaluation was continued for 10 hours: (Gas flow condition 1) composition (volume fraction) was H2S 10%, SO2 4%, O2 1.1%, H2O 25%, the balance was N2, temperature 150-153°C, gas hourly space velocity 980-1020h -1 , (gas flow condition 2) composition (volume fraction) is H2S 10%, SO2 4%, O2 1.3%, H2O 25%, balance N2 and temperature 165-168℃, gas hourly space velocity 1950-2000h -1 .

[0097] During the evaluation of each catalyst, the composition of the outlet gas flow was detected every 30 minutes, and the temperature of the inlet and outlet gas flows of the reaction tube was continuously recorded on the computer of the control system.

[0098] The evaluation results under airflow condition 1 include: for the catalysts of Preparation Examples 1-4, the temperature of the airflow out of the reaction tube gradually increased to above 237°C and was stably maintained for 10 hours, among which the catalyst of Preparation Example 1 reached a stable state of 240°C. From the changes in the composition of the airflow inlet and outlet of the reaction tube, it can be seen that these four catalysts basically had no Claus activity during the evaluation process; for the catalysts of Preparation Comparative Examples 1, 3-5, the maximum temperature of the airflow out of the reaction tube was lower than 207°C; for the catalyst of Preparation Comparative Example 2, the maximum temperature of the airflow out of the reaction tube was 223°C, and then gradually decreased to 208°C.

[0099] The evaluation results under airflow condition 2 include: for the catalysts of Preparation Examples 1-4, the temperature of the airflow out of the reaction tube gradually increased to above 221°C and was stably maintained for 10 hours, among which the catalyst of Preparation Example 1 reached a stable state of 232°C. From the changes in the composition of the airflow inlet and outlet of the reaction tube, it can be seen that these four catalysts basically had no Claus activity during the evaluation process; for the catalysts of Preparation Comparative Examples 1, 3-5, the maximum temperature of the airflow out of the reaction tube was lower than 190°C; for the catalyst of Preparation Comparative Example 2, the maximum temperature of the airflow out of the reaction tube was 205°C, and then gradually decreased to 192°C.

[0100] Before the evaluation of each catalyst is completed, the H2S and SO2 flows are stopped in the circulating gas flow pipeline before the Venturi mixer, and the core tube oxygen-nitrogen mixed flow is replaced with a high-purity nitrogen flow at 140-150°C until the H2S content in the gas flow at the outlet of the reaction tube is less than 5 mg / m 3 When the reaction tube is heated, the heat insulation cotton on the outer jacket is removed, and the catalyst is unloaded after cooling down. The surface and inner layer of each unloaded catalyst are checked, and there is basically no sulfur deposition.

[0101] Prior to the evaluation of the above catalysts, the evaluation device also evaluated the low-temperature oxidation reaction activity of the sulfur recovery catalyst A958 (a product of our company, belonging to the CN1136046C type, also a Φ7-8mm ball material) with deoxidation protection performance described in Example 1, according to the gas flow condition 1. The result was that the temperature of the gas flow out of the reaction tube gradually increased to 172°C and then gradually decreased. From the changes in the gas flow composition at the inlet and outlet of the reaction tube, it can be seen that the Claus activity of the catalyst during the evaluation process is significantly higher than its catalytic activity for the reaction of H2S and O2 to generate elemental sulfur; after the catalyst was unloaded, it was found that there was a lot of sulfur deposition on the inner layer and surface of the catalyst.

[0102] Catalyst Test Example 2

[0103] The catalyst of Preparation Example 1, which has the best effect in Test Example 1, was unloaded and reinstalled into the reaction tube of the evaluation device, and the following steps were continuously performed:

[0104] (1) Perform 230 h of operation performance evaluation according to airflow condition 2 described in Test Example 1;

[0105] (2) Under inlet gas flow condition 3 (volume fraction composition: H2S 10%, SO2 4%, O2 1.5%, H2O 25%, balance N2 and temperature 230-240℃, gas hourly space velocity 960-1020h -1 ) was subjected to heat aging treatment, the temperature of the high-purity nitrogen flow in the core tube of the Venturi airflow mixer was changed to 330-340°C, and the temperature of the heat transfer oil in the liquid sulfur condenser was kept at 150°C) and operated for 60 hours;

[0106] (3) Conducting a 100-hour operating performance evaluation according to the airflow condition 2;

[0107] (4) Under inlet gas flow condition 4 (volume fraction composition: H2S 15%, SO2 4%, balance N2, temperature 230-240℃, gas hourly space velocity 960-1020h -1 ) for reduction and regeneration treatment, the oxygen-nitrogen mixed flow in the core tube of the Venturi air flow mixer is changed to a high-purity nitrogen flow at 240-250℃, and the liquid sulfur condenser heat transfer oil is controlled at 250℃ for 30h;

[0108] (5) Perform an operating performance evaluation for 200 h under airflow condition 2.

[0109] Results include:

[0110] During the 230-hour operation performance evaluation process in step (1), the temperature of the gas flow out of the reaction tube gradually increased and was stably maintained in the range of 231.5-232.2°C;

[0111] During the thermal aging process in step (2), the temperature of the gas flow out of the reaction tube gradually increased to 335°C, and then gradually decreased after being kept stable for 42 hours, indicating that the catalyst had experienced a certain degree of performance degradation;

[0112] During the 100-hour operation performance evaluation process in step (3), the temperature of the gas flow out of the reaction tube gradually decreased and was stably maintained in the range of 212.3-212.9°C;

[0113] During the reduction and regeneration process of step (4), the temperature of the gas flow out of the reaction tube gradually rises and remains in the range of 242-245°C;

[0114] During the 200-hour operation performance evaluation process in step (5), the temperature of the gas flow out of the reaction tube gradually decreased and remained stably within the range of 230.8-231.6°C;

[0115] During the evaluation of the operating performance of steps (1), (3) and (5), it can be seen from the changes in the composition of the gas flow at the inlet and outlet of the reaction tube that the catalyst has basically no Claus activity during the evaluation process.

[0116] In this test example 2, the thermal aging can simulate the operating performance of the catalyst after 1.5-2 years of industrial application. The main reason for the performance degradation is the sulfation of the surface of the active ingredient vanadium iron manganese composite oxide in the inner pores of the catalyst. The results of steps (3)-(5) show that the reduction and regeneration treatment of the catalyst after long-term operation and performance degradation can basically restore the performance of the catalyst. It is believed that the catalyst can have a service life of more than five years in the sulfur recovery process of acid gas in Examples 1-3.

[0117] The results of the above test examples 1 and 2 show that the prepared low-temperature oxidation catalyst has a high catalytic activity for the reaction of H2S and O2 to generate elemental sulfur under the temperature condition of 150-260°C; the catalytic activity for the Claus reaction of H2S and SO2 to generate elemental sulfur under the temperature condition of 150-210°C is very low; the entire low-temperature oxidation catalyst bed is operated at a temperature above the dew point of elemental sulfur, which basically avoids the condensation of liquid sulfur in the pores and surface of the catalyst, thereby ensuring the high temperature of up to 2000h -1 The reason why the low-temperature oxidation catalyst has high catalytic activity for the reaction of H2S and O2 to generate elemental sulfur under the conditions of its use should be due to the loaded vanadium-iron-manganese composite oxide and its content composition and preparation method that are different from the prior art.

Claims

1. A sulfur recovery process for acid gas, wherein the equipment of the sulfur production unit includes a combustion furnace, a waste heat boiler, a primary condenser, an air flow mixer, a primary conversion reactor, a secondary condenser, a first reheater, a secondary conversion reactor and a tertiary condenser along the process gas flow direction; the liquid sulfur captured by the primary to tertiary condensers flows into a liquid sulfur tank for storage, and the outlet process gas flow temperature of the primary condenser is controlled to be 150-170°C; characterized in that, A low-temperature oxidation reactor filled with a low-temperature oxidation catalyst is arranged between the air flow mixer and the primary conversion reactor, or a low-temperature oxidation catalyst layer is arranged at the air flow inlet reaction section of the catalyst bed layer of the primary conversion reactor; the process air flow discharged from the primary condenser is mixed with air or oxygen-enriched air of required flow rate and temperature in the air flow mixer, and then reacted in the low-temperature oxidation reactor or the low-temperature oxidation catalyst layer, and the reaction exothermicity increases the air flow temperature to 210-260°C; In the composition of the gas flow at the outlet of the gas flow mixer, the ratio of the volume fractions of H2S, SO2 and O2 is controlled to be (C H2S -2C SO2 ):C O2 =2:(0.95-1.3) and C H2S :(C SO2 +C O2 )=2:(0.9-1.1); The volume fraction of H2S and SO2 in the process gas stream discharged from the first condenser controls C H2S +C SO2 8-16v% and C H2S -2C SO2 1.2-2.5v; The air or oxygen-enriched air in the air flow mixer has an inlet temperature of 130-180°C; The gas hourly space velocity of the process gas in the low-temperature oxidation catalyst layer of the low-temperature oxidation reactor or the primary conversion reactor is 500-2000h -1 ; The low-temperature oxidation catalyst has a saturated water absorption rate of 0.6-0.7 mL / g and a specific surface area of ​​40-80 m 2 / g, modified alumina containing 10-12wt% of phosphorus oxide as P2O5 as carrier, 5-7wt% of vanadium oxide as V2O5, 4-6wt% of iron oxide as Fe2O3, and 2-4wt% of manganese oxide as MnO; prepared by the following method: A. Pseudo-boehmite powder is calcined at 420-450°C for 2-4h, and then crushed to an average particle size of 5-10μm. The obtained alumina powder is rolled into Φ6-9mm balls in a rolling ball granulator with atomized spray and a phosphoric acid aqueous solution of a required concentration; the dried balls are calcined at 1000-1030°C for 2-3h to obtain a modified alumina ball material of Φ4-6mm; the weight ratio of the alumina powder to the phosphoric acid aqueous solution is 100:(45-50); B. Place the modified alumina ball material in step A in a rotating drum, dynamically spray the vanadium oxalate aqueous solution of required concentration and volume, discharge the material after the surface of the ball becomes dry, seal and place for 5-10 hours, and dry at 110-130° C. to obtain a semi-finished ball material loaded with vanadium compound; the spray volume of the vanadium oxalate aqueous solution is 80-95% of the saturated water absorption volume of the modified alumina ball material; C. placing the semi-finished ball material loaded with vanadium compound in step B in a rotary drum controlled at 80-90° C., dynamically spraying an aqueous ferrous acetate-manganese acetate solution of required concentration and volume at 80-90° C. until the surface of the ball becomes dry, discharging the material, sealing and standing for 15-30 hours, drying at 110-130° C. under ventilation conditions, and roasting at 400-430° C. under ventilation conditions for 4-6 hours to obtain a low-temperature oxidation catalyst; the spraying volume of the aqueous ferrous acetate-manganese acetate solution is 90-95% of the saturated water absorption volume of the semi-finished ball material loaded with vanadium compound.

2. The sulfur recovery process of acid gas according to claim 1, characterized in that: When a low-temperature oxidation reactor is not provided, in the catalyst bed of the primary conversion reactor, a sulfur recovery catalyst having deoxidation protection performance is loaded below the reaction section of the gas flow inlet of the low-temperature oxidation catalyst, and a sulfur recovery catalyst containing TiO2 is loaded below that.

3. The sulfur recovery process of acid gas according to claim 1, characterized in that: When a low-temperature oxidation reactor is not provided, a sulfur recovery catalyst having COS and CS2 hydrolysis performance and deoxidation protection performance is loaded below the gas flow inlet reaction section of the primary conversion reactor filled with the low-temperature oxidation catalyst, and a sulfur recovery catalyst containing TiO2 or an activated alumina sulfur recovery catalyst is loaded below that.

4. The sulfur recovery process of acid gas according to claim 1, characterized in that: When setting up a low-temperature oxidation reactor, in the catalyst bed of the primary conversion reactor, the upper part is filled with a sulfur recovery catalyst with deoxidation protection performance or a sulfur recovery catalyst with COS, CS2 hydrolysis performance and deoxidation protection performance, and the middle and lower part is filled with a TiO2-containing sulfur recovery catalyst or an activated alumina sulfur recovery catalyst.

5. The sulfur recovery process of acid gas according to claim 1, characterized in that: The low-temperature oxidation reactor or the low-temperature oxidation catalyst layer in the primary conversion reactor, after the reaction temperature rise effect decreases to a certain extent, adopts an inlet gas flow with a volume fraction composition of 12-18v% H2S, 3-5v% SO2, and the remainder being inert gas, and a temperature of 220-250°C, at a gas hourly space velocity of 600-1200h -1 Under the above conditions, the reduction and regeneration treatment is carried out for 20-30 hours.