WO2018041171A1 - 烟气脱硝方法 - Google Patents

烟气脱硝方法 Download PDF

Info

Publication number
WO2018041171A1
WO2018041171A1 PCT/CN2017/099794 CN2017099794W WO2018041171A1 WO 2018041171 A1 WO2018041171 A1 WO 2018041171A1 CN 2017099794 W CN2017099794 W CN 2017099794W WO 2018041171 A1 WO2018041171 A1 WO 2018041171A1
Authority
WO
WIPO (PCT)
Prior art keywords
catalyst
denitration
flue gas
ammonia
reactor
Prior art date
Application number
PCT/CN2017/099794
Other languages
English (en)
French (fr)
Inventor
李欣
刘淑鹤
韩天竹
王明星
Original Assignee
中国石油化工股份有限公司
中国石油化工股份有限公司抚顺石油化工研究院
Priority date (The priority date is an assumption and is not a legal conclusion. Google has not performed a legal analysis and makes no representation as to the accuracy of the date listed.)
Filing date
Publication date
Application filed by 中国石油化工股份有限公司, 中国石油化工股份有限公司抚顺石油化工研究院 filed Critical 中国石油化工股份有限公司
Priority to US16/329,111 priority Critical patent/US11213788B2/en
Priority to EP17845473.2A priority patent/EP3498360A4/en
Priority to JP2019511754A priority patent/JP6770176B2/ja
Publication of WO2018041171A1 publication Critical patent/WO2018041171A1/zh

Links

Images

Classifications

    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D53/00—Separation of gases or vapours; Recovering vapours of volatile solvents from gases; Chemical or biological purification of waste gases, e.g. engine exhaust gases, smoke, fumes, flue gases, aerosols
    • B01D53/34—Chemical or biological purification of waste gases
    • B01D53/74—General processes for purification of waste gases; Apparatus or devices specially adapted therefor
    • B01D53/81—Solid phase processes
    • B01D53/83—Solid phase processes with moving reactants
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D53/00—Separation of gases or vapours; Recovering vapours of volatile solvents from gases; Chemical or biological purification of waste gases, e.g. engine exhaust gases, smoke, fumes, flue gases, aerosols
    • B01D53/34—Chemical or biological purification of waste gases
    • B01D53/74—General processes for purification of waste gases; Apparatus or devices specially adapted therefor
    • B01D53/86—Catalytic processes
    • B01D53/8621—Removing nitrogen compounds
    • B01D53/8625—Nitrogen oxides
    • B01D53/8631—Processes characterised by a specific device
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D46/00—Filters or filtering processes specially modified for separating dispersed particles from gases or vapours
    • B01D46/0027—Filters or filtering processes specially modified for separating dispersed particles from gases or vapours with additional separating or treating functions
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D46/00—Filters or filtering processes specially modified for separating dispersed particles from gases or vapours
    • B01D46/30—Particle separators, e.g. dust precipitators, using loose filtering material
    • B01D46/32—Particle separators, e.g. dust precipitators, using loose filtering material the material moving during filtering
    • B01D46/36—Particle separators, e.g. dust precipitators, using loose filtering material the material moving during filtering as a substantially horizontal layer, e.g. on rotary tables, drums, conveyor belts
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D53/00—Separation of gases or vapours; Recovering vapours of volatile solvents from gases; Chemical or biological purification of waste gases, e.g. engine exhaust gases, smoke, fumes, flue gases, aerosols
    • B01D53/34—Chemical or biological purification of waste gases
    • B01D53/46—Removing components of defined structure
    • B01D53/54—Nitrogen compounds
    • B01D53/56—Nitrogen oxides
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D53/00—Separation of gases or vapours; Recovering vapours of volatile solvents from gases; Chemical or biological purification of waste gases, e.g. engine exhaust gases, smoke, fumes, flue gases, aerosols
    • B01D53/34—Chemical or biological purification of waste gases
    • B01D53/74—General processes for purification of waste gases; Apparatus or devices specially adapted therefor
    • B01D53/86—Catalytic processes
    • B01D53/88—Handling or mounting catalysts
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D53/00—Separation of gases or vapours; Recovering vapours of volatile solvents from gases; Chemical or biological purification of waste gases, e.g. engine exhaust gases, smoke, fumes, flue gases, aerosols
    • B01D53/34—Chemical or biological purification of waste gases
    • B01D53/74—General processes for purification of waste gases; Apparatus or devices specially adapted therefor
    • B01D53/86—Catalytic processes
    • B01D53/90—Injecting reactants
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J35/00—Catalysts, in general, characterised by their form or physical properties
    • B01J35/40—Catalysts, in general, characterised by their form or physical properties characterised by dimensions, e.g. grain size
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J35/00—Catalysts, in general, characterised by their form or physical properties
    • B01J35/60—Catalysts, in general, characterised by their form or physical properties characterised by their surface properties or porosity
    • B01J35/61—Surface area
    • B01J35/613—10-100 m2/g
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J35/00—Catalysts, in general, characterised by their form or physical properties
    • B01J35/60—Catalysts, in general, characterised by their form or physical properties characterised by their surface properties or porosity
    • B01J35/61—Surface area
    • B01J35/615—100-500 m2/g
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J8/00—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
    • B01J8/08—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with moving particles
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J8/00—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
    • B01J8/08—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with moving particles
    • B01J8/085—Feeding reactive fluids
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J8/00—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
    • B01J8/08—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with moving particles
    • B01J8/12—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with moving particles moved by gravity in a downward flow
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D2251/00—Reactants
    • B01D2251/20—Reductants
    • B01D2251/206—Ammonium compounds
    • B01D2251/2062—Ammonia
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D2255/00—Catalysts
    • B01D2255/20—Metals or compounds thereof
    • B01D2255/207—Transition metals
    • B01D2255/20707—Titanium
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D2255/00—Catalysts
    • B01D2255/20—Metals or compounds thereof
    • B01D2255/207—Transition metals
    • B01D2255/20723—Vanadium
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D2255/00—Catalysts
    • B01D2255/20—Metals or compounds thereof
    • B01D2255/207—Transition metals
    • B01D2255/20769—Molybdenum
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D2255/00—Catalysts
    • B01D2255/20—Metals or compounds thereof
    • B01D2255/207—Transition metals
    • B01D2255/20776—Tungsten
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D2255/00—Catalysts
    • B01D2255/90—Physical characteristics of catalysts
    • B01D2255/92—Dimensions
    • B01D2255/9202—Linear dimensions
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D2255/00—Catalysts
    • B01D2255/90—Physical characteristics of catalysts
    • B01D2255/92—Dimensions
    • B01D2255/9207—Specific surface
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D2257/00—Components to be removed
    • B01D2257/30—Sulfur compounds
    • B01D2257/302—Sulfur oxides
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D2257/00—Components to be removed
    • B01D2257/40—Nitrogen compounds
    • B01D2257/404—Nitrogen oxides other than dinitrogen oxide
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D2258/00—Sources of waste gases
    • B01D2258/02—Other waste gases
    • B01D2258/0283—Flue gases

Definitions

  • the invention relates to the field of flue gas denitration technology, in particular to a flue gas denitration method.
  • NO x Collectively referred to as nitrogen oxides NO x, is one of the major sources of air pollution. The most harmful ones are NO and NO 2 .
  • SCR selective catalytic reduction
  • the FCC regenerative flue gas denitrification treatment of coal-fired power plants and refineries mainly adopts the SCR method, and is equipped with wet scrubbing desulfurization and dust removal.
  • FCC flue gas as an example, the main process is as follows: FCC regeneration flue gas at 500-600 °C is first recovered by waste heat boiler, and the flue gas temperature is lowered to 320-400 ° C.
  • the SCR fixed-bed reactor is used for denitration reaction to remove smoke.
  • the flue gas temperature is lowered to 150 ⁇ 200 °C, and then into dust desulfurization scrubber with an alkaline absorption liquid while the SO x and dust in flue gas wash down
  • the flue gas temperature is reduced to 55-60 ° C.
  • the desulfurization waste absorption liquid should be subjected to liquid-solid separation by steps of sedimentation, filtration and concentration.
  • the liquid after solid-liquid separation is oxidized by air aeration, COD is discharged to the standard, and the solid is landfilled.
  • the existing SCR denitration process uses a fixed bed denitration reactor, the catalyst is honeycomb, plate or corrugated, and the catalyst is placed in the reactor in the form of a module.
  • reductant NH 3 is first injected in the previous reaction bed, so that NH 3 and NO x in the flue gas sufficiently mixed by a denitration catalyst bed, catalytic reduction of NO x to N 2.
  • the flue gas generally contains SO 2 , SO 3 , O 2 and water vapor
  • the ammonia in the reaction zone when the ammonia in the reaction zone is excessive (ammonia escape), it will react with SO 3 to form an ammonium salt, and the resulting ammonium salt (NH 4 HSO 4 )
  • NH 4 HSO 4 ammonium salt
  • It is liquid at 180 ⁇ 240 °C, has viscosity, and is easy to adhere to the heat exchange tube of the economizer of the downstream device of the SCR denitration reactor. It binds the dust in the flue gas and causes fouling of the heat exchange tube layer. Corrosion affects the operating cycle of the device.
  • the uniformity of ammonia injection at the inlet of the SCR fixed bed reactor generally requires a positive or negative deviation of less than 5%.
  • the catalyst activity decreased, when the reactor is not compliance outlet NO x emissions, it is necessary to replace the catalyst.
  • the operating cycle of a typical SCR device requires at least 3-4 years, otherwise it will affect the operation of the main device.
  • the denitration rate of the SCR device is required to be at least 60 to 90%.
  • the activity of the catalyst is at least about 60%. It can be seen that the utilization rate of the catalyst by the fixed bed SCR reactor is low.
  • CN102008893A discloses the use of a conventional moving bed reactor for SCR reaction. Since the pressure of flue gas passing through the bed requires a pressure drop, it is not suitable for a large amount of flue gas and dust in the flue gas. The working condition of low residual pressure of flue gas. Moreover, the flue gas moves downward by gravity. When the reactor is large, there is a problem of uniformity of movement, and the bed layer is easily blocked by the bridge. The particle size of the catalyst is relatively high, and the uniformity of the reaction is also problematic.
  • the object of the present invention is to overcome the shortcomings of the existing SCR process and equipment, such as low catalyst utilization, low operational flexibility, and the need to use wet scrubbing after flue gas denitration, to provide a flue gas denitration method.
  • the present invention provides a flue gas denitration method, which comprises: in a denitration reactor, in the presence of ammonia gas, passing flue gas from bottom to top through a plurality of catalyst beds for denitration a reaction wherein each catalyst bed comprises a catalyst support member and a particulate denitration catalyst deposited on the catalyst support member, wherein in the adjacent upper and lower catalyst beds, the two catalyst support members are operated in opposite directions, the upper layer
  • the particulate denitration catalyst moves to the end of the catalyst support member with the catalyst support member and falls freely to the starting end of the underlying catalyst support member by gravity.
  • the invention also provides a flue gas denitration method, the method comprises: the flue gas enters from the bottom of the denitration reactor, and the mixed gas containing ammonia gas is injected into the flue gas through the ammonia spray grid, and the air flow passes through from bottom to top.
  • multiple horizontal cross catalyst bed arrangement the denitration reaction is carried out to remove NO X, while being in the flue gas dust removal by filtration of the catalyst bed, after the flue gas denitrification and dust discharged from the top of the reactor, the next step desulfurization process;
  • the catalyst bed layer is composed of a mesh conveyor belt and a granular denitration catalyst accumulated on the conveyor belt.
  • the operation direction of the adjacent upper and lower conveyor belts is opposite.
  • the upper granular denitration catalyst moves the belt end with the conveyor belt and falls freely to the operation of the lower conveyor belt by gravity. At the beginning of the direction, the particulate denitration catalyst falls to the catalyst recovery unit at the end of the last conveyor belt for recovery.
  • the method of the present invention has the following advantages:
  • the method of the present invention is flexible and adaptable, and the residence time of the catalyst in the reactor can be adjusted by adjusting the moving speed of the denitration catalyst on the catalyst supporting member, by adjusting the bed thickness of the catalyst on the catalyst supporting member. adjusting the flue gas through the catalyst bed of the reaction time, it is possible to handle larger flue gas NO X concentration range to maximize catalyst utilization.
  • the catalyst can be repeatedly used, and the catalyst can be renewed at any time. Therefore, the catalyst utilization rate is much higher than that of the conventional fixed bed reactor, and the amount of the catalyst is greatly reduced, and the catalyst can be displaced online to ensure the stable activity of the catalyst inside the reactor.
  • the catalytic particles are in reverse contact with the flue gas in the reactor, and the catalyst bed in the upper portion of the reactor can adsorb excess ammonia gas, and the catalyst particles react with the flue gas to consume ammonia during the downward movement or
  • the ammonia gas is adsorbed in the catalyst particles and taken out of the reactor, no ammonia escape occurs, the uniformity of the initial ammonia distribution of the bed is not required, the secondary pollution caused by ammonia escape and the problem of blocking the bed by ammonium hydrogen sulfate are avoided.
  • the operating cycle of the device is extended.
  • the invention adopts a granular catalyst bed layer to filter dust in the flue gas, and has simpler process than the conventional technology, and the residual dust in the catalyst bed can leave the reaction system along with the catalyst to achieve the effect of simultaneous dust removal.
  • the large-diameter spherical particle catalyst has a higher specific surface area of contact with the flue gas than the conventional fixed-bed reactor, and thus the denitration efficiency is high.
  • the thickness of the bed on the catalyst supporting member of the present invention can be adjusted to adapt to the conditions of large smoke volume, flue gas dusting, and low residual pressure of the flue gas. And the bed moves along the catalyst support member in the reactor, avoiding reactor catalyst retention and no clogging.
  • the out-selling catalyst is gradually moved from the top of the reactor to the next bed in a liquid phase similar to the plate column, and an activity gradient is established in the longitudinal direction of the reactor, which is advantageous for deep denitration, full utilization of catalyst activity, and uniform reaction.
  • FIG. 1 is a schematic view showing the structure of a denitration reactor used in the flue gas denitration method of the present invention.
  • the flue gas denitration method of the present invention comprises: in a denitration reactor, in a presence of ammonia gas, passing the flue gas from bottom to top through a plurality of catalyst beds for denitration reaction, wherein each catalyst bed layer
  • the catalyst supporting member and the granular denitration catalyst deposited on the catalyst supporting member wherein the two catalyst supporting members are operated in opposite directions in the adjacent upper and lower catalyst beds, and the upper granular denitration catalyst moves with the catalyst supporting member To the end of the catalyst support member, it is free to fall by gravity to the starting end of the underlying catalyst support member.
  • the catalyst support member is reticulated.
  • the catalyst support member is a mesh conveyor belt, in which case the particulate denitration catalyst on the mesh conveyor belt can be transferred from one end to the other.
  • the movement speed of the particulate denitration catalyst is flexibly adjusted by controlling the conveying speed of the conveyor belt.
  • a web conveyor belt is separately provided for each catalyst bed so that the moving speed and thickness of each catalyst bed can be independently controlled, and the mutual interference between the beds is small.
  • the plurality of catalyst beds are arranged parallel to each other from top to bottom, that is, a plurality of mesh conveyor belts as catalyst supporting members are arranged in parallel with each other, thereby ensuring a granular denitration catalyst on a single mesh conveyor belt. It is evenly distributed, which is more conducive to making full use of the catalyst activity.
  • the mesh belt may be a conventional metal mesh belt in the art, preferably a stainless steel mesh belt.
  • the mesh size of the mesh conveyor belt is required to be smaller than the particle size of the particulate denitration catalyst to ensure that the catalyst particles do not fall from the mesh.
  • the mesh size on the mesh belt can be from 0.1 to 3 mm, preferably from 1.5 to 2.5 mm.
  • the number and width of the catalyst bed can be selected according to actual needs and reactor size.
  • the number of the catalyst beds is preferably from 3 to 10, more preferably from 3 to 8.
  • the gap between the catalyst supporting member and the reactor wall may be 2 to 50 mm, preferably 2 to 5 mm in the width direction of the catalyst supporting member (preferably the mesh belt).
  • the vertical distance between adjacent two catalyst beds may be from 50 to 2000 mm.
  • the catalyst supporting member is a mesh belt and the plurality of catalyst supporting members are arranged in parallel with each other, the vertical distance between adjacent two catalyst beds is 1200 to 2000 mm, preferably 1400. ⁇ 1600mm.
  • the stacking height of the particulate denitration catalyst on the mesh belt may be 50 to 500 mm. It is preferably 200 to 300 mm.
  • the conveying speed of the web belt may be from 0.1 to 10 mm/s, preferably from 0.5 to 2 mm/s.
  • the flue gas may be selected from the group consisting of coal-fired power plant flue gas, FCC regenerative flue gas, refinery process furnace flue gas, and chemical furnace flue gas (such as ethylene cracking furnace flue gas, etc.), generally containing NO x, SO x and impurities.
  • the impurities are generally dust, water, CO 2 and O 2 and the like.
  • the temperature of the flue gas entering the denitration reactor may be from 300 to 420 ° C, preferably from 340 to 400 ° C.
  • ammonia gas can be introduced in a manner conventional in the art, for example, it can be introduced in the form of ammonia gas.
  • the ammonia gas is introduced in the form of a mixed gas containing ammonia and air, and the volume concentration of the ammonia gas in the mixed gas may be from 0.5 to 10%, preferably from 3 to 7%.
  • the mixed gas containing ammonia and air is preferably injected below the catalyst bed, that is, the mixture containing ammonia and air is at the lowest catalyst bed. Inject below.
  • the mixed gas containing ammonia and air may be injected through an ammonia spray grid provided at the inlet of the flue gas.
  • the ammonia spray grid may be a conventional ammonia spray grid in the art, but in the conventional SCR process and equipment, the ammonia spray grid performance requirement ensures that the ammonia gas concentration distribution deviation is less than 5%, and in the present invention
  • the concentration range deviation of the ammonia gas is not critical, as long as the concentration deviation of the ammonia gas is 30% or less, and may be, for example, 5% to 30%, preferably 12% to 18%.
  • the molar ratio of ammonia gas to nitrogen oxides in the flue gas as nitrogen atoms may be from 0.9 to 1.15:1.
  • the flow rate of the flue gas may be 2-15 m/s, preferably 4-10 m/s; and the reaction residence time may be 0.5-20 s.
  • the particulate denitration catalyst has a property of a particle size of 3 to 6 mm, a bulk density of 0.2 to 0.8 g/cm 3 and a specific surface area of 80 to 120 m 2 /g.
  • the composition of the denitration catalyst may be a component commonly used in the art, and the content of each component is 0.01 to 1% by weight of V, 88 to 99% by weight of Ti, 0.1 to 10 by weight, based on the weight of the catalyst. % W and 0.01 to 1% by weight of Mo.
  • the preparation process of the denitration catalyst may be: material dry mixing-kneading-filtration-practice-extrusion granulation-dry-dry-baking-finished product, wherein the conditions involved in each step of catalyst preparation are well known to those skilled in the art.
  • Extrusion granulation can be equipped with different types of kneading extrusion granulators depending on the catalyst particle size.
  • the flue gas denitration method of the present invention comprises: the flue gas enters from the bottom of the denitration reactor, and the mixed gas containing the ammonia gas and the air is injected into the flue gas through the ammonia spray grid, and the air flow is self-injected.
  • the catalyst bed consists of a mesh conveyor belt and a particulate denitration catalyst accumulated on the conveyor belt, the adjacent upper and lower layers of the mesh conveyor belts operate in opposite directions, and the upper granular denitration catalyst moves to the end of the conveyor belt with the conveyor belt, relying on Gravity falls freely to the beginning of the running direction of the lower conveyor belt, and the particulate denitration catalyst falls to the catalyst recovery unit at the end of the last conveyor belt for recovery.
  • the particulate denitration catalyst is fed to the first layer of the conveyor belt on the top of the denitration reactor via a catalyst addition tube, and the catalyst is dropped on the network conveyor to form a catalyst bed.
  • the catalyst recovery device generally adopts a common equipment such as a catalyst storage tank and a catalyst hopper; the recovered catalyst particles are sieved to remove the dust and the crushed catalyst particles, and can be repeatedly used.
  • the denitration reactor of the present invention comprises a reactor housing 13 and a plurality of catalyst supporting members staggered horizontally from top to bottom in the reactor housing 13, and the plurality of catalyst supporting members It is configured that, in the adjacent upper and lower two catalyst supporting members, the two catalyst supporting members are operated in opposite directions, and the granular denitration catalyst on the catalyst supporting member moves to the end of the catalyst supporting member with the catalyst supporting member, and is free to rely on gravity. Fall to the beginning of the underlying catalyst support member.
  • the catalyst support member is a mesh conveyor belt.
  • the plurality of catalyst beds are arranged in parallel with each other from top to bottom, that is, a plurality of mesh belts as catalyst supporting members are arranged in parallel with each other.
  • the mesh conveyor belt may be a conventional metal mesh conveyor belt in the art, preferably a stainless steel mesh conveyor belt.
  • the mesh size of the mesh conveyor belt is required to be smaller than the particle size of the particulate denitration catalyst to ensure that the catalyst particles do not fall from the mesh.
  • the mesh size on the mesh belt can be from 0.1 to 3 mm, preferably from 1.5 to 2.5 mm.
  • the denitration reactor may further include an inner cylinder 12 penetrating the upper end and the lower end of the reactor casing 13, and the mesh belt is transversely penetrated The inner cylinder 12.
  • the mesh belt is usually driven by a motor, and the belt drive wheel drives the belt to rotate.
  • the drive wheel 9 of the mesh belt is disposed in a chamber between the inner cylinder 12 and the reactor housing 13.
  • the lower end of the inner cylinder 12 is provided with a flue gas inlet
  • the upper end is provided with a flue gas outlet
  • the ammonia spraying unit 6 is disposed at the flue gas inlet.
  • the head (starting end) of the uppermost catalyst supporting member is correspondingly provided with a catalyst addition pipe 7 for adding a particulate denitration catalyst, and a tail portion (end of the lowermost catalyst supporting member)
  • a catalyst collecting hopper is provided for recovering the catalyst particles.
  • v% is a volume fraction
  • wt% is a mass fraction
  • the specific surface area and pore volume of the catalyst were measured by ASAP 2420 large multi-station automatic surface area and pore analyzer.
  • the deviation of the ammonia concentration distribution in the test case is an index of the design, which can be obtained through experiments and calculations.
  • the calculation uses the CFD fluid dynamics calculation software to simulate the structure of the ammonia spray grid to determine the deviation of the ammonia concentration distribution, and the average ammonia concentration on the cross section. And distribution range calculation.
  • Escape ammonia and dust content were measured using a CEMS flue gas online analyzer.
  • the out-of-stock reactor comprises a reactor shell 13, a reactor inner cylinder 12, an ammonia spray grid 6, a catalyst addition tube 7, and a mesh conveyor.
  • the conveyor drive wheel 9 is in the outer sealed chamber of the reactor
  • the conveyor belt 8 is attached to the conveyor drive wheel 9, across the reactor inner cylinder 12
  • the catalyst addition tube 7 is at the top of the outer chamber of the reactor
  • the catalyst is added to the bottom outlet of the tube 7. Facing one end of the conveyor belt
  • the catalyst hopper 11 is at The outside of the reactor seals the bottom of the chamber, and the catalyst discharge pipe 10 is at the bottom of the catalyst hopper 11.
  • the operation process of the flue gas denitration is as follows: the granular denitration catalyst 3 is filled on the first layer of the mesh belt 8 through the catalyst addition pipe 7 to form a bed, and the belt driving wheel 9 drives the bed movement on the conveyor belt 8, the bed
  • the layer passes through the inner tube 12 of the reactor, enters the outer sealed chamber of the reactor, falls by gravity to the next conveyor belt, forms a bed layer, and moves in the opposite direction under the driving of the belt driving wheel 9, according to the above operation mode.
  • Continuously running conveyor belt bed flue gas 1 enters from the bottom of the denitration bed reactor, and ammonia-containing gas mixture 2 is filled into the flue gas 1 through the ammonia spray grid 6, and the two are mixed from bottom to top.
  • the particulate catalyst denitration catalyst falls within hopper 11 at the end of the last layer of the conveyor belt, particulate after the denitration catalyst of the denitration catalyst by discharge pipe 10 exiting the reactor is recovered, and the dust removing NO x purge gas 5 discharged from the top of the reactor.
  • the catalyst fine particle active component of the granular denitration catalyst is an oxide of V, an oxide of Ti, an oxide of W and an oxide of Mo, and the active component is an oxide, and the catalyst is a spherical particle, and the mass ratio is as follows: V is 0.01 wt%, Ti was 99 wt%, W was 0.1 wt%, Mo was 0.02 wt%; the catalyst particle size was 5 mm; the bulk density was 0.68 g/cm 3 and the specific surface area was 40 m 2 /g.
  • the FCC regeneration flue gas is heated by the boiler, and the temperature is lowered from 650 ° C to 400 ° C of the SCR denitration reaction temperature; the flow rate of the mixed gas containing ammonia gas provided by the raw material supply area is 1120 Nm 3 /h, wherein the ammonia concentration is 4 v%, and the reaction
  • the inner sealing cavity of the device is 8m long ⁇ 6m wide ⁇ 8m high; the reaction time is 0.5s.
  • Three layers of conveyor belt are set.
  • the height of the catalyst bed on each conveyor belt is 300mm.
  • the conveyor belt is 9m long and 5.8m wide. Stainless steel is used.
  • the mesh conveyor belt has a gap diameter of 3 mm and a drive wheel diameter of 300 mm.
  • the space between the upper and lower conveyor belts is 1300 mm, leaving sufficient space for maintenance.
  • purification of flue gas can be guaranteed NO x in an amount of 100mg / Nm 3, the dust content of less than 10mg / Nm 3, the focus control area meet environmental requirements; Flue Gas and SO 2 in the flue gas desulfurization and dust removal by Dust can be discharged through the chimney.
  • composition of the catalyst was the same as in Example 1.
  • the FCC regeneration flue gas is heated by the boiler, and the temperature is lowered from 650 ° C to 300 ° C of the SCR denitration reaction temperature; the flow rate of the mixed gas containing ammonia gas provided by the raw material supply area is 1000 Nm 3 /h, wherein the ammonia concentration is 3 v%;
  • the inner sealed cavity size is 8m long ⁇ 6m wide ⁇ 15m high; the reaction time is 2s, 10 layers of conveyor belt are set, the height of the catalyst bed on each conveyor belt is 500mm, the conveyor belt size is 9m long ⁇ 5.8m wide, stainless steel mesh is used.
  • the gap diameter is 3mm
  • the drive wheel diameter is 300mm
  • the space between the upper and lower conveyor belts is 1500mm, leaving sufficient maintenance space.
  • purification of flue gas can be guaranteed NO x in an amount of 100mg / Nm 3, the dust content of less than 5mg / Nm 3, the focus control area meet environmental requirements; Flue Gas and SO 2 in the flue gas desulfurization and dust removal by Dust can be discharged through the chimney.
  • the catalyst composition was the same as in Example 1, and the catalyst particles had a diameter of 3 mm.
  • the FCC regeneration flue gas is heated by the boiler, and the temperature is lowered from 650 ° C to 300 ° C of the SCR denitration reaction temperature; the flow rate of the mixed gas containing ammonia gas provided by the raw material supply area is 2000 Nm 3 /h, wherein the ammonia gas concentration is 2.8 v%;
  • the inner sealed chamber of the reactor is 8m long ⁇ 6m wide ⁇ 6m high; the reaction time is 0.5s, and 3 layers of conveyor belt are provided.
  • the height of the catalyst bed on each conveyor belt is 50mm, and the conveyor belt is 9m long and 5.8m wide.
  • Stainless steel mesh conveyor belt the gap diameter is 2.5mm, the drive wheel diameter is 500mm, and the space between the upper and lower conveyor belts is 2000mm, leaving enough room for maintenance.
  • purification of flue gas can be guaranteed NO x in an amount of 100mg / Nm 3, the dust content of less than 5mg / Nm 3, the focus control area meet environmental requirements; Flue Gas and SO 2 in the flue gas desulfurization and dust removal by Dust can be discharged through the chimney.
  • the catalyst composition was the same as in Example 1, and the catalyst particles had a diameter of 6 mm.
  • the FCC regeneration flue gas is heated by the boiler, and the temperature is lowered from 650 ° C to 300 ° C of the SCR denitration reaction temperature; the flow rate of the mixed gas containing ammonia gas provided by the raw material supply area is 2000 Nm 3 /h, wherein the ammonia gas concentration is 2.8 v%;
  • the inner sealed chamber of the reactor is 8m long ⁇ 6m wide ⁇ 6m high; the reaction time is 0.5s.
  • Three layers of conveyor belt are set. The height of the catalyst bed on each conveyor belt is 500mm.
  • the conveyor belt is 9m long and 5.8m wide.
  • Stainless steel mesh conveyor belt the diameter of the gap is 0.2mm, the diameter of the driving wheel is 300mm, and the space between the upper and lower conveyor belts is 2000mm, leaving enough room for maintenance.
  • purification of flue gas can be guaranteed NO x in an amount of 100mg / Nm 3, the dust content of less than 5mg / Nm 3, the focus control area meet environmental requirements; Flue Gas and SO 2 in the flue gas desulfurization and dust removal by Dust can be discharged through the chimney.
  • the catalyst composition was the same as in Example 1, and the catalyst particles had a diameter of 5 mm.
  • the FCC regenerated flue gas is heated by the boiler, and the temperature is lowered from 650 ° C to 400 ° C of the SCR denitration reaction temperature; the flow rate of the ammonia-containing mixture gas supplied from the raw material supply area is 70 Nm 3 /h, wherein the ammonia concentration is 4 v%, and the reaction
  • the inner sealed cavity size is 1.2m long ⁇ 1m wide x 1.2m high; the reaction time is 0.5s, three layers of conveyor belt are set, the height of the catalyst bed on each conveyor belt is 40mm, and the conveyor belt is 1.5m long and 0.8m wide.
  • the stainless steel mesh belt is used, the gap diameter is 3mm, the driving wheel diameter is 50mm, and the space between the upper and lower conveyor belts is 200mm.
  • Example 2 Same as Example 1, except that the reactor was replaced with a conventional fixed bed reactor, the catalyst was a honeycomb catalyst, and its composition was the same as that of Example 1, using a modular packing, the height of the single catalyst module was 1 m, and the reactor size was 4.4 m ⁇ 4.6m, the catalyst is filled with three layers.
  • the FCC regeneration flue gas is heated by the boiler, the temperature is lowered from 650 °C to the SCR denitration reaction temperature of 350 ° C; the flow rate of the mixed gas containing ammonia gas provided by the raw material supply area is 1000 Nm 3 /h, wherein The ammonia concentration was 3 v%.
  • the mixture containing ammonia gas is added in the upstream flue at a certain distance from the inlet of the reactor.
  • the concentration of ammonia in the flue gas at the inlet of the reactor is less than 5%, and then enters the SCR reactor.
  • the reaction, after denitration reaction, purification of flue gas can be guaranteed NO x in an amount of 100mg / Nm 3, of flue gas entering the downstream denitration apparatus continues to heat, dust desulfurization to meet the environmental requirements focus on the control areas.
  • Example 2 Same as Example 2 except that the reactor was replaced with a conventional fixed bed reactor.
  • Example 3 Same as Example 3 except that the reactor was replaced with a conventional fixed bed reactor.
  • Example 4 Same as Example 4 except that the reactor was replaced with a conventional fixed bed reactor.
  • Example 5 Same as Example 5 except that the reactor was replaced with a conventional fixed bed reactor, the catalyst was a honeycomb catalyst, the composition of which was the same as that of Example 5, and the modular catalyst was used.
  • the height of the single catalyst module was 150 mm, and the reactor size was 700 mm ⁇ 800 mm.
  • the catalyst is filled in three layers. First, the FCC regenerated flue gas is heated by the boiler, and the temperature is lowered from 650 ° C to 350 ° C of the SCR denitration reaction temperature; the flow rate of the mixed gas containing ammonia gas provided by the raw material supply area is 60 Nm 3 /h, wherein ammonia gas The concentration is 3v%.
  • the mixture containing ammonia gas is added in the upstream flue at a certain distance from the inlet of the reactor. After the mixed diffusion of the ammonia spray grid, the concentration of ammonia in the flue gas at the inlet of the reactor is less than 5%, and then enters the SCR reactor. the reaction, after denitration reaction, purification of flue gas can be guaranteed NO x in an amount of 100mg / Nm 3, of flue gas entering the downstream denitration apparatus continues to heat, dust desulfurization to meet the environmental requirements focus on the control areas.
  • the reactor is still used in Comparative Example 1, due to the increased concentration of NO x in the flue gas are too great to ensure that ammonia slip ⁇ 3mgNm 3, the denitration the NO x concentration of 1000 ⁇ 1300mg / Nm 3, not discharge standards, and dust still Enter the desulfurization and dust removal system for treatment.
  • Example 1 to 4 and Comparative Examples 1 to 4 are shown in Table 1.
  • the operating cycles and catalyst dosages of Examples 5 and 5 are shown in Table 2.
  • the requirements for the ammonia distribution of the ammonia spray grid are shown in Table 3. .
  • the flue gas denitration method according to the present invention can improve the utilization rate of the catalyst and greatly reduce the amount of the catalyst; moreover, the method of the present invention does not cause ammonia slippage.
  • the uniformity of the initial ammonia distribution of the bed is not required; in addition, the method of the present invention can obtain a better dust removal effect.

Landscapes

  • Chemical & Material Sciences (AREA)
  • Engineering & Computer Science (AREA)
  • Chemical Kinetics & Catalysis (AREA)
  • Environmental & Geological Engineering (AREA)
  • General Chemical & Material Sciences (AREA)
  • Analytical Chemistry (AREA)
  • Biomedical Technology (AREA)
  • Oil, Petroleum & Natural Gas (AREA)
  • Health & Medical Sciences (AREA)
  • Organic Chemistry (AREA)
  • Materials Engineering (AREA)
  • Exhaust Gas Treatment By Means Of Catalyst (AREA)
  • Devices And Processes Conducted In The Presence Of Fluids And Solid Particles (AREA)
  • Catalysts (AREA)

Abstract

一种烟气脱硝方法，包括：在脱硝反应器中，在氨气的存在下，使烟气从下到上穿过多个催化剂床层，进行脱硝反应，其中，每个催化剂床层包括催化剂支撑部件和堆积在催化剂支撑部件上的颗粒状脱硝催化剂，并且在单个催化剂床层上，颗粒状脱硝催化剂在催化剂支撑部件上沿着同一个方向移动；在相邻的两个催化剂床层中，颗粒状脱硝催化剂从上一个催化剂支撑部件的尾部落至下一个催化剂支撑部件的头部，使得颗粒状脱硝催化剂沿着各个催化剂支撑部件往复折返移动。

Description

烟气脱硝方法 技术领域
本发明涉及烟气脱硝技术领域，具体涉及烟气脱硝方法。
背景技术
氮氧化物总称为NOx，是大气污染的主要污染源之一。危害最大的主要是NO和NO2。NOx的主要危害如下：(1)对人体有毒害作用；(2)对植物有毒害作用；(3)可形成酸雨、酸雾；(4)与碳氢化合物形成光化学烟雾；(5)破坏臭氧层。
烟气脱硝是指脱除烟气中的NOx，按治理工艺可分为湿法脱硝和干法脱硝。主要包括：酸吸收法、碱吸收法、选择性催化还原法、非选择性催化还原法、吸附法、离子体活化法等。国内外一些科研人员还开发了用微生物来处理NOx废气的方法。但是具有工业价值，运用最广泛的是选择性催化还原法(SCR)。
目前燃煤电厂、炼油厂的FCC再生烟气脱硝治理主要采用SCR法，并配套湿法洗涤脱硫除尘。以FCC烟气为例，主要流程如下：500～600℃的FCC再生烟气先经过余热锅炉进行回收热量，烟气温度降低至320～400℃进入SCR固定床反应器进行脱硝反应，脱除烟气中的NOx，然后再回到余热锅炉进行回收热量，烟气温度降至150～200℃，然后进入脱硫除尘洗涤塔，采用碱性吸收液同时将烟气中的SOx与粉尘洗下来，烟气温度降低至55～60℃排放。脱硫废吸收液要进行沉降、过滤、浓缩等步骤进行液固分离，液固分离后的清液采用空气曝气氧化，COD达标排放，固体进行填埋。
现有的SCR脱硝工艺均采用固定床脱硝反应器，催化剂采用蜂窝式、板式或波纹式，催化剂以模块的形式放置在反应器内。在反应床层前面首 先注入还原剂NH3，让NH3与烟气中的NOx充分混合，通过脱硝催化剂床层，将NOx催化还原为N2。
现有技术存在以下问题：
(1)由于烟气中一般含有SO2、SO3、O2和水蒸气，当反应区氨过剩(氨逃逸)时会与SO3反应生成铵盐，生成的铵盐(NH4HSO4)，在180～240℃温度下呈液态，具有粘性，容易附着在SCR脱硝反应器的下游装置省煤器的换热管上，粘接烟气中的粉尘，引发换热管层的结垢堵塞与腐蚀，影响装置运行周期。为了避免氨逃逸，SCR固定床反应器入口喷氨均匀性一般要求正负偏差小于5％。
(2)烟气中的NOx含量与主装置的工艺条件相关，变化波动范围较大，而SCR固定床反应器的催化剂量是固定的，一旦NOx浓度范围超出设计值，则净化烟气的NOx不能达标排放。因此固定床的操作弹性较小。
(3)固定床反应器在运行期间，催化剂的活性逐渐下降，当反应器出口NOx无法达标排放时，就需要更换催化剂。一般SCR装置的运行周期至少要求3-4年，否则会影响主装置的运转。一般SCR装置脱硝率需求至少在60～90％以上，当更换催化剂时，催化剂的活性至少还有60％左右。由此可见，采用固定床SCR反应器对催化剂的利用率较低。
(4)一般烟气脱硝以后，还要采用湿法洗涤除尘，与脱硫一起进行，除尘后，脱硫废液还要进行液固分离，流程繁长，操作复杂，投资与操作费用高。
目前，尚未报道工业化的移动床SCR反应器，CN102008893A公开了采用传统的移动床反应器进行SCR反应，由于烟气穿过床层需要压降，不适合烟气量大、烟气中含尘、烟气余压不高的工况。并且烟气通过重力向下移动，在反应器较大时，存在移动均匀性问题，床层容易架桥堵塞，对催化剂的颗粒度要求比较高，反应均匀性也存在问题。
发明内容
本发明的目的是为了克服现有的SCR工艺和设备存在的催化剂利用率较低、操作弹性较小以及在烟气脱硝后还需要采用湿法洗涤除尘的缺陷，提供一种烟气脱硝方法。
为了实现上述目的，本发明提供了一种烟气脱硝方法，该方法包括：在脱硝反应器中，在氨气的存在下，使烟气从下到上穿过多个催化剂床层，进行脱硝反应，其中，每个催化剂床层包括催化剂支撑部件和堆积在该催化剂支撑部件上的颗粒状脱硝催化剂，在相邻的上下两个催化剂床层中，两个催化剂支撑部件的运行方向相反，上层颗粒状脱硝催化剂随催化剂支撑部件移动至催化剂支撑部件末端，依靠重力自由落至下层催化剂支撑部件的起始端。
本发明还提供了一种烟气脱硝方法，该方法包括：烟气从脱硝反应器底部进入，含氨气的混合气经喷氨格栅加注到烟气中，气流自下而上穿过多层水平交错排列的催化剂床层，进行脱硝反应脱除NOX，烟气中的粉尘同时被催化剂床层过滤除尘，经过脱硝与除尘的烟气从反应器顶部排出，进行下一步脱硫处理；其中催化剂床层由网状传送带和传送带上堆积的颗粒状脱硝催化剂组成，相邻上下两层传送带的运行方向相反，上层颗粒状脱硝催化剂随传送带移动传送带末端，依靠重力自由落至下层传送带的运行方向的起始端，颗粒状脱硝催化剂在最后一层传送带的末端落入催化剂回收装置进行回收。
与现有技术相比较，本发明所述的方法具有以下优点：
(1)本发明所述的方法灵活性和适应性强，可以通过调节脱硝催化剂在催化剂支撑部件上的移动速度调节催化剂在反应器内的停留时间，通过调节催化剂支撑部件上催化剂的床层厚度，调节烟气通过催化剂床层的反应时间，因此可以处理NOX浓度变化范围较大的烟气，最大限度地提高催化剂的利用率。
(2)本发明中催化剂可以反复使用，催化剂随时可以更新，因此催化剂使用率大大高于传统固定床反应器，催化剂的用量大大降低，可实现催化剂在线置换，保证反应器内部催化剂稳定的活性。
(3)本发明中催化颗粒在反应器内与烟气逆向接触，反应器上部的催化剂床层可以吸附过量的氨气，催化剂颗粒在向下移动的过程中与烟气反应把氨消耗掉或将氨气吸附在催化剂颗粒内带出反应器，不会出现氨逃逸，对床层初始氨分布的均匀性要求不高，避免了氨逃逸造成二次污染以及硫酸氢铵堵塞床层的问题，延长了装置的运行周期。
(4)本发明采用颗粒状的催化剂床层对烟气中的粉尘有过滤作用，相比常规技术流程简单，同时催化剂床层内残存的粉尘可以随催化剂一起离开反应系统，达到同时除尘的效果；同时大直径的球型颗粒催化剂与烟气接触比表面积相比传统的固定床反应器高，因此脱硝效率高。
(5)与传统的移动床反应器相比，本发明的催化剂支撑部件上的床层厚度可以调节，适应大烟气量、烟气带尘、烟气余压低的工况。并且床层在反应器内沿着催化剂支撑部件移动，避免了反应器催化剂滞留，不会出现堵塞现象。
(7)脱销催化剂以类似于板式塔液相从反应器顶部逐步移动到下一床层，在反应器内纵向上建立活性梯度，有利于深度脱硝，充分利用催化剂活性，并且反应均匀。
附图说明
图1是本发明所述的烟气脱硝方法所采用的脱硝反应器的结构示意图。
附图标记说明
1       烟气                    2      含有氨气的混合气
3       新加入的催化剂颗粒      4      脱硝后的催化剂颗粒
5       净化气                  6      喷氨格栅
7       催化剂加入管            8      催化剂支撑部件
9       传送带驱动轮            10     催化剂排出管
11      催化剂料斗              12     反应器内筒
13      反应器外壳
具体实施方式
在本文中所披露的范围的端点和任何值都不限于该精确的范围或值，这些范围或值应当理解为包含接近这些范围或值的值。对于数值范围来说，各个范围的端点值之间、各个范围的端点值和单独的点值之间，以及单独的点值之间可以彼此组合而得到一个或多个新的数值范围，这些数值范围应被视为在本文中具体公开。
在本发明中，在未作相反说明的情况下，使用的方位词如“上、下”通常是指参考附图所示的上、下；“内、外”是指相对于各部件本身的轮廓的内、外。
本发明所述的烟气脱硝方法包括：在脱硝反应器中，在氨气的存在下，使烟气从下到上穿过多个催化剂床层，进行脱硝反应，其中，每个催化剂床层包括催化剂支撑部件和堆积在该催化剂支撑部件上的颗粒状脱硝催化剂，在相邻的上下两个催化剂床层中，两个催化剂支撑部件的运行方向相反，上层颗粒状脱硝催化剂随催化剂支撑部件移动至催化剂支撑部件末端，依靠重力自由落至下层催化剂支撑部件的起始端。
在所述脱硝反应器中，为了确保烟气能够穿过催化剂床层，所述催化剂支撑部件上要求带有开孔。在优选情况下，所述催化剂支撑部件为网状的。
根据本发明的一种优选实施方式，所述催化剂支撑部件为网状传送带，在这种情况下，网状传送带上的颗粒状脱硝催化剂可以从一端传送到另一 端，并且通过控制传送带的传送速度来灵活调节颗粒状脱硝催化剂的移动速度。另外，每个催化剂床层都单独设置网状传送带，使得各个催化剂床层的移动速度和厚度能够独立地控制，床层之间的相互干扰较小。进一步优选地，所述多个催化剂床层从上到下相互平行排布，也即作为催化剂支撑部件的多个网状传送带相互平行排布，这样可以确保单个网状传送带上的颗粒状脱硝催化剂分布均匀，从而更有利于充分利用催化剂活性。
在本发明中，所述网状传送带可以为本领域常规的金属网状传送带，优选为不锈钢网状传送带。所述网状传送带的网孔大小要求小于颗粒状脱硝催化剂的颗粒尺寸，以保证催化剂颗粒不从网孔掉落。通常情况下，所述网状传送带上的网孔尺寸可以为0.1～3mm，优选为1.5～2.5mm。
在本发明中，所述催化剂床层的个数和宽度可以根据实际需要及反应器大小进行选择。所述催化剂床层的个数优选为3～10，更优选为3～8。在催化剂支撑部件(优选网状传送带)的宽度方向上，催化剂支撑部件与反应器器壁之间的间隙可以为2～50mm，优选为2～5mm。
在本发明中，相邻两层催化剂床层之间的垂直距离可以为50-2000mm。在优选情况下，当所述催化剂支撑部件为网状传送带，且多个催化剂支撑部件之间相互平行排布时，相邻两层催化剂床层之间的垂直距离为1200～2000mm，优选为1400～1600mm。
在本发明中，当所述催化剂支撑部件为网状传送带，且多个催化剂支撑部件之间相互平行排布时，所述颗粒状脱硝催化剂在网状传送带上的堆积高度可以为50～500mm，优选为200～300mm。
在本发明中，当所述催化剂支撑部件为网状传送带时，所述网状传送带的传送速度可以为0.1-10mm/s，优选为0.5～2mm/s。
在本发明所述的方法中，所述烟气可以选自燃煤电厂烟气、FCC再生烟气、炼油厂工艺炉烟气和化工炉烟气(如乙烯裂解炉烟气等)，一般含有NOx、SOx以及杂质。所述杂质一般为粉尘、水、CO2和O2等。进入所述脱 硝反应器的烟气的温度可以为300-420℃，优选为340～400℃。
在本发明所述的方法中，氨气可以以本领域常规的方式引入，例如可以以氨气的形式引入。在优选情况下，氨气以含有氨气和空气的混合气的形式引入，且氨气在所述混合气中的体积浓度可以为0.5-10％，优选为3-7％。
在本发明所述的方法中，所述含有氨气和空气的混合气优选在所述催化剂床层的下方注入，也即所述含有氨气和空气的混合气在位于最下方的催化剂床层的下方注入。所述含有氨气和空气的混合气可以通过设置于烟气入口处的喷氨格栅注入。所述喷氨格栅可以为本领域常规的喷氨格栅，但是，在常规的SCR工艺和设备中，喷氨格栅性能要求保证氨气的浓度分布偏差小于5％，而本发明中对氨气的浓度分布偏差范围要求并不苛刻，只要氨气的浓度分布偏差达到30％以下即可，例如可以为5％～30％，优选为12％～18％。
在本发明所述的方法中，氨气与烟气中以氮原子计的氮氧化物的摩尔比可以为0.9-1.15：1。
在本发明所述的方法中，所述烟气的流速可以为2-15m/s，优选为4-10m/s；反应停留时间可以为0.5-20s。
在本发明所述的方法中，优选地，所述颗粒状脱硝催化剂具有如下性质：颗粒尺寸为3～6mm，堆积密度为0.2～0.8g/cm3，比表面积为80～120m2/g。所述脱硝催化剂的组成可以为本领域常用的组分，以催化剂重量计，各组分含量以氧化物计为：0.01～1重量％的V、88～99重量％的Ti、0.1～10重量％的W和0.01～1重量％的Mo。所述脱硝催化剂的制备过程可以为：物料干混-捏合-过滤-练泥-挤出造粒-阴干-干燥-焙烧-成品，其中催化剂制备各步骤所涉及的条件为本领域技术人员熟知，挤出造粒可以根据催化剂颗粒尺寸配备不同型号的捏合式挤出造粒机。
在一种具体实施方式中，本发明所述的烟气脱硝方法包括：烟气从脱硝反应器底部进入，含有氨气和空气的混合气经喷氨格栅加注到烟气中， 气流自下而上穿过多层水平交错排列的催化剂床层，进行脱硝反应脱除NOx，烟气中的粉尘同时被催化剂床层过滤除尘，经过脱硝与除尘的烟气从反应器顶部排出，进行下一步脱硫处理；其中催化剂床层由网状传送带和传送带上堆积的颗粒状脱硝催化剂组成，相邻上下两层网状传送带的运行方向相反，上层颗粒状脱硝催化剂随传送带移动至传送带末端，依靠重力自由落至下层传送带的运行方向的起始端，颗粒状脱硝催化剂在最后一层传送带的末端落入催化剂回收装置进行回收。
所述颗粒状脱硝催化剂经催化剂加入管加入脱硝反应器顶部的第一层网状传送带上，催化剂掉落在网状传送带上堆积形成催化剂床层。
所述催化剂回收装置一般采用催化剂储罐、催化剂料斗等常用设备；回收的催化剂颗粒进行筛分脱除粉尘与破碎的催化剂颗粒后可重复使用。
如图1所示，本发明所述的脱硝反应器包括反应器壳体13和在反应器壳体13内从上到下水平交错排列的多个催化剂支撑部件，并且所述多个催化剂支撑部件被配置为：在相邻的上下两个催化剂支撑部件中，两个催化剂支撑部件的运行方向相反，催化剂支撑部件上的颗粒状脱硝催化剂随着催化剂支撑部件移动至催化剂支撑部件末端，依靠重力自由落至下层催化剂支撑部件的起始端。
在一种优选实施方式，所述催化剂支撑部件为网状传送带。进一步优选地，所述多个催化剂床层从上到下相互平行排布，也即作为催化剂支撑部件的多个网状传送带相互平行排布。
所述网状传送带可以为本领域常规的金属网状传送带，优选为不锈钢网状传送带。所述网状传送带的网孔大小要求小于颗粒状脱硝催化剂的颗粒尺寸，以保证催化剂颗粒不从网孔掉落。通常情况下，所述网状传送带上的网孔尺寸可以为0.1～3mm，优选为1.5～2.5mm。
在本发明中，如图1所示，所述脱硝反应器还可以包括内筒12，所述内筒12贯穿所述反应器壳体13的上端和下端，所述网状传送带横向贯穿 所述内筒12。所述网状传送带通常采用电机驱动，由传送带驱动轮带动传送带转动。为了避免所述网状传送带的驱动轮9处于高温下被损坏，优选地，将所述网状传送带的驱动轮9设置在所述内筒12和所述反应器壳体13之间的腔室内。进一步地，所述内筒12的下端设置有烟气入口，上端设置有烟气出口，且喷氨部件6设置在烟气入口处。
在本发明所述的脱硝反应器中，最上方的催化剂支撑部件的头部(起始端)对应设置有催化剂加入管7，用于加入颗粒状脱硝催化剂，最下方的催化剂支撑部件的尾部(末端)对应设置有催化剂收集料斗，用于回收催化剂颗粒。
下面通过具体实施例对本发明的烟气的脱硝除尘方法做详细说明，但并不因此限制本发明。
实施例及比较例中的v％为体积分数，wt％为质量分数。
催化剂比表面积及孔容采用ASAP 2420大型多站式全自动比表面积及孔隙分析仪测量。
测试例中的氨浓度分布偏差为设计的一个指标，可通过实验与计算获得；计算采用CFD流体力学计算软件对喷氨格栅结构进行模拟来确定氨浓度分布偏差，以截面上氨浓度平均值及分布范围计算。
逃逸氨量及粉尘含量采用CEMS烟气在线分析仪进行测量。
实施例1-5中使用的脱销反应器的结构示意图如图1所示，该脱销反应器包括反应器外壳13、反应器内筒12、喷氨格栅6、催化剂加入管7、网状传送带8、传送带驱动轮9、催化剂排出管10和催化剂料斗11；其中反应器外壳13和反应器内筒12之间为反应器外密封腔，喷氨格栅6在反应器内筒12的底部入口处，传送带驱动轮9在反应器外密封腔内，传送带8贴在传送带驱动轮9上，横穿反应器内筒12，催化剂加入管7在反应器外密封腔顶部，催化剂加入管7底部出口正对传送带一端，催化剂料斗11在 反应器外密封腔底部，催化剂排出管10在催化剂料斗11底部。
所述烟气脱硝的操作过程如下：颗粒状脱硝催化剂3通过催化剂加入管7加注到第一层网状传送带8上堆积形成床层，传送带驱动轮9带动传送带8上的床层运动，床层穿过反应器内筒12，进入反应器外密封腔，在重力作用下落到下一个传送带上，形成床层，并在传送带驱动轮9驱动下向相反的方向运动，按照上述运行方式，形成连续不断的运行的传送带床层；烟气1从脱硝床反应器底部进入，含氨气的混合气2通过喷氨格栅6加注到烟气1中，两者混合自下而上穿过上述的传送带床层，进行脱硝反应，脱除NOx，同时粉尘被床层过滤下来，颗粒状脱硝催化剂在最后一层传送带的末端落入催化剂料斗11中，脱硝后的颗粒状脱硝催化剂通过催化剂排出管10排出反应器进行回收，脱除NOx与粉尘的净化气5从反应器顶部排出。
实施例1
FCC再生烟气流量为15万Nm3/h，温度为650℃，压力为10kPa，NOx浓度为600mg/Nm3，SO2浓度为1000mg/Nm3，SO3浓度为20mg/Nm3，粉尘含量为200mg/Nm3。NOx排放标准为200mg/Nm3。
颗粒状脱硝催化剂的催化剂微粒活性组分为V的氧化物、Ti的氧化物、W的氧化物和Mo的氧化物，活性组分以氧化物计，催化剂为球形颗粒，质量比例如下：V为0.01wt％，Ti为99wt％，W为0.1wt％，Mo为0.02wt％；其催化剂颗粒大小为5mm；堆积密度为0.68g/cm3，比表面积为40m2/g。
首先FCC再生烟气经过锅炉取热，温度由650℃降低至SCR脱硝反应温度400℃；原料供给区提供的含有氨气的混合气流量为1120Nm3/h，其中氨气浓度为4v％，反应器的内密封腔大小为长8m×宽6m×高8m；反应时间为0.5s，设置3层传送带，每个传送带上催化剂床层高度为300mm，传送带尺寸为长9m×宽5.8m，选用不锈钢网状传送带，空隙直径为3mm，驱动 轮直径300mm，上下两层传送带之间空高1300mm，留有足够的检修空间。经过脱硝反应后，可保证净化烟气的NOx含量为100mg/Nm3，粉尘含量小于10mg/Nm3，满足重点控制地区环保要求；然后烟气通过脱硫除尘脱除烟气中的SO2与粉尘，即可通过烟囱排放。
实施例2
FCC再生烟气流量、温度、压力同实施例1，NOx浓度为2000mg/Nm3，SO2浓度为2000mg/Nm3，SO3浓度为200mg/Nm3，粉尘含量为400mg/Nm3。NOx排放标准为100mg/Nm3。
催化剂组成同实施例1。
首先FCC再生烟气经过锅炉取热，温度由650℃降低至SCR脱硝反应温度300℃；原料供给区提供的含有氨气的混合气流量为1000Nm3/h，其中氨气浓度为3v％；反应器的内密封腔大小为长8m×宽6m×高15m；反应时间为2s，设置10层传送带，每个传送带上催化剂床层高度为500mm，传送带尺寸为长9m×宽5.8m，选用不锈钢网状传送带，空隙直径为3mm，驱动轮直径300mm，上下两层传送带之间空高1500mm，留有足够的检修空间。经过脱硝反应后，可保证净化烟气的NOx含量为100mg/Nm3，粉尘含量小于5mg/Nm3，满足重点控制地区环保要求；然后烟气通过脱硫除尘脱除烟气中的SO2与粉尘，即可通过烟囱排放。
实施例3
FCC再生烟气流量、温度、压力同实施例1，NOx浓度为300mg/Nm3，SO2浓度为600mg/Nm3，SO3浓度为10mg/Nm3，粉尘含量为100mg/Nm3。NOx排放标准为200mg/Nm3。
催化剂组成同实施例1，催化剂颗粒直径为3mm。
首先FCC再生烟气经过锅炉取热，温度由650℃降低至SCR脱硝反应 温度300℃；原料供给区提供的含有氨气的混合气流量为2000Nm3/h，其中氨气浓度为2.8v％；反应器的内密封腔大小为长8m×宽6m×高6m；反应时间为0.5s，设置3层传送带，每个传送带上催化剂床层高度为50mm，传送带尺寸为长9m×宽5.8m，选用不锈钢网状传送带，空隙直径为2.5mm，驱动轮直径500mm，上下两层传送带之间空高2000mm，留有足够的检修空间。经过脱硝反应后，可保证净化烟气的NOx含量为100mg/Nm3，粉尘含量小于5mg/Nm3，满足重点控制地区环保要求；然后烟气通过脱硫除尘脱除烟气中的SO2与粉尘，即可通过烟囱排放。
实施例4
FCC再生烟气流量、温度、压力同实施例1，NOx浓度为800mg/Nm3，SO2浓度为3600mg/Nm3，SO3浓度为1000mg/Nm3，粉尘含量为200mg/Nm3。NOx排放标准为100mg/Nm3。
催化剂组成同实施例1，催化剂颗粒直径为6mm。
首先FCC再生烟气经过锅炉取热，温度由650℃降低至SCR脱硝反应温度300℃；原料供给区提供的含有氨气的混合气流量为2000Nm3/h，其中氨气浓度为2.8v％；反应器的内密封腔大小为长8m×宽6m×高6m；反应时间为0.5s，设置3层传送带，每个传送带上催化剂床层高度为500mm，传送带尺寸为长9m×宽5.8m，选用不锈钢网状传送带，空隙直径为0.2mm，驱动轮直径300mm，上下两层传送带之间空高2000mm，留有足够的检修空间。经过脱硝反应后，可保证净化烟气的NOx含量为100mg/Nm3，粉尘含量小于5mg/Nm3，满足重点控制地区环保要求；然后烟气通过脱硫除尘脱除烟气中的SO2与粉尘，即可通过烟囱排放。
实施例5
FCC再生烟气流量为8500Nm3/h，温度为650℃，压力为10kPa，NOx浓度 为600mg/Nm3，SO2浓度为1000mg/Nm3，SO3浓度为20mg/Nm3，粉尘含量为200mg/Nm3。NOx排放标准为200mg/Nm3。
催化剂组成同实施例1，催化剂颗粒直径为5mm。
首先FCC再生烟气经过锅炉取热，温度由650℃降低至SCR脱硝反应温度400℃；原料供给区提供的含有氨气的混合气流量为70Nm3/h，其中氨气浓度为4v％，反应器的内密封腔大小为长1.2m×宽1m×高1.2m；反应时间为0.5s，设置3层传送带，每个传送带上催化剂床层高度为40mm，传送带尺寸为长1.5m×宽0.8m，选用不锈钢网状传送带，空隙直径为3mm，驱动轮直径50mm，上下两层传送带之间空高200mm。经过脱硝反应后，可保证净化烟气的NOx含量为100mg/Nm3，粉尘含量小于10mg/Nm3，满足重点控制地区环保要求；然后烟气通过脱硫除尘脱除烟气中的SO2与粉尘，即可通过烟囱排放。
对比例1
同实施例1，只是反应器替换为传统的固定床反应器，催化剂采用蜂窝状催化剂，其组分同实施例1，采用模块化装填，单催化剂模块高度为1m，反应器大小为4.4m×4.6m，催化剂装填三层，首先FCC再生烟气经过锅炉取热，温度由650℃降低至SCR脱硝反应温度350℃；原料供给区提供的含有氨气的混合气流量为1000Nm3/h，其中氨气浓度为3v％。含有氨气的混合气在距离反应器入口一定距离的上游烟道加入，经过喷氨格栅的混合扩散后，保证反应器入口烟气中的氨气浓度偏差小于5％，再进入SCR反应器反应，经过脱硝反应后，可保证净化烟气的NOx含量为100mg/Nm3，脱硝后的烟气继续进入下游装置进行换热、脱硫除尘，满足重点控制地区环保要求。
对比例2
同实施例2，只是反应器替换为传统的固定床反应器。
对比例3
同实施例3，只是反应器替换为传统的固定床反应器。
对比例4
同实施例4，只是反应器替换为传统的固定床反应器。
对比例5
同实施例5，只是反应器替换为传统的固定床反应器，催化剂采用蜂窝状催化剂，其组分同实施例5，采用模块化装填，单催化剂模块高度为150mm，反应器大小为700mm×800mm，催化剂装填三层，首先FCC再生烟气经过锅炉取热，温度由650℃降低至SCR脱硝反应温度350℃；原料供给区提供的含有氨气的混合气流量为60Nm3/h，其中氨气浓度为3v％。含有氨气的混合气在距离反应器入口一定距离的上游烟道加入，经过喷氨格栅的混合扩散后，保证反应器入口烟气中的氨气浓度偏差小于5％，再进入SCR反应器反应，经过脱硝反应后，可保证净化烟气的NOx含量为100mg/Nm3，脱硝后的烟气继续进入下游装置进行换热、脱硫除尘，满足重点控制地区环保要求。
对比例6
仍采用对比例1中的反应器，由于烟气中NOx浓度提高幅度太大，保证氨逃逸﹤3mgNm3，脱硝后NOx浓度为1000～1300mg/Nm3，不能达标排放，并且粉尘仍需进入脱硫除尘系统进行处理。
测试例
实施例1～4及对比例1～4的运行周期及催化剂用量见表1，实施例5和对比例5的运行周期及催化剂用量见表2，喷氨格栅对氨分布的要求见表3。
表1
Figure PCTCN2017099794-appb-000001
表2
  运行周期 催化剂用量(运行1年)
实施例5 1年以上 0.12m3
对比例5 0.5年 0.25m3
表3
  氨浓度分布偏差 逃逸氨量(mg/Nm3) 净化烟气粉尘含量(mg/Nm3)
实施例1 ＜10％ 0 ﹤10
实施例2 ＜30％ 0 ﹤5
实施例3 ＜7％ 0 ﹤5
实施例4 ＜15％ 0 ﹤5
实施例5 ＜15％ 0 ﹤5
对比例1 ＜5％ 1.0 200
对比例2 ＜5％ 1.0 200
对比例3 ＜5％ 1.2 400
对比例4 ＜5％ 0.9 100
对比例5 ＜5％ 1.0 200
由上表1-3的数据可以看出，按照本发明所述的烟气脱硝方法可以提高催化剂的利用率，大大降低了催化剂的用量；而且，本发明所述的方法不会出现氨逃逸，对床层初始氨分布的均匀性要求不高；另外，本发明所述的方法可以获得较好的除尘效果。
以上详细描述了本发明的优选实施方式，但是，本发明并不限于此。 在本发明的技术构思范围内，可以对本发明的技术方案进行多种简单变型，包括各个技术特征以任何其它的合适方式进行组合，这些简单变型和组合同样应当视为本发明所公开的内容，均属于本发明的保护范围。

Claims (16)

  1. 一种烟气脱硝方法，该方法包括：在脱硝反应器中，在氨气的存在下，使烟气从下到上穿过多个催化剂床层，进行脱硝反应，其中，每个催化剂床层包括催化剂支撑部件和堆积在该催化剂支撑部件上的颗粒状脱硝催化剂，在相邻的上下两个催化剂床层中，两个催化剂支撑部件的运行方向相反，上层颗粒状脱硝催化剂随催化剂支撑部件移动至催化剂支撑部件末端，依靠重力自由落至下层催化剂支撑部件的起始端。
  2. 一种烟气脱硝方法，该方法包括：烟气从脱硝反应器底部进入，含氨气的混合气经喷氨格栅加注到烟气中，气流自下而上穿过多层水平交错排列的催化剂床层，进行脱硝反应脱除NOX，烟气中的粉尘同时被催化剂床层过滤除尘，经过脱硝与除尘的烟气从反应器顶部排出，进行下一步脱硫处理；其中催化剂床层由网状传送带和传送带上堆积的颗粒状脱硝催化剂组成，相邻上下两层传送带的运行方向相反，上层颗粒状脱硝催化剂随传送带移动传送带末端，依靠重力自由落至下层传送带的运行方向的起始端，颗粒状脱硝催化剂在最后一层传送带的末端落入催化剂回收装置进行回收。
  3. 根据权利要求1所述的方法，其中，所述催化剂支撑部件为网状传送带。
  4. 根据权利要求2或3所述的方法，其中，所述网状传送带上的网孔尺寸为0.1-3mm。
  5. 根据权利要求1-4中任意一项所述的方法，其中，所述催化剂床层的个数为3-10。
  6. 根据权利要求1-5中任意一项所述的方法，其中，相邻两个催化剂床层之间的垂直距离为1200-2000mm。
  7. 根据权利要求1-6中任意一项所述的方法，其中，进入所述脱硝反应器的烟气的温度为300-420℃。
  8. 根据权利要求1-7中任意一项所述的方法，其中，氨气以含有氨气和空气的混合气的形式引入，且氨气在所述混合气中的体积浓度为0.5-10％。
  9. 根据权利要求8所述的方法，其中，所述混合气在所述催化剂床层的下方注入。
  10. 根据权利要求1-9中任意一项所述的方法，其中，氨气与烟气中以氮原子计的氮氧化物的摩尔比为0.9-1.15：1。
  11. 根据权利要求1-10中任意一项所述的方法，其中，所述烟气的流速为2-15m/s，反应停留时间为0.5-20s。
  12. 根据权利要求1-11中任意一项所述的方法，其中，所述颗粒状脱硝催化剂的颗粒尺寸为3-6mm，堆积密度为0.2-0.8g/cm3，比表面积为80-120m2/g。
  13. 根据权利要求1-12中任意一项所述的方法，其中，所述颗粒状脱硝催化剂在所述催化剂支撑部件上的堆积高度为50-500mm。
  14. 根据权利要求2-13中任意一项所述的方法，其中，所述网状传送带的传送速度为0.1-10mm/s。
  15. 根据权利要求1所述的方法，其中，所述氨气通过喷氨格栅注入脱硝反应器中。
  16. 根据权利要求2或15所述的方法，其中，所述喷氨格栅喷氨的要求为：氨气的浓度分布偏差范围为5％～30％。
PCT/CN2017/099794 2016-08-31 2017-08-31 烟气脱硝方法 WO2018041171A1 (zh)

Priority Applications (3)

Application Number Priority Date Filing Date Title
US16/329,111 US11213788B2 (en) 2016-08-31 2017-08-31 Method of flue gas denitrification
EP17845473.2A EP3498360A4 (en) 2016-08-31 2017-08-31 PROCESS FOR THE DENITRATION OF COMBUSTION GAS
JP2019511754A JP6770176B2 (ja) 2016-08-31 2017-08-31 排煙脱硝方法

Applications Claiming Priority (2)

Application Number Priority Date Filing Date Title
CN201610767731.8A CN107789983A (zh) 2016-08-31 2016-08-31 一种烟气脱硝方法及脱硝反应器
CN201610767731.8 2016-08-31

Publications (1)

Publication Number Publication Date
WO2018041171A1 true WO2018041171A1 (zh) 2018-03-08

Family

ID=61301750

Family Applications (1)

Application Number Title Priority Date Filing Date
PCT/CN2017/099794 WO2018041171A1 (zh) 2016-08-31 2017-08-31 烟气脱硝方法

Country Status (5)

Country Link
US (1) US11213788B2 (zh)
EP (1) EP3498360A4 (zh)
JP (1) JP6770176B2 (zh)
CN (1) CN107789983A (zh)
WO (1) WO2018041171A1 (zh)

Cited By (1)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN110052162A (zh) * 2019-05-15 2019-07-26 北京国电龙源环保工程有限公司 Scr脱硝反应器中的脱硝催化剂系统及其脱硝方法

Families Citing this family (6)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN110052157A (zh) * 2019-04-01 2019-07-26 浙江菲达环保科技股份有限公司 一种基于改性赤泥的Hg、SO3协同脱除装置
GB2593207B (en) * 2020-03-19 2022-06-01 Narasimhamurthy Prakashkumar A continuous type process method to increase the rate of reaction between solids, liquids, and gasses per area of the land occupied by two reactors
CN112370964B (zh) * 2020-10-28 2023-03-10 国网山东省电力公司电力科学研究院 一种改善燃煤电站锅炉scr系统综合性能的方法及系统
CN113617218A (zh) * 2021-09-03 2021-11-09 唐山鑫联环保科技有限公司 一种脱硝药剂胶囊及其制备方法与用途
CN114887477B (zh) * 2022-05-17 2023-08-22 广东国鑫实业股份有限公司 一种烧结机脱硝系统
CN115475508B (zh) * 2022-09-28 2024-04-30 天津中材工程研究中心有限公司 Sncr旋风筒喷枪的插入位置判断方法及自动调整系统

Citations (6)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
JPS5466380A (en) * 1977-11-07 1979-05-28 Hitachi Plant Eng & Constr Co Ltd Treating apparatus for exhaust gas containing hydrogen chloride
JPS60122028A (ja) * 1983-12-02 1985-06-29 Gadelius Kk ガス吸収装置
JPH0952015A (ja) * 1995-08-18 1997-02-25 Shinko Pantec Co Ltd 溶剤処理装置及び溶剤処理方法
CN2350119Y (zh) * 1998-10-30 1999-11-24 中国科学院化工冶金研究所 多层流化移动床反应器
CN1911491A (zh) * 2006-08-03 2007-02-14 长春惠工净化工业有限公司 移动床烟气脱硫脱硝除尘工艺
CN105859097A (zh) * 2016-05-24 2016-08-17 华南理工大学 一种结合热电厂余热利用污泥干燥方法及其装置

Family Cites Families (13)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
BE636744A (zh) * 1962-08-29
FR1578125A (zh) * 1968-06-19 1969-08-14
JPS51129871A (en) * 1975-05-08 1976-11-11 Asahi Glass Co Ltd Adust removing and denitration process by catalytic reaction apparatus
JPS5850769B2 (ja) 1976-06-04 1983-11-12 川崎重工業株式会社 集塵・脱硝装置
JPS55106514A (en) * 1979-02-08 1980-08-15 Tomoaki Tanaka Air cleaner
US5494500A (en) * 1990-11-26 1996-02-27 The Social Welfare Foundation Hokkaido Rehabily Activated carbon, production thereof and adsorption using activated carbon
JPH04277006A (ja) * 1991-03-01 1992-10-02 Sumitomo Heavy Ind Ltd ゴミ焼却炉の乾式排ガス処理装置
DE4140612C1 (zh) * 1991-12-10 1993-06-17 Fa. Carl Freudenberg, 6940 Weinheim, De
KR930017610A (ko) * 1992-02-17 1993-09-20 셀 인터나쵸 나아레 레사아치 마아츠샤피 비이부이 화학적 전환을 수행하기 위한 방법 및 장치
JPH08182918A (ja) * 1994-12-28 1996-07-16 Kawasaki Heavy Ind Ltd 微粉炭素による流動層脱硝方法及び装置
JPH11156189A (ja) * 1997-11-28 1999-06-15 Unitika Ltd 排ガス処理方法
CN101613101B (zh) * 2008-03-06 2012-02-01 项瑞芳 具有净化自身尾气功能的化学法制活性炭方法
EP2789677A1 (en) * 2013-04-12 2014-10-15 Kymi Baltic Consulting Oü Torrefaction plant, its operation and maintenance

Patent Citations (6)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
JPS5466380A (en) * 1977-11-07 1979-05-28 Hitachi Plant Eng & Constr Co Ltd Treating apparatus for exhaust gas containing hydrogen chloride
JPS60122028A (ja) * 1983-12-02 1985-06-29 Gadelius Kk ガス吸収装置
JPH0952015A (ja) * 1995-08-18 1997-02-25 Shinko Pantec Co Ltd 溶剤処理装置及び溶剤処理方法
CN2350119Y (zh) * 1998-10-30 1999-11-24 中国科学院化工冶金研究所 多层流化移动床反应器
CN1911491A (zh) * 2006-08-03 2007-02-14 长春惠工净化工业有限公司 移动床烟气脱硫脱硝除尘工艺
CN105859097A (zh) * 2016-05-24 2016-08-17 华南理工大学 一种结合热电厂余热利用污泥干燥方法及其装置

Non-Patent Citations (1)

* Cited by examiner, † Cited by third party
Title
See also references of EP3498360A4 *

Cited By (1)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
CN110052162A (zh) * 2019-05-15 2019-07-26 北京国电龙源环保工程有限公司 Scr脱硝反应器中的脱硝催化剂系统及其脱硝方法

Also Published As

Publication number Publication date
US20190224621A1 (en) 2019-07-25
JP2019531881A (ja) 2019-11-07
EP3498360A4 (en) 2020-04-01
EP3498360A1 (en) 2019-06-19
CN107789983A (zh) 2018-03-13
JP6770176B2 (ja) 2020-10-14
US11213788B2 (en) 2022-01-04

Similar Documents

Publication Publication Date Title
WO2018041171A1 (zh) 烟气脱硝方法
KR102003148B1 (ko) 코크스로 연도 가스의 탈황 및 탈질 통합정화공정 및 장치
CN102895839B (zh) 一种集成净化烟气的装置及方法
CN208356505U (zh) 一种用于催化裂化装置烟气的半干法脱硫除尘系统
CN103007681A (zh) 活性炭烟气脱硫及再生装置和方法
CN108855023A (zh) 一种烟气脱硝催化剂的制备方法及脱硝工艺
CN107469600A (zh) 一种复合床层净化塔
CN109126815B (zh) 一种低温脱硝催化剂及应用
CN107998859A (zh) 一种脱硫脱硝除尘一体化湿法超净处理烟气的工艺
CN102309920A (zh) 一种脱除FCC烟气中NOx和SOx的方法
CN108786438A (zh) 一种烟气脱硝方法及反应器
CN103768930A (zh) Fcc再生烟气防尘脱硝反应器
CN102688731B (zh) 一种流态化反应器及使用该反应器的烟气催化脱硝装置
CN207769527U (zh) 一种脱硫脱硝除尘一体化湿法超净处理烟气的装置
CN204952626U (zh) 气相氧化系统以及应用该系统的烟气净化设备
CN108855079B (zh) 一种烟气脱硝催化剂及其制备方法及脱硝工艺
CN111298593B (zh) 一种活性焦吸附塔及活性焦吸附净化烟气的方法
CN108786403B (zh) 一种脱硝反应器及脱硝脱硫方法
CN207254073U (zh) 一种复合床层净化塔
CN107875824A (zh) 一种脱硫脱硝除尘一体化湿法超净处理烟气的装置
CN109126453B (zh) 一种低温脱硝工艺
CN202844887U (zh) 一种集成净化烟气的装置
CN102698594B (zh) 一种垃圾焚烧烟气预热催化脱硝反应装置
CN108786439A (zh) 一种烟气脱硝方法及脱硝反应器
CN107569937A (zh) 一种焚烧炉烟气脱硫脱硝装置

Legal Events

Date Code Title Description
121 Ep: the epo has been informed by wipo that ep was designated in this application

Ref document number: 17845473

Country of ref document: EP

Kind code of ref document: A1

ENP Entry into the national phase

Ref document number: 2019511754

Country of ref document: JP

Kind code of ref document: A

NENP Non-entry into the national phase

Ref country code: DE

ENP Entry into the national phase

Ref document number: 2017845473

Country of ref document: EP

Effective date: 20190312