WO2011009283A1 - 煤层气脱氧催化剂机器制备方法和应用 - Google Patents

煤层气脱氧催化剂机器制备方法和应用 Download PDF

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WO2011009283A1
WO2011009283A1 PCT/CN2010/000528 CN2010000528W WO2011009283A1 WO 2011009283 A1 WO2011009283 A1 WO 2011009283A1 CN 2010000528 W CN2010000528 W CN 2010000528W WO 2011009283 A1 WO2011009283 A1 WO 2011009283A1
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catalyst
gas
coalbed methane
deoxidation
oxygen
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English (en)
French (fr)
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王树东
王胜
袁中山
张纯希
倪长军
李德意
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Dalian Institute of Chemical Physics of CAS
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Dalian Institute of Chemical Physics of CAS
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Priority claimed from CN200910012669A external-priority patent/CN101613627B/zh
Priority claimed from CN200910012670A external-priority patent/CN101664679B/zh
Application filed by Dalian Institute of Chemical Physics of CAS filed Critical Dalian Institute of Chemical Physics of CAS
Priority to US12/737,342 priority Critical patent/US20120003132A1/en
Publication of WO2011009283A1 publication Critical patent/WO2011009283A1/zh
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    • 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
    • B01J23/00—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00
    • B01J23/38—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00 of noble metals
    • B01J23/54—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00 of noble metals combined with metals, oxides or hydroxides provided for in groups B01J23/02 - B01J23/36
    • B01J23/56—Platinum group metals
    • B01J23/63—Platinum group metals with rare earths or actinides
    • 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
    • B01J23/00—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00
    • B01J23/002—Mixed oxides other than spinels, e.g. perovskite
    • 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
    • B01J23/00—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00
    • B01J23/38—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00 of noble metals
    • B01J23/54—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group B01J21/00 of noble metals combined with metals, oxides or hydroxides provided for in groups B01J23/02 - B01J23/36
    • B01J23/56—Platinum group metals
    • B01J23/58—Platinum group metals with alkali- or alkaline earth metals
    • 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
    • B01J37/00—Processes, in general, for preparing catalysts; Processes, in general, for activation of catalysts
    • B01J37/02—Impregnation, coating or precipitation
    • B01J37/024—Multiple impregnation or coating
    • B01J37/0242—Coating followed by impregnation
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10L—FUELS NOT OTHERWISE PROVIDED FOR; NATURAL GAS; SYNTHETIC NATURAL GAS OBTAINED BY PROCESSES NOT COVERED BY SUBCLASSES C10G OR C10K; LIQUIFIED PETROLEUM GAS; USE OF ADDITIVES TO FUELS OR FIRES; FIRE-LIGHTERS
    • C10L3/00—Gaseous fuels; Natural gas; Synthetic natural gas obtained by processes not covered by subclass C10G, C10K; Liquefied petroleum gas
    • C10L3/06—Natural gas; Synthetic natural gas obtained by processes not covered by C10G, C10K3/02 or C10K3/04
    • C10L3/10—Working-up natural gas or synthetic natural gas
    • C10L3/101—Removal of contaminants
    • 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
    • B01J21/00—Catalysts comprising the elements, oxides, or hydroxides of magnesium, boron, aluminium, carbon, silicon, titanium, zirconium, or hafnium
    • B01J21/06—Silicon, titanium, zirconium or hafnium; Oxides or hydroxides thereof
    • B01J21/066—Zirconium or hafnium; Oxides or hydroxides thereof
    • 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
    • B01J2523/00—Constitutive chemical elements of heterogeneous catalysts

Definitions

  • the invention belongs to the field of chemistry, and particularly provides a coal bed gas deoxidation catalyst, a preparation method thereof and a process for catalytic deoxidation of an oxygen-containing coal bed gas.
  • Coalbed methane is a combustible gas adsorbed in coal seams, and its main component is high purity methane. Because coalbed methane does not contain harmful impurities such as sulfur (H 2 S) in conventional natural gas, and does not contain toxic substances such as benzene, mercury and lead, CBM is recognized as a high-quality clean energy in the world.
  • harmful impurities such as sulfur (H 2 S) in conventional natural gas, and does not contain toxic substances such as benzene, mercury and lead
  • CBM is recognized as a high-quality clean energy in the world.
  • the quality of coalbed methane produced by both ground drilling and downhole drainage is very different, and its utilization is also very different.
  • the C-concentration of coalbed methane extracted by the former is more than 90%. It is suitable for input into natural gas pipeline systems and can be used as fuel for power generation, industrial fuels, vehicle fuels, chemical raw materials and residential fuels.
  • the C-concentration of the coalbed methane recovered by the latter is 30-50% on average, and the pressure is also low. Usually, it can only be used as a domestic fuel, and most of it is burned and emptied, resulting in great energy waste. At the same time, the contradiction between supply and demand of natural gas resources in China has become increasingly serious. Therefore, such medium and low concentrations of coalbed methane must be processed and purified, pressurized for long distance transportation and use.
  • the CBM purification technology refers to the separation of N 2 or air from C3 ⁇ 4, so that the methane content in the coalbed methane is correspondingly increased, thereby increasing the calorific value of the coalbed methane and reducing the transportation cost.
  • CBM purification technologies mainly include low temperature cryogenic separation, pressure swing adsorption and membrane separation.
  • Patent CN1952569A and CN1908559A disclose a low-temperature two-stage rectification liquefaction separation process containing air coalbed methane, the liquefaction and separation are carried out at low temperature, and the products of liquefied natural gas are produced. The purity can reach more than 99%.
  • the separation process as the concentration of methane increases, the oxygen content of the exhaust gas is also concentrated. Inevitably, one stage is just the range of combustion and explosion of methane, and there is a great safety risk.
  • a safer separation and purification scheme is to selectively remove oxygen from the coalbed methane by catalytic combustion and then purify it.
  • This method can reduce the oxygen content in the coalbed methane to less than 0.5%, and completely eliminate the safety hazard of the operation process.
  • the desulfurization modes of coalbed methane mainly include catalytic deoxidation (ZL02113628.9, CN101139239A, etc.), coke combustion method (ZL02113627.0, CN1919986A) and the like.
  • the coalbed methane coke combustion deoxidation process can effectively remove the o 2 in the oxygen-containing coalbed methane
  • the process uses coke as a fuel (such as the use of anthracite instead of coke to bring about S0 2 emissions), and the energy consumption is higher;
  • the coke and dust removal process is also relatively complicated; the higher reaction temperature not only puts higher requirements on the reactor material, but also may cause side reactions such as C3 ⁇ 4 pyrolysis and reforming, which reduces the C3 ⁇ 4 recovery rate in the coalbed methane.
  • the essence of the catalytic deoxidation process is the catalytic combustion of C3 ⁇ 4 under a rich oxygen-poor atmosphere.
  • the main reaction of this process is C +20 2 ⁇ C0 2 + 2H 2 0. It can be seen from the above reaction that if about 10% of 0 2 is directly removed, about 5% of C is consumed, which may cause the catalyst bed temperature to reach 1000 ⁇ or more (the gas adiabatic temperature rise is about 700 ° C or so). Therefore, the use of a circulating reactor and a partial product gas circulation process is an inevitable choice, not only can the reaction bed temperature be controlled below 650 ⁇ to facilitate reactor material selection and catalyst life extension, and can effectively eliminate the reaction temperature is too high.
  • Patent CN101139239A discloses a sulfur-tolerant catalytic deoxidation process rich in methane gas, which reduces the oxygen concentration to control the reaction temperature by circulating a partially deoxidized cooled gas.
  • the process employs a manganese-based sulfur-tolerant deoxidation catalyst, and in order to maintain the catalyst activity to achieve the desired deoxidation depth, a higher reaction temperature and a lower reaction space velocity are inevitable. Higher reaction temperatures increase the chance of side reactions and reduce the recovery of T C3 ⁇ 4.
  • C3 ⁇ 4 catalytic combustion means that under oxygen-rich conditions (fuel/air molar ratio can be as low as 1-5%), flameless combustion is carried out at a low light-off temperature (200-350 ° C) by means of a catalyst. , and the process of converting C oxidation to C0 2 and 3 ⁇ 40.
  • supported noble metal catalysts are widely used in C catalytic combustion processes (especially low temperature) due to their higher catalytic activity, lower light-off temperature and better anti-toxic properties. Burning section).
  • the activity of the noble metal Pd is higher than that of Pt and Rh, and A1 2 0 3 , Si0 2 , Ti0 2 and Zr0 2 are generally used as carriers.
  • a non-negligible phenomenon for noble metal ruthenium catalysts is that Pd is easily converted to PdO under oxygen-rich reaction conditions, and PdO can be reductively decomposed into metal Pd as the reaction temperature increases.
  • the decomposition conversion temperature of PdO ⁇ Pd is affected by the composition of the feed gas and the catalyst. It is different, usually around 700-800 °C.
  • This balance and mutual conversion between PdO and Pd may, in some cases, cause the catalytic combustion process to become unstable and have an Oscillatory behavior, that is, the catalytic activity increases as the reaction temperature increases. A sudden drop in activity.
  • a polymetallic active component catalyst such as Pd-Pt (K. Persson et al., J.
  • Pd-Pt-Ni Pd-Pt-Ni
  • a Pd-based catalyst is added with a rare earth element such as Ce, La, Nd, Sm (US Patent 5216875), and a special catalyst preparation method using a noble metal PdO/Pd multilayer distribution (JP63088041) to delay the transformation of PdO ⁇ Pd morphology.
  • the above means inhibit the attenuation of the catalyst activity to a certain extent.
  • the above-mentioned catalytically active oscillation phenomenon not only exists but also appears more frequently and intensely. This is because, on the one hand, as the reaction progresses and the concentration of 02 decreases, the decomposition conversion temperature of PdO ⁇ Pd gradually advances; on the other hand, under the overall reducing atmosphere, PdO will be at a lower temperature. The catalyst is rapidly and completely reduced to Pd, resulting in loss of catalyst activity. In severe cases, the reaction bed temperature may be lower than the catalyst light-off temperature, resulting in termination/extinction of the reaction, thereby enabling deoxygenation purification by catalytic combustion of coalbed methane. The solution is difficult to achieve.
  • the above various supported precious metal Pd catalysts developed under the conditions of oxygen enrichment may not be suitable for the C catalytic combustion condition under the rich and oxygen-poor conditions of the present invention, and it is necessary to specifically target the rich and oxygen-poor
  • a new type of supported precious metal Pd catalyst was developed in a reducing atmosphere.
  • the object of the present invention is to provide a coalbed methane deoxidation catalyst, a preparation method thereof and a process for catalytic deoxidation of oxygen-containing coalbed methane, and solve the safety caused by the presence of o 2 in the process of liquefaction, storage and transportation of coalbed methane, It can be applied to the catalytic deoxidation of oxygen-containing coalbed methane and the hidden danger of catalytic deoxidation of other oxygen-containing gases.
  • the invention provides a coalbed methane deoxidation catalyst comprising a main catalytically active component, a catalytic auxiliary agent and a catalyst carrier; wherein the main catalytically active component and the catalytic auxiliary agent are supported in the form of a coating on a structurally inert carrier The whole catalyst is prepared.
  • the coalbed methane deoxidation catalyst provided by the invention mainly comprises one or a combination of platinum group noble metals Pd, Pt, Ru, Rh, Ir, preferably Pd>Pd-Rh, Pd-Pt, Pd- One of Rh-Pt; the content of the main catalytically active component is 0.01-5% (preferably 0.1-1%) based on the total weight of the catalyst; in the precious metal or a combination thereof, the Pd content is based on the simple substance. , 50-100% (preferably 70-90%) of the total weight of the precious metal.
  • the coalbed methane deoxidation catalyst provided by the invention is an alkali metal/alkaline earth metal oxide and a Ce0 2 based composite oxide; the alkali metal/alkaline earth metal oxide content is 1-10% (preferably 2-5) of the total weight of the catalyst.
  • Ce0 2 -based composite oxide content is 1-70% (preferably 5-30%) of the total weight of the catalyst; in the Ce0 2 -based composite oxide, Ce0 2 content 30-100% (preferably 40-75%) of the total weight of the 00 2 -based composite oxide;
  • the alkali metal/alkaline earth metal oxide is one of Na 2 0, K 2 0, MgO, CaO, SrO, BaO or a combination thereof, preferably M g O, K 2 0, CaO;
  • a CeO 2 -based composite oxide is CeO 2 and a lanthanide rare earth element Pr, Nd, Sm, Eu, Gd or/and a transition element Y, Zr, La or/and
  • a binary or multicomponent complex of ⁇ - ⁇ 1 2 0 3 is preferably a Ce-Zr, Ce-Sm, Ce-Zr-AK Ce-Zr-Y composite oxide.
  • the coal bed gas deoxidation catalyst provided by the invention is characterized in that the catalyst carrier is one or more selected from the group consisting of cordierite honeycomb ceramics, mullite honeycomb ceramics, A1 2 3 honeycomb ceramics, metal honeycombs, metal foams and the like.
  • the present invention provides a method for preparing the above catalyst, the steps are as follows: (1) preparing a Ce0 2 -based composite oxide auxiliary agent, which is supported on a structurally inert catalyst carrier, dried and calcined to obtain a catalyst precursor A; (2) supporting an alkali metal/alkaline earth metal oxide to the catalyst precursor A obtained in the above step (1), and drying and calcining to obtain a catalyst precursor B; (3) a platinum group noble metal active component The catalyst precursor B obtained in the above step (2) is dried and calcined to prepare an oxidation catalyst C; (4) the oxidation catalyst C is reduced to obtain a final catalyst D.
  • the method for preparing a catalyst provided by the present invention wherein the Ce0 2 -based composite oxide auxiliary agent is an oxide of Ce0 2 and an lanthanide rare earth element or/and an oxide of a transition element or/and a ⁇ - ⁇ 1 2 3 3 A binary or multivariate crystallite mixture having a diameter of less than 500 nm.
  • the method for preparing a catalyst provided by the present invention is a method using a coprecipitation method, a homogeneous precipitation method, a reverse microemulsion method, a high temperature hydrothermal synthesis method, a rapid decomposition method, or the like (preferably Homogeneous precipitation method for the preparation of Ce0 2 and lanthanide rare earth elements
  • An oxide compound and / or transition elements and / or ⁇ - ⁇ 1 2 0 3 is formed entirely group Ce0 2 binary composite or polyhydric compound; as a powder form prepared Ce0 2 based composite oxide was dispersed in deionized water, using Wet high-energy ball milling to obtain a water-soluble slurry containing a Ce0 2 -based composite oxide auxiliary in a weight percentage of 20-40%, and adjusting the pH of the slurry to 3-4 with nitric acid, and then The slurry is coated on an inert catalyst carrier, dried and calcined to obtain a catalyst precursor A; the drying and calcine
  • Drying and roasting methods are preferably slow drying and slow calcination, such as drying in a vacuum oven at 60 ° C for more than 15 hours, rising to 500 ° C in a muffle furnace at a heating rate of 2.5 ° C / min. 2-4 hour.
  • the step (2) is an alkali metal/alkaline earth
  • the metal oxide is supported on the catalyst precursor A by impregnation with an aqueous solution of a precursor containing an alkali metal or alkaline earth metal oxide auxiliary component (for example, impregnating the catalyst precursor A with a Mg(N0 3 ) 2 solution to carry MgO), after drying and calcination, the catalyst precursor B is obtained; likewise, the drying and calcination of the catalyst precursor B is preferably selected from rapid drying and slow calcination, such as rapid drying in a microwave oven for 3-10 minutes in a muffle furnace. The temperature was raised to 700 ° C at a heating rate of 2.5 ⁇ / min for 2-4 hours; this step can be repeated until the required loading is obtained.
  • step (3) is that the catalytic active component of the platinum group noble metal is supported on the catalyst precursor B by impregnation with the aqueous solution of the precursor/mixed aqueous solution containing the precious metal component (for example, The mixed solution of PdCl 2 , RhCl 3 and PtCl 2 is impregnated with the catalyst precursor B to support Pd-Pt-Rh), and dried and calcined to prepare an oxidation catalyst C; likewise, the catalyst C is preferably dried and calcined.
  • the preparation method of the catalyst provided by the invention, the reduction mode of the oxidation state catalyst C in the step (4) may be 3 ⁇ 4 reduction and hydrazine hydrate reduction, preferably under the atmosphere of 450-550 C in a 10% H 2 -90% N 2 atmosphere. Restore for 2-4 hours.
  • the catalyst provided by the present invention is applied to a methane catalytic combustion process for the purpose of deoxidizing and purifying coalbed methane.
  • the catalyst of the present invention should have the following characteristics: stable combustion under a rich oxygen-poor atmosphere, high activity, long life, low ignition light-off temperature, and low side reactions (C reforming, C3 ⁇ 4) Pyrolysis carbon deposition reaction) and so on. Among them, stable combustion under a rich and oxygen-poor atmosphere is the most critical point of the catalyst and catalytic deoxidation process of the present invention.
  • the methane catalytic combustion catalyst for the coalbed methane deoxidation purification process provided by the invention has a conversion rate of 0 2 or more in the laboratory for nearly 3000 hours of life test, and is conventionally used for methane catalysis under oxygen-rich conditions. Compared with the Pd/Al 2 0 3 catalyst system in the combustion process, the active oscillation phenomenon of the catalyst is eliminated, indicating that the methane catalytic combustion catalyst provided by the invention has the advantages of high activity, good combustion stability and long service life.
  • the interaction with the noble metal Pd is used to realize the self-regulation of the micro-oxidation-reduction atmosphere on the catalyst, and the decomposition and conversion temperature of PdO ⁇ Pd is improved.
  • the reduced Pd can be rapidly oxidized to PdO, so that the ratio of the catalytically active component PdO/Pd is kept stable, the catalytic activity oscillation phenomenon is alleviated, and the purpose of stabilizing the combustion process is achieved.
  • the catalyst is reduced before use, and the obtained catalyst not only improves the stability of the combustion process, but also can greatly improve the low-temperature ignition start performance.
  • the above-mentioned excellent properties of the catalyst of the present invention indicate that the catalyst is particularly suitable for use in the catalytic combustion of methane for the purpose of deoxidizing and purifying coalbed methane.
  • the catalyst provided by the present invention can further expand the catalytic combustion process of CO and low-carbon hydrocarbons in a rich oxygen-depleted reducing atmosphere.
  • the concentration of oxygen in the gas of the qualified coalbed methane product is less than 0.2% (preferably 0.1%);
  • the operating pressure (gauge pressure) of the deoxidation reactor is 0-10 MPa, the inlet temperature of the catalyst bed during steady state operation is 250-450 ° C, and the outlet temperature of the catalyst bed is 450-650 ° C,
  • the volumetric reaction space velocity is ⁇ , ⁇ - ⁇ , ⁇ - 1 ; the preferred condition is that the operating pressure (gauge pressure) of the deoxidation reactor is 0.01-0.03 MPa, and the inlet temperature of the catalyst bed during steady-state operation is 285-325 ⁇ , catalyst
  • the exit temperature of the bed is 550-650 ° C, and the volume reaction space velocity is ⁇ , ⁇ , ⁇ 1 .
  • the CBM product gas is subjected to at least two stages of heat exchange/cooling to reduce the temperature to 30-500 and remove the moisture contained therein;
  • the ratio of the volumetric flow rate of the CBM product gas returned to the initial oxygenated coalbed methane is 0:1 to 6:1.
  • the coalbed methane deoxidation catalyst provided by the invention is applied to a process for catalytic deoxidation of an oxygen-containing coalbed methane, the heat exchange/cooling device comprising at least one high-temperature gas-gas heat exchanger or waste heat boiler, and at least one low-temperature gas-liquid heat exchanger
  • the high temperature gas heat exchanger or waste heat boiler can cool the degassing reactor outlet gas temperature to 150-500 ° C; the low temperature gas liquid heat exchanger can The temperature of the high temperature gas heat exchanger or the waste heat boiler outlet gas is cooled to 30-50 °C.
  • the coalbed methane deoxidation catalyst provided by the invention is applied to the process of catalytic deoxidation of oxygen-containing coalbed methane, and the ratio of the volumetric flow rate of the recycled coalbed methane product gas to the initial oxygenated coalbed methane is 0:1 to 4:1.
  • Gas circulation can be done in a variety of ways.
  • the CBM product gas returned by the cycle is a CBM product gas after heat exchange/cooling dehydration, and the gas and the high temperature reaction gas are exchanged for preheating, and then mixed with the normal temperature feed gas to enter the reactor.
  • the circulating CBM product gas is a high temperature gas at the outlet of the deoxidation reactor, and the gas and the normal temperature feed gas are mixed into the reactor.
  • the coalbed methane deoxidation catalyst provided by the invention is applied to the process of catalytic deoxidation of oxygen-containing coalbed methane, and the low-temperature starting process has two ways, one way is to directly introduce the volume flow rate of the raw material gas of the coalbed methane into the initial coalbed methane raw material gas.
  • the coalbed methane deoxidation catalyst provided by the invention is applied to a process for catalytic deoxidation of an oxygen-containing coalbed methane, wherein the recycled coalbed methane product gas is a coalbed methane product gas after heat exchange/cooling dehydration, and the gas and high temperature reaction gas are exchanged.
  • the heat is preheated and then mixed with the normal temperature feed gas into the reactor; or the recycled coalbed methane product gas is a high temperature gas at the outlet of the deoxidation reactor, and the gas and the normal temperature feed gas are mixed into the reactor.
  • the process of the invention can realize the ignition start of the catalytic deoxidation reaction at a low temperature and can Stable and efficient deoxygenation at low pressure, high space velocity and temperature less than 650 °C, and finally remove the volume percentage of oxygen in the oxygen-containing coalbed methane to less than 0.2%.
  • High catalyst activity, reaction space velocity and low catalyst bed pressure drop increase the treatment capacity of oxygen-containing coalbed methane per unit volume of catalyst, thereby reducing the cost of catalytic deoxygenation; low reaction temperature avoids high reaction temperature of non-precious metal catalysts.
  • the occurrence of side reactions such as C3 ⁇ 4 cracking carbon and steam reforming increases the recovery rate of C in coalbed methane.
  • the process of the invention is particularly suitable for the catalytic deoxidation process of oxygen-containing coalbed methane with large processing volume, low pressure head and frequent and severe concentration change of 0 2 .
  • Figure 1 is a H 2 -TPR spectrum of the catalyst samples Example-1, Example-2 and Example-3 of the present invention (curve (1) is Example-1, curve (2) is Example-2, and curve (3) is Example -3 ; experimental conditions are 10 vol% H 2 /90 vol% Ar mixed atmosphere, heating rate 10 ° C / min);
  • Example-2 is a graph showing the catalyst bed temperature as a function of reaction time during the deoxidation reaction of the catalyst sample Example-1 of the present invention (the dry molar composition of the raw material gas is 50% C, 2.85% 0 2 , N 2 equilibrium; The water vapor (3 ⁇ 40) molar content is 9.1%; the raw material gas has a GHSV of OOOhr' 1 (dry space air velocity));
  • Figure 3 is a graph showing the catalyst bed temperature as a function of reaction time during the deoxidation reaction of the catalyst sample Example-2 of the present invention (the dry molar composition of the feed gas is 50% C, 2.85% 0 2 , N 2 equilibrium; The water vapor (3 ⁇ 40) molar content is 9.1%; the raw material gas GHSV is AOOOOhf 1 (dry space airspeed));
  • Figure 4 is a catalyst bed of the catalyst sample Example-3 of the present invention during the deoxygenation reaction
  • the layer temperature varies with the reaction time (the dry molar composition of the feed gas is 50% C, 2.85% 0 2 , N 2 equilibrium; the water vapor (H 2 0) molar content in the feed gas is 9.1%; the GHSV of the feed gas Is ⁇ OOOhr' 1 (dry basis airspeed));
  • Figure 5 is a graph showing the catalyst bed temperature as a function of reaction time during the deoxidation reaction of the catalyst sample Comparative-1 of the present invention (the dry molar composition of the raw material gas is 50% C, 2.85% 0 2 , N 2 equilibrium; The water vapor (H 2 0) molar content is 9.1%; the raw material gas has a GHSV of OOOhr' 1 (dry space air velocity));
  • Figure 6 is a graph showing the catalyst bed temperature as a function of reaction time during the deoxidation reaction of the catalyst sample Comparative-5 of the present invention (the dry molar composition of the raw material gas is 50% C, 2,85% 0 2 , N 2 equilibrium; The water vapor (H 2 0) molar content is 9.1%; the raw material gas has a GHSV of OOOhr' 1 (dry basis space velocity).
  • Figure 7 is a graph showing the catalyst bed temperature as a function of reaction time during the deoxidation reaction of the catalyst sample Compariscm-8 of the present invention (the dry molar composition of the raw material gas is 50% € 3 ⁇ 4, 2.85% 0 2 , N 2 equilibrium; The water vapor (H 2 0) molar content is 9.1%; the raw material gas GHSV is ⁇ OOOhf 1 (dry space airspeed));
  • Figure 8 is a graph showing the catalyst bed temperature as a function of reaction time during the deoxidation reaction of Example-2-l of the catalyst of the present invention over 3000 hours (the dry molar composition of the feed gas is 50% C, 2.85% 0 2 , N 2 Balance; the water vapor (H 2 0) molar content in the feed gas is 9.1%; the GHSV of the feed gas is OOOhr' 1 (dry space airspeed));
  • Figure 9 is a graph showing the catalyst bed temperature as a function of reaction time during the deoxygenation reaction of Example-2-2 of the catalyst sample of the present invention at a high concentration of 0 2 (the dry molar composition of the feed gas is 39.15% C3 ⁇ 4, 12.60% O 2 , N 2 balance; the GHSV of the raw material gas is OOOh 1 (dry Base airspeed).
  • the invention provides an oxygen-containing coal bed gas catalytic deoxidation process, comprising a system low temperature starting process, a process flow and a process operation parameter;
  • Figures 10 and 11 show the oxygen-containing coalbed methane cyclic catalytic deoxidation process of the present invention, including two circulation modes.
  • Figure 10 is a low-temperature cycle process of partial coalbed methane product gas, wherein: 1 is a reactor; 2 is a circulating booster; 3 is a waste heat boiler or a high temperature heat exchanger; 4 is a water-cooled heat exchanger; 5 is a water-distributing tank;
  • the returned coalbed methane product gas is a coalbed methane product gas after heat exchange/cooling dehydration, and the gas and the high temperature reaction gas are exchanged for preheating, and then mixed with the normal temperature raw material gas and sent to the reactor by the low temperature circulating fan;
  • the ambient temperature feed gas can also be heat exchanged with the helium temperature reaction gas in a high temperature heat exchanger for preheating, and then mixed with the recycled product gas to enter the reaction.
  • Figure 11 is a high-temperature cycle process for partial coalbed methane product gas, wherein: 1 is a reactor; 2 is a circulating booster; 3 is a waste heat boiler or a high temperature heat exchanger; 4 is a water-cooled heat exchanger; 5 is a water separator
  • the coalbed methane product gas returned by the cycle is a high temperature gas at the outlet of the deoxidation reactor, and the gas is mixed with the normal temperature feed gas and sent to the reactor by a high temperature circulation fan.
  • the catalysts present in the specific examples of the present specification (excluding the comparative examples) have a de-o 2 conversion of 96% or more under the reaction conditions described in the specification of the present invention.
  • the combustion stability thereof is It is expressed by the temperature change of the upper, middle and lower parts of the catalyst bed. All numbers appearing in the specification and claims of the present invention, such as the inlet and outlet temperature ranges of the respective unit devices, the pressure ranges, the volume percentages representing the composition of the gas components, etc., should not be construed as being absolutely accurate. It is within the margin of error allowed by those skilled in the art to be understood by those skilled in the art.
  • the reaction space velocity of the present invention is determined by dividing the volumetric flow rate of the reaction gas feedstock (dry basis) into the reaction system per hour divided by the volume of the catalyst. Expressed as GHSV, the unit is hr'
  • the catalyst ignition light-off temperature of the present invention means that under the reaction process conditions described in the specification of the present invention, when the catalyst bed reaches a certain temperature, the bed temperature suddenly rises sharply and finally The combustion of the catalyst is stabilized. The temperature is determined to be the ignition temperature of the catalyst.
  • Off 02 of the present invention as Yi in the conversion of a given mole percent of the feed gas 02 to be converted, i.e., feed gas and product gas 02 difference in the number of moles with respect to mole percent of the feed gas 02, unit for%.
  • the cycle ratio referred to in the present invention refers to the ratio of the volumetric flow rate of the coalbed methane product gas that is recycled back to the initial oxygenated coal bed gas, expressed as R.
  • this patent changes the composition of the catalytic component of the catalyst, and introduces a rare earth catalytic component with a certain oxygen storage and storage function into the catalyst system to replace A1 2 0 3 , and uses its interaction with the noble metal Pd to achieve microscopic on the catalyst. Self-regulation of redox atmosphere, mitigating catalytic activity oscillation Like, to achieve the purpose of a stable combustion process.
  • Ce0 2 and Ce-containing solid solution have been widely studied and applied in three-way catalysts for automobile exhaust gas purification.
  • Ce 3+ and Ce 4+ conversion of Ce can be stored under lean conditions 02, 02 is released under the rich conditions to facilitate automobile exhaust CO and HC oxidation.
  • the presence of Ce0 2 can also inhibit the sintering of the A1 2 3 3 carrier and increase the dispersion of the precious metal catalytic component.
  • the role of Ce0 2 is slightly different from that in the three-way catalyst for automobile exhaust purification, mainly by using the mutual conversion between Ce 3+ and Ce 4+ to increase the decomposition conversion temperature of PdO ⁇ Pd. The reduced Pd can be rapidly oxidized to PdO.
  • Ce0 2 is introduced in the other lanthanide metals and / or other transition metals and / or ⁇ - ⁇ 1 2 0 3, and (3 ⁇ 40 2 form a double-or multiple composite oxide, interaction between the metal, The stability of Ce0 2 , the oxygen exchange capacity of Ce0 2 can be increased, the specific surface area can be increased, and the ignition start performance of the catalyst can be improved.
  • Ce0 2 -based composite oxide catalytic materials Some physical properties of Ce0 2 -based composite oxide catalytic materials such as specific surface area, particle size and distribution, pore size distribution, formation of single-phase solid solution, etc. will directly affect the oxygen exchange capacity of the CeO 2 -based composite oxide, thereby affecting the catalyst. Activity and stability.
  • Composition of a preferred Ce0 2 -based composite oxide catalytic material provided by certain embodiments of the present invention and the preparation method can make it have better properties such as high specific surface area, high oxygen exchange capacity and thermal stability.
  • the catalyst needs to be previously reduced.
  • the pre-reduced catalyst will maintain a better PdO/Pd ratio during ignition start-up and can be ignited at normal temperature (25 ° C).
  • the catalyst bed in order to adapt to the large flow of coalbed methane and the source conditions of the low pressure head, the catalyst bed must also have a low resistance drop.
  • Catalyst structures having a regular geometry, such as honeycomb catalysts, are advantageous in achieving a lower catalyst bed resistance drop.
  • the essence of the CBM catalytic deoxidation process is the catalytic combustion of C under a rich oxygen-poor atmosphere. It is well known that the C molecule has a regular tetrahedral structure and is an organic substance that is difficult to activate. Therefore, how to achieve the ignition start of the catalytic deoxygenation reaction of the coalbed methane at a lower temperature is the primary problem to be solved in the technical solution of the present invention.
  • Supported precious metal catalysts have higher catalytic activity, lower light-off temperature and better anti-toxic properties than various metal oxide, perovskite and hexaaluminate methane combustion catalysts. It is widely used in the low temperature light-off stage of C catalytic combustion process.
  • the main reactions of the coalbed methane catalytic deoxidation process are as follows -
  • coalbed methane catalytic deoxidation process may also have the following side reactions within a certain temperature range (B) - (F):
  • the lower reaction temperature helps to inhibit the occurrence of C-cracking carbon deposition reaction and steam reforming reaction. , reduce the 3 ⁇ 4 and (0 content in the deoxygenated coalbed methane product gas, increase the yield of formazan and the safety of operation. Control the catalyst bed temperature at a relatively low level (such as 650 ° C) to reduce the occurrence of side reactions. It is another key point of the catalytic deoxidation process of the present invention.
  • the catalytic deoxidation process of the present invention will employ a supported noble metal catalyst to achieve the above object.
  • the catalyst bed in order to adapt to the high flow rate of oxygen-bearing coalbed methane and the gas source conditions of the low pressure head, the catalyst bed must also have a low resistance drop.
  • catalyst structures having a regular geometry, such as honeycomb catalysts have advantages in achieving a lower catalyst bed resistance drop, allowing the deoxygenation reaction to operate at higher volumetric reaction space velocities.
  • the oxygen-containing coalbed gas treatment amount per unit volume of the catalyst is increased, thereby reducing the deoxidation cost.
  • a first aspect of the present invention provides a catalytic deoxygenation cycle process for an oxygen-containing coal bed gas, see Figure 10 and Figure 11.
  • Figure 10 and Figure 11 are only simplified schematic views of the process flow of the present invention, only the most basic features of the process of the present invention are disclosed, with many details being omitted, such as automatic control systems, sensor components, valves, and the like.
  • Those skilled in the art will be able to design more detailed integrated process drawings based on the basic characteristics of the process flow disclosed in the drawings.
  • the oxygen-containing coalbed methane catalytic deoxidation cycle process provided by the present invention, in the steady state operation, the oxygen-containing coalbed methane feed gas and the coalbed methane product gas sent back by the pressurized circulation fan 2 are mixed into the deoxidation reactor 1, the coalbed methane.
  • the heat exchange/cooling device comprises at least one high temperature gas heat exchanger or waste heat boiler 3, and at least one low temperature gas liquid heat exchanger 4.
  • the high temperature gas heat exchanger or waste heat boiler 3 can cool the degassing reactor outlet gas temperature to 150-500 °C.
  • the low temperature gas-liquid heat exchanger 4 can cool the temperature of the outlet gas of the high-temperature gas heat exchanger or the waste heat boiler 3 to 30-50 °C.
  • the gas circulation can be carried out in two ways.
  • the recycled coalbed methane product gas is a heat exchange/cooling dehydrated coalbed methane product gas, and the gas and the high temperature reaction gas are exchanged for preheating.
  • the product gas is low temperature cycle; in other embodiments, the recycled coalbed methane product gas is a high temperature gas at the outlet of the deoxidation reactor, and the gas and the normal temperature feed gas are mixed into the reaction.
  • the product gas is a high temperature cycle; in other embodiments, the ambient temperature feed gas is mixed with the recycle product gas, exchanged with the high temperature reaction gas for preheating, and then passed to the reactor.
  • a second aspect of the invention provides a set of operating process parameters and conditions suitable for use in the above-described oxygen-containing coalbed methane catalytic deoxygenation cycle process.
  • the deoxidation reactor is a fixed bed adiabatic reactor equipped with a noble metal monolithic catalyst, wherein the noble metal monolithic catalyst refers to a platinum group noble metal Pd, Pt.
  • the noble metal monolithic catalyst refers to a platinum group noble metal Pd, Pt.
  • a preferred supported noble metal monolith catalyst example is a platinum group noble metal Pd as a main catalytic active component, a Ce0 2 -La 2 0 3 binary composite oxide as a catalytic auxiliary, and a cordierite honeycomb ceramic as a physical carrier.
  • the catalyst may be, but not limited to, the above preferred embodiment, any noble metal monolithic catalyst having high low temperature catalytic deoxidation activity and stability at a temperature of less than 650 Torr. It is applied in the deoxidation process of the present invention.
  • the volume percentage concentration of 02 in the oxygen-containing coalbed methane raw material may vary between 1-15%, and the change in the concentration of o 2 in the coalbed methane is more suitable. Big features.
  • the operating pressure (gauge pressure) of the deoxidation reactor is 0-10 MPa
  • the inlet temperature of the catalyst bed during steady state operation is 250-450 ° C
  • the outlet temperature of the catalyst bed is 450-650 ⁇
  • the volume reaction space velocity is ⁇ , ⁇ - ⁇ , ⁇ 1 .
  • the operating pressure (gauge pressure) of the deoxidation reactor is 0.01-0.03 MPa
  • the inlet temperature of the catalyst bed during steady-state operation is 285-325 ° C, the outlet temperature of the catalyst bed.
  • the volume reaction space velocity is SO OO-SO ⁇ OOhr ⁇
  • the ratio of the volumetric flow rate of the recycled coalbed methane product gas to the initial oxygenated coalbed methane is from 0:1 to 6:1. In a preferred embodiment of the oxygen-containing coalbed methane catalytic deoxidation process of the present invention, the ratio of the volumetric flow rate of the recycled coalbed methane product gas to the initial oxygenated coalbed methane is from 0:1 to 4:1. The cycle ratio should be minimized to reduce the energy consumption of the booster fan while meeting the conditions of use of the catalyst.
  • a third aspect of the invention provides a method of achieving low temperature start-up of the oxygen-containing coalbed methane catalytic deoxygenation process system of the present invention.
  • the invention utilizes the characteristics of low ignition light temperature of the hydrogen-oxygen catalytic combustion reaction, and introduces small-sized hydrogen gas (3 ⁇ 4) into the oxygen-containing coalbed gas raw material gas to make
  • the deoxidation catalyst reacts with o 2 in the coalbed methane, and the combustion exothermic preheating catalyst bed reaches the light-off temperature of c catalytic combustion, so that the entire deoxidation system is smoothly started, and the supply is stopped when the system is stably operated.
  • the catalytic deoxygenation reaction is initiated by introducing a small volume of 4-10% of the volumetric flow rate of the coalbed methane feed gas directly into the oxygen-containing coalbed methane feed gas.
  • 0 2 and 3 ⁇ 4 of the coalbed methane burn the exothermic preheated bed on the deoxidation catalyst to 250-450 ° C to reach the catalytic combustion light-off temperature of methane.
  • the catalytic deoxygenation reaction is initiated by introducing into the initial coalbed methane feed gas that has been preheated to 30-50 Torr into a small amount of 4-10% of the volumetric flow rate of the coalbed methane feed gas, in the coalbed methane.
  • Oxygen and hydrogen are burned on the deoxygenation catalyst to exotherm the preheated bed to 250-450 C to reach the light-off temperature of the catalytic combustion of formazan.
  • the heating rate of 2.5 ° C / min was raised to 500 C for 2 hours to obtain 34.314 g of a composite oxide powder having a composition of 50% CeO 2 -35% ZrO 2 -15% Al 2 O 3 .
  • the powder had a BET specific surface area of 154.7 m 2 /g, and its 3 ⁇ 4-TPR spectrum is shown in Fig. 1, and the reduction process was 638.5 Mmol/g.
  • the above powder was mixed with 15 ml of a HNO 3 solution having a pH of 1.2 and 30 ml of deionized water, and ball-milled by wet ball milling for 18 hours to obtain a water-soluble slurry containing the above Ce-Zr-Al composite oxide.
  • the obtained slurry was adjusted with an appropriate amount of deionized water and a HN0 3 solution having a pH of 1.2, so that the pH was controlled within a range of 3-4, and the weight percentage of the solid matter was about 34%, and about 100 ml was obtained for the honeycomb.
  • Carrier coated water soluble slurry was adjusted with an appropriate amount of deionized water and a HN0 3 solution having a pH of 1.2, so that the pH was controlled within a range of 3-4, and the weight percentage of the solid matter was about 34%, and about 100 ml was obtained for the honeycomb. Carrier coated water soluble slurry.
  • a cordierite honeycomb ceramic carrier having a weight of 0.3764 g was immersed in the above slurry containing Ce-Zr-Al composite oxide, and the slurry was appropriately agitated, and the honeycomb was taken out after immersion for 3 minutes, and the honeycomb ceramic passage was purged with compressed air. The excess slurry was then quickly dried in a microwave oven for 3 minutes, and then calcined at 700 ° C for 2 hours in a muffle furnace to obtain a catalyst having a Ce-Zr-Al composite oxide loading of 0.03 lg. body. This procedure was repeated twice to obtain a catalyst intermediate having a Ce-Zr-Al composite oxide supporting amount of 0.0565 g.
  • the obtained catalyst intermediate was immersed in 50 ml of a 2.7 M Mg(N0 3 ) 2 solution, and the catalyst intermediate was supported with MgO of O.OMg in the same manner as above. Then, using the same method as above on the catalyst intermediate supporting MgO The precious metal catalytic component PdO was supported, and the impregnation liquid used was 50 ml of a PdCl 2 solution containing 7 mg/ml of Pd.
  • the desired oxidation state precious metal honeycomb ceramic catalyst is obtained after microwave drying and calcination for 700 Torr for 2 hours.
  • the above catalyst was reduced with a 10% 3 ⁇ 4-90% N 2 mixed gas at 450 ° C for 2 hours to obtain a noble metal elemental catalyst, sample number is Example-1, and its specific composition was 0.18% Pd / 3.13% MgO / 12.62% Ce -Zr-Al-Ox/84.07% Cordierite o
  • the prepared precipitate was centrifuged, and the filter cake was thoroughly washed twice with 60 L of boiling water in the reaction vessel with stirring, each time after washing. Centrifugal filtration was carried out. After washing twice with deionized water, the obtained filter cake was thoroughly dispersed in 10 L of isopropanol solvent to remove residual water in the precipitate, and the isopropanol was centrifuged and cleaned. Drying in a vacuum oven at 60 ° C for 20 hours, rising to 500 ° C in a muffle furnace at a heating rate of 2.5 ° C / min, and calcining for 2 hours to obtain a composition of 700 g by weight of 58% Ce0 2 - 42% Zr0 2 Ce-Zr composite oxide powder.
  • 100 g of ⁇ - ⁇ 1 2 3 powder was obtained at 750 g of ⁇ 1 ( ⁇ 0 3 ) 3 ⁇ 93 ⁇ 40.
  • the above 700 g of Ce-Zr composite oxide powder and 100 g of ⁇ - ⁇ 1 2 3 powder were uniformly dispersed in 500 ml of HN0 3 solution having a pH of 1.2 and 600 ml of deionized water.
  • the ball was ground by wet ball milling for 18 hours to obtain a water-soluble slurry of the ternary crystallite mixture containing the above Ce-Zr oxide and ⁇ - ⁇ 1 2 3 3 .
  • the slurry was adjusted with an appropriate amount of deionized water and a HN0 3 solution having a pH of 1.2, so that the pH was controlled within a range of 3-4, and the weight percentage of the solid was about 34%, which was about 2.3 L.
  • the particle size of the crystallite mixture in the slurry was less than 500 nm as determined by a particle size analyzer.
  • a part of the above slurry sample was dried and calcined at 500 ° C, and then characterized by BET and 3 ⁇ 4-TPR.
  • the BET specific surface area of the crystallite mixture was 143.2 m 2 /g, and the H 2 -TPR spectrum was shown in Fig. 1.
  • the process 3 ⁇ 4 consumption is 647.6 ⁇ 11 ⁇ 2.
  • a cordierite honeycomb ceramic carrier having a weight of 790 g was immersed in the above slurry containing the Ce-Zr-Al microcrystal mixture, and the slurry was appropriately agitated, and after immersion for 3 minutes, the honeycomb was taken out, and the honeycomb ceramic passage was purged with compressed air.
  • the slurry was then quickly dried in a microwave oven for 15 minutes and then calcined in a muffle furnace for 700 hours for 2 hours to obtain a catalyst intermediate having a Ce-Zr-Al crystallite mixture loading of 63.2 g. This procedure was repeated twice to prepare a catalyst intermediate having a Ce-Zr-Al crystallite mixture loading of 118.5 g.
  • the obtained catalyst intermediate was immersed in 2 L of a 2.7 M Mg(N0 3 ) 2 solution, and 29.1 g of M g O was supported on the catalyst intermediate by the same method as above.
  • the noble metal catalytic component PdO was supported on the catalyst intermediate supporting MgO by the same method as above, and the impregnation liquid used was 2 L of a PdCl 2 solution containing 7 mg of Pd.
  • the desired oxidation state of the noble metal honeycomb ceramic catalyst P was obtained.
  • the above catalyst was reduced with a 10% H 2 -90% N 2 mixed gas at 450 Torr for 4 hours to obtain a noble metal elemental catalyst.
  • a sample having a weight of 0.4521 g was randomly cut from the catalyst prepared above, and the sample code was Example-2, and the specific composition thereof was 0.18% Pd / 3.09% MgO / 12.62% Ce-Zr-Al-Ox / 84.11% cordierite.
  • a Ce-Zr composite oxide powder having a composition by weight of 58% Ce0 2 -42% Zr0 2 was prepared by the method for preparing a Ce-Zr composite oxide described in Example 2 and sampled for BET and H 2 -TPR characterization.
  • the Ce-Zr composite oxide has a BET specific surface area of 120.4 m 2 /g, and its 3 ⁇ 4-TPR spectrum is shown in Fig. 1, and the reduction process consumption is 810.7 ⁇ 1/3.
  • Example 3 The CeO 2 -based composite oxide slurry prepared in Example 1 and Example 2 was replaced by the above-prepared Ce-Zr composite oxide powder in place of the Ce-Zr-Al ternary composite or crystallite mixture in Example 1 and Example 2.
  • the Ce-Zr composite oxide slurry was prepared, and the Ce-Zr composite oxide, MgO and noble metal Pd were coated by the same coating coating method to obtain a specific composition of 0.18% Pd / 3.12 ° /.
  • a catalyst sample of MgO/12.76% Ce-Zr-Ox/83.94% cordierite, sample code is Example-3.
  • Catalyst samples Examples-1, Example-2 and Example-3 prepared in Example 1, Example 2 and Example 3 were subjected to CBM deoxidation performance test evaluation on a fixed bed reactor.
  • the cycle process for desulfurization of coalbed methane according to the present invention sets the following reaction evaluation conditions: the molar percentage composition of the feed gas is 50% C, 2.85% 0 2 , N 2 equilibrium (dry basis composition); The water vapor (H 2 0) molar content is 9.1%; the raw material gas has a GHSV of OOOhr (dry basis space velocity). The raw material gas together with the water is preheated to 30 (TC is introduced into the catalyst bed for deoxidation catalytic combustion reaction.
  • the external electric furnace of the reactor is maintained at a temperature of 330 ° (C in the raw material gas and the product gas, N 2 , C0 2 , CO and H 2 are detected by gas chromatography thermal conductivity detector; in the feed gas and product 0 2 Online detection via the PROLINE® Process Mass Spectrometer.
  • the catalytic bed temperature collects one data every 3 seconds.
  • Four thermocouples were placed in a fixed bed reactor to detect the temperature of the catalyst upper, middle, lower and gas flow bodies (represented by T in , T mid , T ut and T g , respectively). Unless otherwise stated, the catalytic combustion deoxygenation performance tests of the catalysts in the specific examples of the catalysts of the present invention are carried out under the above experimental conditions.
  • the concentration of 0 2 in the product gas detected by the PROLINE® process mass spectrometer was always maintained within 0.1%, ie the conversion of 0 2 was 96. %the above.
  • the typical product gas composition analyzed by gas chromatography is as follows: 49.02% C, 1.61% C0 2 , 0.2% H 2 , 0.14% CO, N 2 equilibrium. 3 ⁇ 4 and CO are derived from side reactions such as C3 ⁇ 4 partial oxidation or C3 ⁇ 4 steam reforming associated with the process, but the amount of side reactions is extremely small.
  • the temperature profiles of the catalyst samples E X ample-l, Example-2 and Example-3 during the reaction are shown in Figures 2, 3 and 4.
  • Example 1 the catalyst samples prepared in Example 1, Example 2 and Example 3 on a fixed bed reactor were Example-1, Example-2 and Example-3.
  • the ignition start performance test was performed.
  • the cycle process of desulfurization purification of coalbed methane described in the present invention sets the following ignition process conditions:
  • the molar percentage composition of the feed gas is 50% C, 6% 0 2 , N 2 balance (dry Further, in order to ensure successful ignition at room temperature (25 ⁇ ), it is necessary to introduce H 2 which accounts for 6% of the total flow rate of the above-mentioned raw material gas ; the GHSV of all gases is SOOOltf 1 (dry space air velocity).
  • the ignition start performance tests of the catalysts in the specific examples of the catalysts of the present invention are carried out under the above experimental conditions.
  • the success of the catalyst ignition at room temperature is based on whether the temperature of the catalyst bed rises under ignition conditions and reaches a stable combustion state. It can be seen from the experiment that the catalyst samples Example-1 and Example-2 prepared in Example 1 and Example 2 can be smoothly ignited and started at room temperature under the condition that the gas GHSV is SOOOhr" 1 (dry space air velocity), and the embodiment 3 Preparation of the catalyst sample Example-3 is difficult to start at room temperature, it needs to be preheated to above 50 °C to start smoothly. This is more than the Ce0 2 based composite oxide additive in Example-1 and Example-2. The high specific surface area is thus associated with a better precious metal dispersion.
  • This example attempts to clarify the effect of the composition and content of different CeO 2 -based composite oxide promoters on the catalytic deoxidation performance and ignition performance of the catalyst of the present invention.
  • a series of catalysts were then prepared by the method of preparing a catalyst as described in Example 2.
  • the detailed composition of the catalysts is shown in Table 1 below. Catalyst samples with less Ce0 2 content, Comparative-1 and Comparison ⁇ , and catalyst samples without Ce ⁇ 3 ⁇ 4 promoter Comparison ⁇ and Comparison-4 were used as a comparison of the catalyst of the present invention.
  • Example-4 8°/.Pd/3.35%MgO/13.52./.Ce0 2 /82, 95°/. Cordierite 100% CeO 2
  • Example-6 51%Ce0 2 -49%La 2 0 3
  • Example-8 41%CeO 2 -29%ZrO 2 -30%Al 2 O3
  • Example-4 to Example-9 the bed temperature of the catalyst is higher during the deoxygenation reaction
  • Example-8 is 25
  • This example attempts to clarify the effect of the composition and content of different noble metal catalytically active components on the catalytic deoxidation performance and ignition performance of the catalyst of the present invention.
  • the preparation and composition of the catalyst samples are the same as those of Example 3 except that the noble metal Pd is replaced by precious metal units or multi-component noble metals of different compositions and contents (due to the difference in the content of precious metals and the preparation of different batches of samples, the composition of the catalyst components may be slightly Differences).
  • the detailed composition of the catalyst is shown in Table 3 below. a catalyst sample containing no precious metal active component Pd therein Comparison-5 to Comparison-7 was used as a comparative sample of the catalyst of the present invention.
  • the stability of the catalytic combustion process under oxygen conditions is essential.
  • Example- 11 is 25
  • Example- 12 is 25
  • This example attempts to clarify the effect of different catalyst reduction conditions on the catalytic deoxidation performance and ignition performance of the catalyst of the present invention.
  • the preparation and composition of the catalyst samples were the same as in Example 3 except for the reduction conditions (due to the sample preparation of the same batch of wood, the composition of the catalyst components may be slightly different).
  • the detailed composition of the catalyst is shown in Table 5 below.
  • the catalyst sample Comparative-8 in the oxidation state was used as a comparative sample of the catalyst of the present invention.
  • Table 5 Catalyst sample sample code composition with different reduction conditions, wt% reduction conditions
  • the catalytic deoxidation performance test results of the above catalyst samples show that the catalyst samples of Examples-14 to Example-16 within the composition range of the catalyst formulation of the present invention have a relatively stable bed temperature during the deoxidation reaction. It indicates that the combustion process is stable; while the comparative Comparison-8 in this example shows a large degree of temperature fluctuation during the reaction (Fig. 7).
  • the above phenomenon can be explained by the fact that the catalyst sample reduced by reduction or hydrazine hydrate has a higher decomposition conversion temperature between PdO and Pd, and the re-reduced Pd can be rapidly oxidized to PdO, thereby making the combustion process more stable.
  • This example gives the experimental results of the long-term stability of the catalyst of the present invention in the catalytic deoxygenation reaction.
  • the experiment was carried out in a laboratory fixed bed reactor using the parallel sample Example-2-1 of Catalyst Example-2 in Example 2.
  • Catalyst at room temperature (25 ° C) In the atmosphere where the molar percentage composition is 45% C, 6% 0 2 , 6% H 2 , N 2 equilibrium (dry basis composition), the GHSV of all gases is SOOOhr- 1 (dry space air velocity), and When the combustion is stable, it is switched to the composition of the reaction raw material gas.
  • the molar percentage composition of the feed gas is 50.5% CH4, 2.83%0 2 , N 2 equilibrium (dry basis composition); the water vapor (H 2 0) molar content in the feed gas is 9.1%; the GHSV of the feed gas is OOOhr' 1 (dry basis airspeed).
  • concentration of 0 2 in the product gas detected by the PROLINE® process mass spectrometer was always maintained within 0.1%, that is, the conversion rate of 0 2 was above %%.
  • the typical product gas composition analyzed by gas chromatography is as follows: 49.13% C, 1.56% C0 2 , 0.18% 3 ⁇ 4, 0.15% CO, N 2 equilibrium. See Figure 8.
  • the above-described excellent properties of the catalyst of the present invention indicate that the catalyst is particularly suitable for use in the catalytic combustion of methane for the purpose of desulfurization of coalbed methane.
  • the results of the catalytic deoxygenation reaction of the catalyst of the present invention at a high concentration of 0 2 are given in this example.
  • the experiment was carried out in a laboratory fixed bed reactor using the parallel sample Example-2-2 of Catalyst Example-2 in Example 2.
  • the catalyst has a molar composition of 45% C3 ⁇ 4, 6%0 2 , 6% H 2 , N 2 at room temperature (25 ° C) (dry basis composition), and the GHSV of all gases is SOOOhr' 1 (dry space velocity)
  • the ignition is started in the atmosphere, and is switched to the composition of the reaction raw material gas when the combustion is stable.
  • the molar percentage composition of the feed gas is 39.15% C, 12.60% O 2 , N 2 equilibrium (dry basis composition); the feed gas contains no water vapor (H 2 0); the feed gas GHSV is OOOh 1 (dry basis air velocity) .
  • the combustion process is stable, and the concentration of 0 2 in the product gas detected online by the PROLINE® process mass spectrometer is always maintained at 0.1%. Within, that is, the conversion rate of 0 2 is above 96%.
  • the above experimental results show that the catalyst of the present invention can also be applied to a rich oxygen-rich oxygen-reducing atmosphere at a higher concentration of 0 2 and further extended to a catalytic combustion process of CO and low-carbon hydrocarbons.
  • Example 10 - Example 17 shows different gas circulation modes of coalbed methane products in the process of the present invention, as well as 0 2 concentration of different coalbed methane feed gas, reaction bed inlet temperature, outlet temperature, inlet pressure, cycle ratio R, reaction space
  • the effects of operating parameters such as speed on the gas composition of the deoxygenated coalbed methane product, wherein Example 15 and Example 17 are comparative examples, are not covered by the present invention.
  • the catalysts used in the experiments of Example 10 - Example 17 were all honeycomb ceramic monolith catalysts having a composition by weight of 0.2% Pd / 5% Ce0 2 - 5% La 2 0 3 / 79.8% cordierite.
  • C, N 2 , C0 2 , CO and 3 ⁇ 4 in the feed gas and product gas are detected by a gas chromatography thermal conductivity detector; 0 2 in the feed gas and product is detected online by the PROLINE® Process Mass Spectrometer.
  • the total amount of the above-mentioned raw material gas is 6% into the coalbed methane raw material gas, and the GHSV of all the gases is SOOOhf 1 (dry space air velocity) and the reactor inlet temperature is 25 ° C, 3 ⁇ 4 and The 0 2 in the coalbed methane begins to react on the catalyst, and the combustion exothermic preheating catalyst bed reaches the light-off temperature of the C catalytic combustion, so that the entire deoxidation system is smoothly started, and the supply is stopped when the system is stably operated.
  • Table 6 The experimental data for stable operation under various conditions are listed in Table 6 below. Among them, the product gas circulation modes of Examples 10 to 15 are low temperature cycles, and the cycle modes of Example 16 and Example 17 are high temperature cycles.
  • Example 15 Comparative Example 15
  • Example 17 Comparative Example 17
  • the examples of the operation of the process parameters according to the present invention all obtain a better deoxidation effect, and the 0 2 content of the coalbed methane product gas is less than 1000 ppm, that is, the depolarization ratio of 0 2 is greater than 98.5%;
  • the 3 ⁇ 4 and CO content are lower, the C loss is smaller, and the theoretical recovery is calculated according to the complete conversion of c3 ⁇ 4 and o 2 , thus ensuring a higher
  • Example 10 11 13 13 14 15 16 17 Oxygenated coalbed methane feed gas:
  • Example 18 - Example 21 gives a comparison of the ignition start performance of a catalytic deoxygenation process system under different conditions.
  • the catalyst used in the experiment is composed of a weight percentage.
  • More than 280 can make the deoxygenation reaction start. It is necessary to add the preheater to preheat the reaction raw materials before entering the deoxidation reactor, which undoubtedly increases the complexity of the deoxidation process.

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Description

煤层气脱氧催化剂及其制备方法和应用 技术领域
本发明属于化学领域，具体提供了一种煤层气脱氧催化剂、其制 备方法及其应用于含氧煤层气催化脱氧的工艺。
背景技术
煤层气是一种吸附于煤层中的可燃气体， 主要成份为高纯度甲 烷。 由于煤层气不含常规天然气中的硫（H2S)等有害杂质， 也不含 苯、汞、铅等可致癌的有毒物， 所以煤层气是世界公认的优质清洁能 源。
通过地面钻采和井下抽放两种技术开采出的煤层气质量差异很 大， 其利用方式也有很大不同。 前者开采出的煤层气中 C 浓度在 90%以上， 适合输入天然气管道系统， 可用作发电燃料、 工业燃料、 车用燃料、 化工原料和居民生活燃料。 而后者回收的煤层气中 C 浓度平均为 30-50%, 压力也较低， 通常只能就地作为民用燃料， 绝 大部分被燃烧放空， 造成了极大的能源浪费。 同时， 我国天然气资源 供求矛盾日趋严重。 因此， 这种中、低浓度的煤层气必须要考虑加工 提纯、 加压以便长距离输送、 使用。
煤层气提纯技术是指将 N2或空气与 C¾分离， 使煤层气中甲烷 含量相应增加，从而可以提高煤层气的热值及降低输送成本。煤层气 提纯技术主要包括低温深冷分离、 变压吸附和膜分离等三种。 专利 CN1952569A和 CN1908559A公开了一种含空气煤层气的低温双级精 馏液化分离工艺，其液化和分离都在低温下进行，液化天然气的产品 纯度可以达到 99%以上。然而该技术在分离过程中，随着甲烷浓度的 提高，排放废气的氧含量也被浓缩提高，不可避免地有一个阶段正好 是属于甲烷的燃烧和爆炸的范围， 存在着很大的安全风险。
一种较为安全的分离提纯方案是先以催化燃烧的方法选择性脱 除煤层气中的氧气再进行提纯処理。这种方法可以将煤层气中含氧量 降到 0.5%以下， 并彻底消除了操作过程的安全隐患。 目前可采用的 煤层气脱氧方式主要包括催化脱氧 （ZL02113628.9、 CN101139239A 等） 、 焦炭燃烧法 (ZL02113627.0、 CN1919986A)等。 煤层气焦炭 燃烧法脱氧工艺虽然能够有效脱除含氧煤层气中的 o2， 但是该工艺 采用焦炭做燃料（如采用无烟煤代替焦炭则带来 S02排放等问题） ， 能耗较高；补焦和除尘工艺也相对比较复杂；较高的反应温度不仅对 反应器材质提出了更高的要求，同时可能导致 C¾高温裂解及重整等 副反应发生，使煤层气中 C¾回收率降低。这些都增加了焦炭燃烧法 脱氧工艺的成本。
催化脱氧工艺的本质是富燃贫氧气氛下 C¾的催化燃烧，该过程 发生的主要反应为 C +202→C02+2H20。 由上述反应可见， 若直接 脱除 10%左右的 02， 则需消耗 5%左右的 C ， 由此可能导致催化剂 床层温度达 1000Ό以上 （气体绝热温升约 700°C左右）。 因此， 采用 循环反应器和部分产品气循环工艺是必然的选择，不仅可以将反应床 层温度控制在 650Ό以下以利于反应器材质选择和催化剂寿命的延 长， 同时可以有效消除由于反应温度过高而引起的副反应发生（由反 应体系热力学分析可知， 650°C以下甲烷的水蒸汽重整反应和裂解积 碳反应发生的可能性较小）。 专利 CN101139239A公开了一种富含甲 烷气体的耐硫催化脱氧工艺，通过循环部分脱氧冷却后的气体来降低 氧浓度控制反应温度。但是，该工艺采用锰系耐硫脱氧催化剂， 为维 持催化剂活性以达到要求的脱氧深度，必然要采用较高的反应温度和 较低的反应空速。较高的反应温度会增加 C 发生副反应的机会，降 低 T C¾回收率；较低的反应空速会使反应器设备体积庞大，增加煤 层气脱氧成本。此外，非贵金属催化剂的使用会使得催化脱氧反应的 点火起动变得困难；颗粒型催化剂则增加了床层的阻力降，不利于入 口压力较低的煤层气脱氧工艺。上述问题均使得该催化脱氧工艺技术 方案的可行性降低。相应地， 寻求一种起燃温度较低、能在富燃贫氧 气氛下长时间稳定操作的 C 催化燃烧催化剂，是该催化脱氧工艺的 关键所在。
通常意乂上的 C¾催化燃烧是指在富氧的条件下 （燃料 /空气摩 尔比可低至 1-5%)， 借助催化剂在低起燃温度下 （200-350°C )进行 无焰燃烧， 并将 C 氧化转化为 C02和 ¾0的过程。 与各种金属氧 化物催化剂相比， 负载型贵金属催化剂由于其具有更高的催化活性、 更低的起燃温度以及更好的抗毒性能而被广泛地应用于 C 催化燃 烧过程（特别是低温燃烧段）。 对 C 的氧化而言， 贵金属 Pd的活 性高于 Pt和 Rh，且普遍使用 A1203、 Si02、 Ti02以及 Zr02等为载体。 针对贵金属 Ϊ 催化剂，一个不可忽略的现象是，在富氧反应条件下， Pd极易向 PdO转变， 随着反应温度的升高， PdO又可被还原分解为 金属 Pd。 PdO→Pd的分解转换温度因原料气及催化剂组成的影响而 有所不同， 一般在 700-800 °C左右。 这种 PdO和 Pd之间的平衡及相 互转换，某些情况下会使得催化燃烧过程变得很不稳定，有振荡现象 产生（Oscillatory behavior), 即催化活性随着反应温度升高而上升又 伴随着突然的活性下降。为缓解这种催化活性振荡现象， 人们通过诸 如采用多金属活性组分催化剂如 Pd-Pt (K. Persson et al.， J. Catal, 231 (2005) 139.)和 Pd-Pt-Ni (JP61033233) 在 Pd基催化剂中加入 Ce、 La、Nd、Sm等稀土元素（US Patent 5216875),以及采用贵金属 PdO/Pd 多层分布 （JP63088041 ) 的特殊催化剂制备方法等来延缓 PdO→Pd 形态的转变， 上述手段在一定程度上均抑制了催化剂活性的衰减。
针对本发明所涉及的富燃贫氧条件下的 C 催化燃烧工况,上述 催化活性振荡现象不仅依然存在，而且表现得更为频繁和剧烈。这是 因为， 一方面随着反应的进行和 02浓度的逐渐降低， PdO→Pd的分 解转换温度会逐渐前移； 另一方面， 在整体的还原性气氛下， PdO会 在较低的温度下就被迅速完全还原为 Pd, 造成催化剂活性位丧失， 严重时可能引起反应床层温度低于催化剂起燃温度， 导致反应终止 / 熄火，从而使通过煤层气催化燃烧的方式进行脱氧净化的技术方案难 以实现。
鉴于此， 上述针对富氧条件下开发的各种负载型贵金属 Pd催化 剂可能并不适用于本发明所涉及的富燃贫氧条件下的 C 催化燃烧 工况，有必要专门针对富燃贫氧的还原性气氛，开发新型的负载型贵 金属 Pd催化剂。在催化剂中引入具有一定储放氧功能的稀土催化组 分， 通过其与 Pd的相互作用来调节催化剂上的微观氧化还原气氛， 提高 PdO— Pd的分解转换温度并使完全还原后的 Pd能够快速氧化为 PdO, 从而稳定催化剂上 PdO/Pd的比例， 缓解催化活性振荡现象， 达到稳定燃烧过程的目的。针对和煤层气処理量大、压头低、氧气浓 度变化频繁剧烈等工艺特点，有必要开发出更为高效实用的煤层气催 化脱氧循环工艺， 迸一步提高技术可行性和降低脱氧成本。
发明内容
本发明的目的是提供了一种煤层气脱氧催化剂、其制备方法及其 用于含氧煤层气催化脱氧的工艺，解决了煤层气液化、储运过程中由 于 o2的存在而导致的安全， 可应用于含氧煤层气催化脱氧以及其它 含氧气体的催化脱氧过程隐患。
本发明提供了一种煤层气脱氧催化剂，该催化剂包括主要催化活 性组分、催化助剂以及催化剂载体；其中主要催化活性组分和催化助 剂以涂层的形式担载在结构规整的惰性载体上制成整体催化剂。
本发明提供的煤层气脱氧催化剂，主要催化活性组分为铂族贵金 属 Pd、 Pt、 Ru、 Rh、 Ir中的一种或几种的组合， 优选 Pd> Pd-Rh、 Pd-Pt、 Pd-Rh-Pt中的一种； 主要催化活性组分的含量以贵金属单质 计， 占催化剂总重量的 0.01-5% (优选 0.1-1%); 在贵金属或其组合 中， Pd的含量以单质计，占贵金属总重量的 50-100% (优选 70-90%)。
本发明提供的煤层气脱氧催化剂， 催化助剂为碱金属 /碱土金属 氧化物和 Ce02基复合氧化物；碱金属 /碱土金属氧化物含量占催化剂 总重量的 1-10% (优选 2-5%); Ce02基复合氧化物含量占催化剂总重 量的 1-70% (优选 5-30%); 在 Ce02基复合氧化物中， Ce02的含量 占 002基复合氧化物总重量的 30-100% (优选 40-75%); 碱金属 /碱 土金属氧化物为 Na20、 K20、 MgO、 CaO、 SrO、 BaO中的一种或其 组合， 优选 MgO、 K20、 CaO; Ce02基复合氧化物为 Ce02与镧系稀 土元素 Pr、 Nd、 Sm、 Eu、 Gd或 /和过渡元素 Y、 Zr、 La或 /和 γ-Α1203 的双元或多元复合物， 优选 Ce-Zr、 Ce-Sm、 Ce-Zr-AK Ce-Zr-Y复合 氧化物。
本发明提供的煤层气脱氧催化剂，催化剂载体选自堇青石蜂窝陶 瓷、 莫来石蜂窝陶瓷、 A1203蜂窝陶瓷、 金属蜂窝、 金属泡沬等整体 结构载体材料的一种或多种。
本发明提供了上述催化剂的制备方法，步骤如下：（ 1 )制备 Ce02 基复合氧化物助剂，将其担载到结构规整的惰性催化剂载体上，经干 燥和焙烧， 得到催化剂前体 A; (2)将碱金属 /碱土金属氧化物担载 到上述步骤（1 )得到的催化剂前体 A上， 并经干燥和焙烧， 得到催 化剂前体 B; (3 )将铂族贵金属活性组份担载到上述步骤（2)得到 的催化剂前体 B上， 经干燥和焙烧， 制成氧化态催化剂 C; (4)将氧 化态催化剂 C进行还原， 得最终催化剂 D。
本发明提供的催化剂的制备方法， 所述 Ce02基复合氧化物助剂 是 Ce02与镧系稀土元素的氧化物或 /和过渡元素的氧化物或 /和 γ-Α1203形成的粒径小于 500nm的双元或多元微晶混合物。
本发明提供的催化剂的制备方法， 所述步骤（1 ) 为采用共沉淀 法、均相沉淀法、反相微乳液法、高温水热合成法、速分解法等中的 任意一种方法（优选均相沉淀法）制备的 Ce02与镧系稀土元素的氧 化物或 /和过渡元素的氧化物或 /和 γ-Α1203形成完全复合的 Ce02基双 元或多元复合物； 将制备的粉末态 Ce02基复合氧化物分散在去离子 水中， 采用湿法高能球磨制得含有 Ce02基复合氧化物助剂重量百分 含量在 20-40%之间的水溶性浆料，用硝酸调整浆料的 pH值在 3-4之 间， 然后将此浆料涂覆到惰性催化剂载体上， 经过干燥和焙烧，得到 催化剂前体 A; 催化剂前体 A的干燥及焙烧方式最好选择快速干燥 及缓慢焙烧，如在微波炉中快速干燥 3-10分钟，在马弗炉中以 2.5Ό/ 分钟的升温速率升至 700Ό焙烧 2-4小时； 此步骤可重复进行直至获 得所需要的担载量； 此步骤可重复进行直至获得所需要的担载量。
以采用 Ce( 03)3-6¾0、 Zr(N03)4-3H20和 Α1(Ν03)3·9Η20作为前 驱体制备 Ce-Zr-Al复合氧化物为例， 包括以下步骤：
( A-1 ) 配制含有 Ce(N03)3'6H20、 Zr(N03)4'3H20 和 Α1(Ν03)3·9Η20, 以及尿素的混合水溶液；
(Α-2)加热上述步骤（A-1 )中的混合水溶液至尿素分解并进一 步加热至沸腾状态搅拌数小时， 经均相共沉淀制得 Ce02基复合氧化 物前体；
(A-3)陈化、 过滤、 洗涤、干燥及焙烧步骤（A-2)所得的复合 氧化物前体， 得到粉末态的完全复合的 Ce02基三元复合氧化物。 干 燥及焙烧方式最好选择缓慢干燥及缓慢焙烧， 如在 60°C真空干燥箱 中干燥 15小时以上，在马弗炉中以 2.5°C/分钟的升温速率升至 500°C 焙烧 2-4小时。
本发明提供的催化剂的制备方法， 所述步骤（2)为碱金属 /碱土 金属氧化物是通过含有碱金属或碱土金属氧化物助剂组分的前驱体 水溶液浸渍的方式担载在催化剂前体 A上（如以 Mg(N03)2溶液浸渍 催化剂前体 A以担载 MgO)， 经过干燥和焙烧， 得到催化剂前体 B; 同样地， 催化剂前体 B的干燥及焙烧方式最好选择快速干燥及缓慢 焙烧，如在微波炉中快速干燥 3-10分钟，在马弗炉中以 2.5Ό/分钟的 升温速率升至 700°C焙烧 2-4小时； 此步骤可重复进行直至获得所需 要的担载量。
本发明提供的催化剂的制备方法， 所述步骤（3)为铂族贵金属 催化活性组分是通过含有贵金属组分的前驱体水溶液 /混合水溶液浸 渍的方式担载在催化剂前体 B上（如以 PdCl2、 RhCl3、 PtCl2的混合 溶液浸渍催化剂前体 B以担载 Pd-Pt-Rh)， 经过干燥和焙烧， 制成氧 化态催化剂 C; 同样地，催化剂 C的干燥及焙烧方式最好选择快速干 燥及缓慢焙烧， 如在微波炉中快速干燥 3-10分钟， 在马弗炉中以 2.5 °C/分钟的升温速率升至 70(TC焙烧 2-4小时;此步骤可重复进行直 至获得所需要的担载量。
本发明提供的催化剂的制备方法， 所述步骤（4) 中氧化态催化 剂 C的还原方式可以为 ¾还原和水合肼还原， 优选 10%H2-90%N2 气氛下于 450-550 C下还原 2-4小时。
本发明提供的催化剂应用于以煤层气脱氧净化为目的的甲烷催 化燃烧过程。
本发明的催化剂应具有如下特性： 能在富燃贫氧气氛下稳定燃 烧、 活性高、 寿命长、 点火起燃温度低、 副反应少（C 重整、 C¾ 裂解积碳反应）等。其中， 能在富燃贫氧气氛下稳定燃烧是本发明所 涉及的催化剂及催化脱氧工艺的最关键所在。
本发明提供的用于煤层气脱氧净化工艺的甲烷催化燃烧催化剂 在实验室近 3000小时的寿命实验中， 02的转化率始终维持在 96%以 上，与常规的用于富氧条件下甲烷催化燃烧过程的 Pd/Al203催化剂体 系相比，催化剂的活性振荡现象 ¾本被消除，表明本发明所提供的甲 烷催化燃烧催化剂具有活性高、燃烧稳定性好、寿命长等优点。通过 向催化剂体系中引入具有一定储放氧功能的稀土催化组分，利用其与 贵金属 Pd的相互作用实现了催化剂上的微观氧化还原气氛的自我调 节，提高了 PdO→Pd的分解转换温度并使还原后的 Pd能够快速氧化 为 PdO， 从而使催化活性组分 PdO/Pd的比例保持稳定， 缓解了催化 活性振荡现象，达到了稳定燃烧过程的目的。将催化剂使用前进行还 原，所得催化剂不仅会提高燃烧过程稳定性， 同时其低温点火起动性 能可以获得较大幅度提升。本发明催化剂的上述优良性能表明该催化 剂特.别适合于在以煤层气脱氧净化为目的的甲烷催化燃烧过程中应 用。
本发明所提供的催化剂还可进一步拓展应用于富燃贫氧还原性 气氛下的 CO及低碳烃类的催化燃烧过程。
具体如下：
通过向含氧煤层气原料气中引入预热到 25-50Ό的小股氢气， 在 脱氧催化剂上与氧气反应，燃烧放热预热催化剂床层达到甲烷催化燃 烧的起燃温度；稳态操作时，初始含氧煤层气和循环返回的煤层气产 品气混合进入装有贵金属整体结构催化剂的固定床绝热脱氧反应器， 煤层气中的甲烷与氧气在催化剂作用下反应生成二氧化碳和水，产品 气经过换热 /冷却以降温并脱除其所含的水分， 得到合格的煤层气产 品气；部分产品气以一定循环比返回至脱氧反应器入口与初始含氧煤 层气混合以控制脱氧反应器入口的煤层气氧浓度； 其中：
( 1-1 )含氧煤层气中氧气的体积百分比浓度为 1%-15%；
( 1-2)合格的煤层气产品气中氧气体积百分比浓度小于 0.2% (优 选为 0.1%)；
( 1-3)脱氧反应器的操作压力 （表压） 为 0-10MPa， 稳态操作 时催化剂床层的入口温度为 250-450°C， 催化剂床层的出口温度为 450-650°C， 体积反应空速为 Ι,ΟΟΟ-δΟ,ΟΟΟΙιτ-1; 优选条件为脱氧反应 器的操作压力（表压）为 0.01-0.03MPa， 稳态操作时催化剂床层的入 口温度为 285-325Ό , 催化剂床层的出口温度为 550-650°C , 体积反 应空速为 Ο,ΟΟΟ^Ο,ΟΟΟΙιτ·1。
( 1-4)煤层气产品气经过至少两级换热 /冷却使其温度降至 30-500并脱除其所含的水分；
( 1-5)循环返回的煤层气产品气与初始含氧煤层气的体积流量 之比为 0 : 1至 6 : 1。
本发明提供的煤层气脱氧催化剂应用于含氧煤层气催化脱氧的 工艺， 所述换热 /冷却装置包括至少一个高温的气气换热器或废热锅 炉， 以及至少一个低温的气液换热器； 高温气气换热器或废热锅炉可 将脱氧反应器出口气体温度冷却至 150-500°C ; 低温气液换热器可将 高温气气换热器或废热锅炉出口气体温度冷却至 30-50°C。
本发明提供的煤层气脱氧催化剂应用于含氧煤层气催化脱氧的 工艺，所述循环返回的煤层气产品气与初始含氧煤层气的体积流量之 比为 0 : 1至 4 : 1。 气体循环可以采用多种方式。 如循环返回的煤层 气产品气是经过换热 /冷却脱水后的煤层气产品气， 该股气体和高温 反应气体换热以进行预热，然后和常温原料气混合进入反应器。又如， 循环返回的煤层气产品气是脱氧反应器出口的高温气体，该股气体和 常温原料气混合进入反应器。
本发明提供的煤层气脱氧催化剂应用于含氧煤层气催化脱氧的 工艺，所述低温起动过程有两种方式，一种方式是通过直接向初始煤 层气原料气中引入占煤层气原料气体积流量 4-10%的小股 ¾, 煤层 气中的氧气和氢气在脱氧催化剂上燃烧放热预热床层到 250-450 ， 达到甲烷的催化燃烧的起燃温度；另一种方式是通过向经过加热器预 热到 30-50°C的初始煤层气原料气中引入占煤层气原料气体积流量 4-10%的小股 ¾，煤层气中的氧气和氢气在脱氧催化剂上燃烧放热预 热床层到 250-450°C， 达到甲烷的催化燃烧的起燃温度。
本发明提供的煤层气脱氧催化剂应用于含氧煤层气催化脱氧的 工艺， 所述循环返回的煤层气产品气是经过换热 /冷却脱水后的煤层 气产品气，该股气体和高温反应气体换热以进行预热，然后和常温原 料气混合进入反应器；或者循环返回的煤层气产品气是脱氧反应器出 口的高温气体， 该股气体和常温原料气混合进入反应器。
本发明工艺能够在低温下实现催化脱氧反应的点火起动并能在 低压、高空速和小于 650°C的温度条件下进行稳定、高效的脱氧反应， 最终将含氧煤层气中氧气的体积百分比含量脱除到 0.2%以下。 高的 催化剂活性、反应空速以及低的催化剂床层压力降提高了单位体积催 化剂上含氧煤层气的处理量，从而降低了催化脱氧成本；低的反应温 度避免了非贵金属催化剂反应温度高导致的 C¾裂解结碳及水蒸汽 重整等副反应的发生，提高了煤层气中 C 的回收率。本发明工艺特 别适用于処理量大、 压头低、 02浓度变化频繁剧烈的含氧煤层气的 催化脱氧过程。
附图说明
图 1为本发明催化剂样品 Example-1、 Example-2和 Example-3 的 H2-TPR谱图 （曲线 （1 )为 Example-1 , 曲线 （2) 为 Example-2， 曲线（3 )为 Example-3; 实验条件为 10vol%H2/90vol%Ar混合气氛， 升温速率 10°C/min);
图 2为本发明催化剂样品 Example-1在脱氧反应过程中催化剂床 层温度随反应时间变化曲线 （原料气的干基摩尔组成为 50%C ， 2.85%02, N2平衡； 原料气中的水蒸汽（¾0)摩尔含量为 9.1%; 原 料气的 GHSV为 OOOhr'1 (干基空速））；
图 3为本发明催化剂样品 Example-2在脱氧反应过程中催化剂床 层温度随反应时间变化曲线 （原料气的干基摩尔组成为 50%C ， 2.85%02， N2平衡； 原料气中的水蒸汽（¾0)摩尔含量为 9.1%; 原 料气的 GHSV为 AOOOOhf1 (干基空速））；
图 4为本发明催化剂样品 Example-3在脱氧反应过程中催化剂床 层温度随反应时间变化曲线 （原料气的干基摩尔组成为 50%C ， 2.85%02, N2平衡； 原料气中的水蒸汽（H20)摩尔含量为 9.1%; 原 料气的 GHSV为 ^OOOhr'1 (干基空速））；
图 5为本发明催化剂样品 Comparison-1在脱氧反应过程中催化 剂床层温度随反应时间变化曲线 （原料气的干基摩尔组成为 50%C ， 2.85%02, N2平衡； 原料气中的水蒸汽（H20)摩尔含量为 9.1%； 原料气的 GHSV为 OOOhr'1 (干基空速））；
图 6为本发明催化剂样品 Comparison-5在脱氧反应过程中催化 剂床层温度随反应时间变化曲线 （原料气的干基摩尔组成为 50%C ， 2，85%02， N2平衡； 原料气中的水蒸汽 (H20)摩尔含量为 9.1%； 原料气的 GHSV为 OOOhr'1 (干基空速））。
图 7为本发明催化剂样品 Compariscm-8在脱氧反应过程中催化 剂床层温度随反应时间变化曲线 （原料气的干基摩尔组成为 50%€¾, 2.85%02, N2平衡； 原料气中的水蒸汽（H20)摩尔含量为 9.1%； 原料气的 GHSV为 ^OOOhf1 (干基空速））；
图 8为本发明催化剂样品 Example-2-l在 3000小时以上的脱氧 反应过程中催化剂床层温度随反应时间变化曲线（原料气的干基摩尔 组成为 50%C ， 2.85%02, N2平衡； 原料气中的水蒸汽（H20)摩 尔含量为 9.1%; 原料气的 GHSV为 OOOhr'1 (干基空速））；
图 9为本发明催化剂样品 Example-2-2在高 02浓度的脱氧反应 过程中催化剂床层温度随反应时间变化曲线（原料气的干基摩尔组成 为 39.15%C¾， 12.60%O2， N2平衡；原料气的 GHSV为 OOOh 1 (干 基空速)。 本发明提供了一种含氧煤层气催化脱氧工艺， 包括系统低 温起动过程、 工艺流程及工艺操作参数；
附图 10、 11所示为本发明含氧煤层气循环催化脱氧工艺， 包括 两种循环方式。
图 10为部分煤层气产品气低温循环工艺，其中： 1为反应器； 2 为循环增压风机； 3 为废热锅炉或高温换热器； 4为水冷换热器； 5 为分水罐； 循环返回的煤层气产品气是经过换热 /冷却脱水后的煤层 气产品气，该股气体和高温反应气体换热以进行预热，然后和常温原 料气混合由低温循环风机送入反应器；在本发明所述的含氧煤层气催 化脱氧工艺的某些实施方案中，常温原料气还可以在高温换热器中和 髙温反应气体换热以进行预热， 然后和循环产品气混合进入反应器； 图 11为部分煤层气产品气高温循环工艺，其中： 1为反应器； 2 为循环增压风机； 3 为废热锅炉或高温换热器； 4为水冷换热器； 5 为分水罐； 循环返回的煤层气产品气是脱氧反应器出口的高温气体， 该股气体和常温原料气混合后由高温循环风机送入反应器。
本发明最佳具体实施方式
以下实施例将对本发明予以进一步的说明，但并不因此而限制本 发明。
除非另外指出， 在本发明说明书的具体实施例（不含比较例）中 出现的催化剂， 在本发明说明书所描述的反应工艺条件下， 其脱 o2 转化率均在 96%以上。同时，为更好地描述在本发明说明书的具体实 施例及比较例中给出的催化剂的催化燃烧过程，其燃烧稳定性情况均 以催化剂床层上、 中、下三部分的温度变化表示。在本发明说明书和 权利要求书中出现的所有数字，例如各个单元设备的进、出口温度范 围，压力范围，表示气体组分构成的体积百分比等数値均不应该被理 解为绝对精确値，该数値是在本领域内的普通技术人员所理解的、公 知技术所允许的误差范围内。在本发明说明书和权利要求书中出现的 精确的数値应该被理解为构成本发明的部分实施例。尽管在本发明给 出的实例中努力做到保证数値的精确性，但由于各种测量技术的标准 偏差， 任何测量得到的数値都不可避免地存在一定误差。
本发明所述的反应空速定乂为反应气体原料（干基）每小时进入 反应系统的体积流量除以催化剂的体积。以 GHSV表示，单位为 hr' 本发明所述的催化剂点火起燃温度是指在本发明说明书所描述 的反应工艺条件下，催化剂床层达到某一温度时，床层温度突然急剧 上升并最终能够使催化剂的燃烧呈稳定状态。定乂该温度为催化剂点 火起燃温度。
本发明所述的脱 02转化率定乂为原料气中 02被转化的摩尔百分 比，即原料气与产品气中 02的摩尔数之差相对于原料气中 02的摩尔 百分比， 单位为％。
本发明所述的循环比是指循环返回的煤层气产品气与初始含氧 煤层气的体积流量之比， 以 R表示。
通常意乂上的甲垸催化燃烧是在富氧条件下进行的，负载型贵金 属催化剂（如 PdO/Al203等）是应用最广泛的低温燃烧段 C 燃烧催 化剂。 对 PdO/Al203催化剂而言， PdO和 Pd之间的分解转换不可避 免，从而引起催化剂活性振荡现象。这种振荡现象的本质是催化剂上 的催化活性组分 PdO/Pd比例的周期性改变。 目前， 越来越多的证据 表明， 在催化剂上维持一定的 PdO/Pd比例， 不仅可以获得更高的甲 烷燃烧活性， 同时也使得燃烧过程更加稳定 （C. A. Muller et aL, Catalysis Today, 47 (1999) 245.), 这可能与甲烷催化燃烧过程涉及的 氧化还原机理有关。
在富氧燃烧条件下， 由于过量 02的存在， 使得 PdO分解为 Pd 的温度大大后移， 同时分解后的 Pd又较易重新氧化成 PdO。 因此在 富氧燃烧条件下，催化剂的活性振荡现象在低温下表现得并不剧烈和 频繁。 Y. Deng等人的研究表明， 在富氧气氛下甲烷的催化燃烧过程 中， 催化剂的活性振荡现象只在 02/C 比为 1.0-2.0的较小范围内存 在 (Y. Deng et al.， Journal of Molecular Catalysis A: Chemical, 142 (1999) 51.) 。
而在本发明申请所涉及的富燃贫氧条件下，一方面随着反应的进 行和 02浓度的逐渐降低， PdO→Pd的分解转换温度会逐渐前移； 另 一方面，在整体的还原性气氛下， PdO会在较低的温度下就被迅速完 全还原为 Pd, 造成催化剂活性位丧失， 严重时可能引起反应床层温 度低于催化剂起燃温度， 导致反应终止 /熄火， 从而使通过煤层气催 化燃烧的方式进行脱氧净化的技术方案难以实现。为此，本专利通过 改变催化剂的催化组分体系构成，向催化剂体系中引入具有一定储放 氧功能的稀土催化组分替代 A1203， 利用其与贵金属 Pd的相互作用， 实现催化剂上微观氧化还原气氛的自我调节， 缓解催化活性振荡现 象， 达到稳定燃烧过程的目的。
Ce02和含 Ce的固溶体在汽车尾气净化三效催化剂中得到了广泛 的研究和应用。 Ce3+和 Ce4+的相互转换使 Ce能够在贫燃的条件下储 存 02，在富燃的条件下释放 02以促进汽车尾气中 CO和 HC的氧化。 同时， Ce02的存在还能够抑制 A1203载体的烧结和提高贵金属催化组 分的分散度。 在甲烷的催化燃烧反应中， Ce02的作用与在汽车尾气 净化三效催化剂中的作用略有不同，主要是利用 Ce3+和 Ce4+间的相互 转换来提高 PdO→Pd的分解转换温度并使还原后的 Pd能够快速氧化 为 PdO。 Ce02的上述作用已经在富氧条件下的甲垸催化燃烧过程中 得到证明 （P. O. Thevenin et al, Journal of Catalysis, 215 (2003) 78.)， 少量 Ce02 (如 5 wt%) 的应用即可起到上述作用。 但在富燃贫氧的 还原性条件下， PdO和 Pd之间的分解转换更加频繁， 同时转换温度 也将更低， 少量€602的引入将难以达到上述目的。 因此， 在本专利 申请中，将以 002全部或大部取代 A1203用以担载和分散贵金属 Pd。 更进一步地， 在 Ce02中引入其他镧系金属或 /和其他过渡金属或 /和 γ-Α1203, 与 (¾02形成双元或多元复合氧化物， 通过金属之间的相互 作用，可以提高 Ce02的稳定性、 Ce02的氧交换能力并增大比表面积、 改善催化剂的点火起动性能。
Ce02基复合氧化物催化材料的一些物理特性如比表面积、 粒径 大小与分布、 孔径分布、 是否形成单相固溶体等都将直接影响 Ce02 基复合氧化物的氧交换能力，进而影响催化剂的活性和稳定性。本发 明某些实施方案提供的较佳的 Ce02基复合氧化物催化材料组分构成 和制备方法可使其具有较佳的性能，如高比表面积、高的氧交换能力 和热稳定性等。
由热力学分析可知，在本发明专利所涉及的煤层气催化脱氧反应 工艺条件下， 甲烷的水蒸汽重整反应（CH4+H20→CO+3¾)和裂解 积碳反应（C →C+2H2) 发生的可能性较小。 尽管如此， 在本发明 某些实施方案中，仍然引入碱金属和碱土金属氧化物助剂，这将有利 于在反应过程中提高水的吸附强度，促进催化剂表面碳物种与水分子 之间的反应， 从而抑制催化剂表面积碳。
为使本发明催化剂发挥最佳效果并具有良好的点火起动性能，在 本发明某些较佳的实施方案中，需要将催化剂预先进行还原。经过预 还原的催化剂在点火起动过程中将保持较佳的 PdO/Pd比例， 并且能 够在常温下 （25°C )实现点火起动。
此外， 为适应煤层气大流量、低压头的气源条件， 催化剂床层还 必须具有较低的阻力降。具有规整几何形状的催化剂结构如蜂窝状催 化剂等在获得较低的催化剂床层阻力降方面具有优势。
煤层气催化脱氧工艺的本质是富燃贫氧气氛下 C 的催化燃烧。 众所周知， C 分子具有正四面体结构， 是一种较难活化的有机物。 因此，如何在较低的温度下实现煤层气催化脱氧反应的点火起动，是 本发明技术方案中所要解决的首要问题。与各种金属氧化物型、钙钛 矿型以及六铝酸盐型甲烷燃烧催化剂相比，负载型贵金属催化剂由于 其具有更高的催化活性、更低的起燃温度以及更好的抗毒性能而被广 泛地应用于 C 催化燃烧过程的低温起燃阶段。 煤层气催化脱氧过程发生的主要反应如下-
CH4(g)+202(g) = C02(g)+2H20(g) (A) 该反应为强放热反应， 放热量为 802.32kJ/mol。 含氧煤层气中的 02浓度有时可高达 15%, 由热力学计算可知， 若直接脱除 15%左右的 02，气体绝热温升约 1000Ό左右， 由此可能导致催化剂床层温度达到 1300°C以上。如此高的反应温度是绝大多数催化剂和反应器材质所无 法承受的。因此，如何移走反应过程中放出的大量的热， 是本发明技 术方案中所要解决的另一个关键问题。采用具有内构件的等温床反应 器会使得反应器设备造价急剧升高；如采用绝热床反应器，则循环反 应器和部分产品气循环工艺是必然的选择。
煤层气催化脱氧过程除发生上述主要反应 (A)外， 在一定的温度 范围内还可能发生如下副反应 (B)— (F):
CH4+ 0.5O2 = CO + 2H2 (甲垸部分氧化反应） （B) CO + 0.5O2 = CO2 (一氧化碳燃烧反应） （C)
H2 + 0.5O2 = H20 (氢气燃烧反应） （D)
CH4 = C + 2H2 (甲烷裂解反应） （E)
CH4+ H20 = CO + 3H2 (甲烷水蒸汽重整反应） （F) 根据上述各个反应的标准热力学数据，可以计算得出在 250-1450 的温度范围内， C¾完全燃烧反应 (A)占主导地位。 在温度低于 650 °C时， CO和 ¾的燃烧反应 (C)和 (D)也占一定的比例； C 裂解积碳反 应 (E)以及水蒸汽重整反应 (F)基本上不会发生反应。 当温度高于 650 °C时， 反应 (E)和 (F)可能发生， 而且本发明工艺的富 C 气氛更增加 了反应 (E)和 (F)发生的机会。 同时， 随着温度的增加， ¾和(：0的平衡 浓度增加， C¾的收率降低。 由此可知， 较低的反应温度有助于抑制 C 裂解积碳反应和水蒸汽重整反应的发生， 降低脱氧煤层气产品气 中的 ¾和( 0含量， 提高甲垸收率和操作的安全性。 控制催化剂床层 温度在相对较低的水平（如 650°C以内） 以减少副反应的发生， 是本 发明催化脱氧工艺的又一个关键所在。本发明催化脱氧工艺将采用负 载型贵金属催化剂来实现上述目的。
此外， 为适应含氧煤层气大流量、低压头的气源条件，催化剂床 层还必须具有较低的阻力降。与传统的颗粒状催化剂相比，具有规整 几何形状的催化剂结构如蜂窝状催化剂等在获得较低的催化剂床层 阻力降方面具有优势，使得脱氧反应可以在较高的体积反应空速下操 作，提高了单位体积催化剂的含氧煤层气处理量，从而降低了脱氧成 本。
基于以上考虑，本发明的第一个方面提供了一种含氧煤层气的催 化脱氧循环工艺流程， 见附图 10和附图 11。 附图 10和附图 11只是 本发明工艺流程的简单的示意图， 只公开了本发明工艺的最基本特 征，其中省略了许多细节，例如自动控制系统、传感器件、阀门等等。 熟悉本领域工作的技术人员完全可以根据附图披露的工艺流程基本 特点设计出更为详细的集成工艺图纸。
根据本发明提供的含氧煤层气催化脱氧循环工艺流程，稳态操作 时，含氧煤层气原料气和由增压循环风机 2送回的煤层气产品气混合 进入脱氧反应器 1, 煤层气中的 C 与 02在催化剂作用下反应生成 0)2和 ¾0， 产品气经过至少两级换热 /冷却使其温度降至 30-50Ό并 脱除其所含的水分， 得到 02体积百分比浓度小于 0.2%的合格煤层气 产品气；部分产品气以一定循环比返回至脱氧反应器 1入口与初始含 氧煤层气混合以控制脱氧反应器入口的煤层气氧浓度。在上述的含氧 煤层气催化脱氧工艺的实施方案中， 换热 /冷却装置包括至少一个高 温的气气换热器或废热锅炉 3， 以及至少一个低温的气液换热器 4。 高温气气换热器或废热锅炉 3 可将脱氧反应器出口气体温度冷却至 150-500°C。 低温气液换热器 4可将高温气气换热器或废热锅炉 3出 口气体的温度冷却至 30-50°C。 气体循环可以采用两种方式， 如在某 些实施方案中， 循环返回的煤层气产品气是经过换热 /冷却脱水后的 煤层气产品气，该股气体和高温反应气体换热以进行预热，然后和常 温原料气混合进入反应器，即产品气低温循环；在另一些实施方案中， 循环返回的煤层气产品气是脱氧反应器出口的高温气体，该股气体和 常温原料气混合进入反应器，即产品气高温循环；在另一些实施方案 中，常温原料气和和循环产品气混合，与高温反应气体换热以进行预 热， 然后进入反应器。
本发明的第二个方面提供了一套适用于上述含氧煤层气催化脱 氧循环工艺的操作工艺参数和条件。
在本发明所述的含氧煤层气催化脱氧循环工艺的实施方案中，脱 氧反应器为装有贵金属整体结构催化剂的固定床绝热反应器，其中贵 金属整体结构催化剂是指含有铂族贵金属 Pd、 Pt、 Ru、 Rh、 Ir中的 一种或几种催化活性组分、载体为堇青石蜂窝陶瓷、莫来石蜂窝陶瓷、 A1203 、金属 、金属 等结构规整惰性材料的催化剂。 例如一个较佳的负载型贵金属整体催化剂实例是以铂族贵金属 Pd为 主要催化活性组分， 以 Ce02-La203双元复合氧化物为催化助剂， 以 堇青石蜂窝陶瓷为物理载体。然而， 在本发明所述的脱氧工艺中， 催 化剂可以但不限于采用上面的较佳实例， 任何在小于 650Ό的温度条 件下具有较高的低温催化脱氧活性和稳定性的贵金属整体结构催化 剂均可在本发明所述的脱氧工艺中应用。
在本发明所述的含氧煤层气催化脱氧工艺的实施方案中，含氧煤 层气原料中 02的体积百分比浓度可在 1-15%之间变化， 适应煤层气 中 o2浓度变化幅度较大的特点。 脱氧反应器的操作压力 （表压）为 0-10MPa, 稳态操作时催化剂床层的入口温度为 250-450°C， 催化剂 床层的出口温度为 450-650Ό , 体积反应空速为 Ι,ΟΟΟ-δΟ,ΟΟΟΐΜ··1, 低 压和高空速适应煤层气原料气処理量大、压头低等特点，可进行高负 荷操作以降低 ΙΜ氧成本；且 650Ό以下操作能够有效消除 C 裂解及 水蒸汽重整等副反应，提高 C 回收率。在本发明专利的较佳实施方 案中， 脱氧反应器的操作压力（表压）为 0.01-0.03MPa， 稳态操作时 催化剂床层的入口温度为 285-325 °C， 催化剂床层的出口温度为 550-650°C， 体积反应空速为 SO OO-SO^OOhr^
在本发明所述的含氧煤层气催化脱氧工艺的实施方案中，循环返 回的煤层气产品气与初始含氧煤层气的体积流量之比为 0： 1至 6： 1。 在本发明所述的含氧煤层气催化脱氧工艺的较佳实施方案中，循环返 回的煤层气产品气与初始含氧煤层气的体积流量之比为 0： 1至 4： 1， 应在满足催化剂使用条件的前提下尽量减小循环比以降低增压风机 能耗。
本发明的第三个方面提供了一种实现本发明所述含氧煤层气催 化脱氧工艺系统低温起动的方法。
为实现整个含氧煤层气催化脱氧工艺系统的低温起动，本发明利 用氢氧催化燃烧反应起燃温度低的特点，通过向含氧煤层气原料气中 引入小股氢气（¾), 使 ¾在脱氧催化剂上与煤层气中的 o2反应， 燃烧放热预热催化剂床层达到 c 催化燃烧的起燃温度使整个脱氧 系统顺利起动， 系统稳定运行时停止供入 ¾。 在本发明所述的含氧 煤层气催化脱氧工艺的某些实施方案中，催化脱氧反应低温起动是通 过直接向含氧煤层气原料气中引入占煤层气原料气体积流量 4-10% 的小股 ¾， 煤层气中的 02和 ¾在脱氧催化剂上燃烧放热预热床层 到 250-450°C，达到甲烷的催化燃烧起燃温度。在另一些实施方案中， 催化脱氧反应低温起动是通过向已经预热到 30-50Ό的初始煤层气原 料气中引入占煤层气原料气体积流量 4-10%的小股 ¾， 煤层气中的 氧气和氢气在脱氧催化剂上燃烧放热预热床层到 250-450 C， 达到甲 垸的催化燃烧的起燃温度。
实施例 1
将 393.574g Zr(N03)4-3H20分散于 800ml去离子水中， 在搅拌的 情况下加热至 75-80°C使之完全溶解， 待溶液冷却后定容至 1000ml 得浓度为 1M的 Zr(N03)4溶液。 取上述溶液 100ml， 连同 43.465g Ce(N03)3，6¾0、 37.876g Α1(Ν03)3·9¾0 溶于去离子水中定容至 300ml, 在不断搅拌的情况下， 用分液漏斗向上述混合溶液中滴入 25-28%氨水， ίί水量根据 pH値控制， 直至 pH値达到 8-9。 然后将 所制得的沉淀充分搅拌 2小时， 抽滤， 并用 1200ml去离子水分 3次 洗涤滤饼， 将洗涤过的滤饼置于 60°C真空烘箱中干燥 20小时， 在马 弗炉中以 2.5°C/分钟的升温速率升至 500 C焙烧 2小时，得到 34.314g 重量百分比组成为 50%CeO2-35%ZrO2-15%Al2O3的复合氧化物粉体。 该粉体的 BET比表面积为 154.7 m2/g， 其 ¾-TPR谱图示于图 1， 还 原过程 ¾耗量为 638.5Mmol/g。
将上述粉体与 15ml pH値为 1.2的 HN03溶液和 30ml去离子水 混合， 采用湿式球磨法球磨 18小时， 制得含有上述 Ce-Zr-Al复合氧 化物的水溶性浆料。 用适量去离子水和 pH値为 1.2的 HN03溶液调 节所得的浆料， 使其 pH値控制在 3-4的范围内， 固体物重量百分含 量在 34%左右， 得到约 100ml适于蜂窝载体涂敷的水溶性浆料。
将重量为 0.3764g 的堇青石蜂窝陶瓷载体浸没于上述含有 Ce-Zr-Al复合氧化物的浆料中， 并适当搅动浆料， 浸没 3分钟后取 出蜂窝，用压缩空气吹扫蜂窝陶瓷通道内多余的浆料，然后用微波炉 快速干燥上述涂敷过的蜂窝载体 3分钟， 再于马弗炉中 700°C焙烧 2 小时得到 Ce-Zr-Al复合氧化物担载量为 0.03 lg的催化剂中间体。 重 复此过程 2次制得 Ce-Zr-Al复合氧化物担载量为 0.0565g的催化剂 中间体。 然后， 再将得到的催化剂中间体浸没于 50ml 2.7M 的 Mg(N03)2溶液中，采用上述同样方法使该催化剂中间体上担载 O.OMg 的 MgO。接着，再采用上述同样方法在担载 MgO的催化剂中间体上 担载贵金属催化组分 PdO， 使用的浸渍液为 50ml含有 7mg/ml Pd的 PdCl2溶液。经过微波干燥及 700Ό2小时焙烧后得到所需的氧化态贵 金属蜂窝陶瓷催化剂。上述催化剂用 10%¾-90%N2混合气体于 450°C 下还原 2小时， 得到贵金属单质态催化剂， 样品代号为 Example-1, 其具体组成为 0.18%Pd/3.13%MgO/12.62%Ce-Zr-Al-Ox/84.07%堇青 石 (Cordierite) o
实施例 2
在 70L的反应釜内加 30L去离子水， 搅拌状态下将 6459g的尿 素、 1311g的 (NH4)2Ce(N03)6投入釜内，最后将浓度为 1M的 Zr(N(¾)4 溶液 2.39L和另外 12L去离子水加入釜中。加热釜内料液至尿素分解 并进一步加热至沸腾状态（98-100°C )搅拌 4小时， 然后停止加热继 续搅拌陈化 2小时， 经均相共沉淀制得 Ce02-Zr02复合氧化物前体。 将制得的沉淀物离心过滤，用 60L的沸水在搅拌的状态下于反应釜内 充分洗涤滤饼两次，每次洗漆后均进行离心过滤。用去离子水洗涤过 滤两次后，再将所得滤饼充分分散于 10L的异丙醇溶剂中以带走沉淀 物中残余水， 将异丙醇离心过滤干净。 所得的沉淀物在 60°C真空干 燥箱中干燥 20小时，在马弗炉中以 2.5°C/分钟的升温速率升至 500°C 焙烧 2小时制得 700g重量百分比组成为 58%Ce02-42%Zr02的 Ce-Zr 复合氧化物粉体。
采用上述同样的方法， 以 750g的 Α1(Ν03)3·9¾0制得 100g的 γ-Α1203粉体。将上述 700g的 Ce-Zr复合氧化物粉体与 100g的 γ-Α1203 粉体均勻分散在 500ml pH値为 1.2的 HN03溶液和 600ml去离子水 中， 采用湿式球磨法球磨 18 小时， 制得含有上述 Ce-Zr氧化物及 γ-Α1203的三元微晶混合物水溶性浆料。 用适量去离子水和 ρΗ値为 1.2的 HN03溶液调节所得的浆料， 使其 ρΗ値控制在 3-4的范围内， 固体物重量百分含量在 34%左右， 得到约 2.3L适于蜂窝载体涂敷的 水溶性浆料。 经粒度分析仪测定， 该浆料中微晶混合物的粒度小于 500nm。 取部分上述浆料样品经干燥及 500°C焙烧后进行 BET及 ¾-TPR表征，该微晶混合物的 BET比表面积为 143.2 m2/g,其 H2-TPR 谱图示于图 1， 还原过程 ¾耗量为 647.6μιηο1½。
将重量为 790g的堇青石蜂窝陶瓷载体浸没于上述含有 Ce-Zr-Al 微晶混合物的浆料中， 并适当搅动浆料， 浸没 3分钟后取出蜂窝， 用 压缩空气吹扫蜂窝陶瓷通道内多余的浆料，然后用微波炉快速干燥上 述涂敷过的蜂窝载体 15分钟， 再于马弗炉中 700Ό焙烧 2小时得到 Ce-Zr-Al微晶混合物担载量为 63.2g的催化剂中间体。 重复此过程 2 次制得 Ce-Zr-Al微晶混合物担载量为 118.5g的催化剂中间体。然后， 再将得到的催化剂中间体浸没于 2L的 2.7M的 Mg(N03)2溶液中，釆 用上述同样方法使该催化剂中间体上担载 29.1g的 MgO。接着，再采 用上述同样方法在担载 MgO的催化剂中间体上担载贵金属催化组分 PdO,使用的浸渍液为 2L含有 7mg ml Pd的 PdCl2溶液。经过微波干 燥及 700°C2小时焙烧后得到所需的氧化态贵金属蜂窝陶瓷催化剂 P 上述催化剂用 10%H2-90%N2混合气体于 450Ό下还原 4小时， 得到 贵金属单质态催化剂。 从上述制得的催化剂上随机切割下重量为 0.4521g 的样品， 样品代号为 Example-2 , 其具体组成为 0.18%Pd/3.09%MgO/12.62%Ce-Zr-Al-Ox/84.11%堇青石。
实施例 3
采用实例 2中所述的制备 Ce-Zr复合氧化物的方法制备重量百分 组成为 58%Ce02-42%Zr02的 Ce-Zr复合氧化物粉体并取样进行 BET 及 H2-TPR表征。该 Ce-Zr复合氧化物的 BET比表面积为 120.4 m2/g, 其 ¾-TPR谱图示于图 1， 还原过程 ¾耗量为 810.7μιηο1/¾。
以上述制备的 Ce-Zr复合氧化物粉体代替实例 1及实例 2中的 Ce-Zr-Al三元复合物或微晶混合物，采用实例 1及实例 2中制备 Ce02 基复合氧化物浆料的方法制备 Ce-Zr复合氧化物浆料，并以相同的涂 层涂覆方法进行 Ce-Zr复合氧化物、 MgO及贵金属 Pd的涂覆， 制得 具体组成为 0.18%Pd/3.12°/。MgO/12.76%Ce-Zr-Ox/83.94%堇青石的催 化剂样品， 样品代号为 Example-3。
实施例 4
在固定床反应器上对实施例 1、实施例 2及实施例 3所制备的催 化剂样品 Example-l、Example-2和 Example-3进行了煤层气脱氧性能 测试评价。根据本发明所描述的煤层气脱氧净化的循环工艺设定了如 下的反应评价条件： 原料气的摩尔百分比组成为 50%C , 2.85%02, N2平衡 （干基组成）； 原料气中的水蒸汽 （H20)摩尔含量为 9.1%; 原料气的 GHSV为 OOOhr (干基空速）。 原料气连同水被预热到 30(TC通入催化剂床层进行脱氧催化燃烧反应。 为减少反应热损失， 反应器外用电加热炉维持温度 330° ( 。原料气和产品气中的 C 、 N2、 C02、 CO以及 H2通过气相色谱热导捡测器检测； 原料气和产物中的 02通过 PROLINE®过程质谱仪在线检测。 催化床层温度每隔 3秒采 集一个数据。在固定床反应器中布置四根热电偶，分别检测催化剂上、 中、 下以及气流主体的温度（分别以 Tin、 Tmid、 T。ut及 Tg表示）。 除 非另有说明，以下本发明催化剂的具体实施例中催化剂的催化燃烧脱 氧性能测试均在上述实验条件下进行。
在样品 Example-l、Example-2及 Example-3的催化脱氧反应过程 中， 通过 PROLINE®过程质谱仪在线检测的产品气中 02浓度始终维 持在 0.1%以内， 即 02的转化率在 96%以上。 由气相色谱分析的典型 产品气体组成如下： 49.02%C ， 1.61%C02， 0.2%H2， 0.14%CO, N2平衡。 ¾和 CO来自过程伴生的 C¾部分氧化或 C¾的水蒸汽重 整等副反应， 但副反应量极小。 反应过程中催化剂样品 EXample-l、 Example-2及 Example-3的温度变化曲线示于图 2、图 3及图 4。 由图 可见， 在数百小时的反应时间内， C 的催化燃烧过程均较为平稳， 没有观察到催化活性振荡现象。对比图 2至图 4可以发现，催化剂样 品 Example-1、 Example-2的催化剂床层上点温度有随着反应时间的 延长缓慢下降的趋势， 而样品 Example-3的温度则比较平稳， 可见样 品 Example-3的催化脱氧反应稳定性更好， 这与其中的 Ce02基复合 氧化物助剂具有更高的氧化还原能力有关（见图 1的 HrTPR图谱）。
为考察本发明催化剂在煤层气脱氧条件下的点火起动性能，在固 定床反应器上对实施例 1、实施例 2及实施例 3所制备的催化剂样品 Example-1、 Example-2及 Example-3进行了点火起动性能测试。
为最大限度地适应初始煤层气原料气中的 o2浓度之变化（变化 范围为 6-12%)， 本发明所描述的煤层气脱氧净化的循环工艺设定了 如下的点火工艺条件： 原料气的摩尔百分比组成为 50%C ， 6%02, N2平衡（干基组成)； 更进一步地， 为确保室温下（25Ό )点火成功， 需引入占上述原料气总流量 6%的 H2; 全部气体的 GHSV为 SOOOltf1 (干基空速)。 除非另有说明， 以下本发明催化剂的具体实施例中催 化剂的点火起动性能测试均在上述实验条件下进行。
室温下催化剂点火成功与否以点火条件下催化剂床层温度是否 上升并达到稳定燃烧状态为准。 由本实验可知， 实施例 1及实施例 2 所制备的催化剂样品 Example-1和 Example-2均能够在气体 GHSV为 SOOOhr"1 (干基空速） 的条件下于室温顺利点火起动， 而实施例 3制 备的催化剂样品 Example-3则室温下点火起动困难，需要将之预热到 50°C以上才能顺利起动。 这与 Example-1和 Example-2中的 Ce02基 复合氧化物助剂具有更高的比表面积从而具有更好的贵金属^散度 有关。
实施例 5
本实施例试图阐明的是不同的 Ce02基复合氧化物助剂组成及含 量对本发明催化剂的催化脱氧性能及点火性能的影响。采用实施例 2 中制备 Ce02基复合氧化物粉体的均相沉淀方法制备不同组成的 Ce02 基单元或多元复合氧化物粉体， 以及 γ-Α1203粉体及 Zr02粉体；然后 采用实施例 2中所述的制备催化剂的方法制备一系列催化剂，催化剂 的详细组成列于下表 1。 其中 Ce02含量较少的催化剂样品 Comparison-1 和 Comparison^, 以及不含 Ce<¾助剂的催化剂样品 Comparison^和 Comparison-4作为本发明催化剂的对比样。
表 1 Ce02基复合氧化物助剂组成及含量不同的催化剂样品 样品代号 组成， wt% Ce02基复合氧化物组成， wt%
Example-4 ( 8°/。Pd/3.35%MgO/13.52。/。Ce02/82,95°/。堇青石 100%CeO2
0.18%Pd/2.77% gO/12.30%Ce-Zr-Ox/84.75%堇
Example-5 66%Ce02-34%Sm203
青石
0.18%Pd 3.32%MgO/16.55%Ce-Zr-Ox/79.95%堇
Example-6 51%Ce02-49%La203
青石
0.18%Pd/3.87%MgO/14.03%Ce-Zr-Ox/81.92%堇
Example-7 85%Ce02-15%Zr02
青石
0.18%Pd/3.47%MgO/35.18%Ce-Zr-Al-Ox/61.17%
Example-8 41%CeO2-29%ZrO2-30%Al2O3
0.18%Pd/3.96%MgO/56.95%Ce-Zr-Ox/38.91%堇
Example-9 58%Ce02-42%Zr02
青石
0.18%Pd/3.14%MgO/15.58%Ce-Zr-Al-Ox/81.10%
Comparison-1 29%CeO2-21%ZrO2-50%Al2O3
0.18%Pd/2.16%MgO/l 8.25%Ce-Zr-Ox/79.41%堇
Comparison-2 26%Ce02-74%Zr02
青石
0.18%Pd/3.42%MgO/14.25%Y-Al203/82.15%堇青
Comparison-3
石
Comparison-4 0.18%Pd/2.96%MgO/l 3.89% Zr02/82.97%堇青石 对上述催化剂样品分别进行了催化脱氧性能测试。 实验结果表 明， 在本发明权利要求所述催化剂配方组成范围内的催化剂样品
5 Example-4至 Example-9， 在脱氧反应过程中催化剂的床层温度均较
为平稳， 表明燃烧过程稳定； 而本实施例中的对比样 Comparison-1 至 Comparison^,则在反应过程中均出现了较大程度的温度波动。典 型的温度波动见样品 Comparison-1的反应过程温度变化曲线图 5。上 述事实表明， 在本发明专利的催化剂配方组成中， 适量 Ce02的引入
10 对富燃贫氧条件下的催化燃烧过程稳定性起到了决定性的作用。
同样地，对上述催化剂样品分别进行了催化脱氧点火起动性能测 试。 实验结果列于表 2。 由表 2可见， 在本发明专利的催化剂配方组 成中，适量 γ-Α1203的引入对催化剂的低温点火起动起到了好的作用。 表 2 <¾02基复合氧化物助剂组成及含量不同的催化剂样品的点火起 动性能
样品代号 室温（25 C) 点火起动性能 点火起动温'度， °c
Example-4 否 45
Example-5 否 50
Example-6 否 50
Example-7 否 50
Example-8 是 25
Example-9 否 40
Comparison- 1 是 25
Comparison-2 否 50
Comparison-3 是 25
Comparison-4 否 55
实施例 6
本实施例试图阐明的是不同的贵金属催化活性组分组成及含量 对本发明催化剂的催化脱氧性能及点火性能的影响。 除了贵金属 Pd 用不同组成及含量的贵金属单元或多元贵金属代替外，催化剂样品的 制备及组成均与实施例 3相同（由于贵金属含量的不同及不同批次的 样品制备可能造成催化剂组分配方组成略有差异）。 催化剂的详细组 成列于下表 3。 其中不含贵金属活性组分 Pd 的催化剂样品 Comparison-5至 Comparison-7作为本发明催化剂的对比样。
表 3贵金属催化活性组分组成及含量不同的催化剂样品 贵金属活性组分组 样品代号 组成， t%
成， wt%
Example- 10 0.36°/。Pd/3.79%MgO/13.94%Ce-Zr-Ox /81.91%堇青石 100%Pd
0.36%Pd-0.04%Rh/3.77%MgO/13.60%Ce-Zr-Ox/82.23%
Example- 11 90%Pd-10%Rh 堇青石
0.36%Pd-0.04%Rh-0.04%Pt 3.43%MgO/11.96%Ce-Zr-Ox
Example- 12 82%Pd-9%Rh-9%Pt
/84.17%堇青石
0.18%Pd-0.02%Rh-0.02%Pt/3.82%MgO/13.65%Ce-Zr-Ox
Example- 13 82%Pd-9%Rh-9%Pt
/82.31%堇青石
Comparison-5 0.18%Pt//3.46%MgO/12.99%Ce-Zr-Ox/83.37%堇青石 100%Pt
Comparison-6 0, 18%Rh/2.86%MgO/17.85%Ce-Zr-Ox/79.11 %堇青石 100%Rh
0.18%Pt-0.04%Rh/3.61%MgO/15.42%Ce-Zr-Ox//80.75%
Comparison-7 82%Pt-18%Rh
堇青石
对上述催化剂样品进行的催化脱氧性能测试结果表明，在本发明 权利要求所述催化剂配方组成范围内的催化剂样品 Example-10至
5 Example-13, 在脱氧反应过程中催化剂的床层温度均较为平稳， 表明
燃烧过程稳定； 而本实施例中的对比样 Comparison-5 至
Comparison-7, 则在反应过程中均出现了较大程度的温度波动， 其中 对比样 Comparison-5更是出现了反应温度随反应时间延长逐渐降低 的情况， 表明催化剂的燃烧活性逐渐衰减（见图 6)。 上述事实表明，0 在本发明专利的催化剂配方组成中， 适量的贵金属 Pd对维持富燃贫
氧条件下的催化燃烧过程稳定性是必不可少的。
同样地，对上述催化剂样品分别进行了催化脱氧点火起动性能测 试。 实验结果列于表 4。 由表 4可见， 在本发明专利的催化剂配方组 成中，贵金属含量的适当增加将对催化剂的低温点火起动起到好的作 用。
表 4贵金属催化活性组分组成及含量不同的催化剂样品的点火起动 性能
样品代号 室温（25°C ) 点火起动性能 点火起动温度， V
Example- 10 .是 25
Example- 11 是 25
Example- 12 是 25
Example- 13 否 35
Comparison-5 否 40
Comparison-6 否 55
Comparison-7 否 35
实施例 7
本实施例试图阐明的是不同的催化剂还原条件对本发明催化剂 的催化脱氧性能及点火性能的影响。除还原条件不同外， 催化剂样品 的制备及组成均与实施例 3相同（由于木同批次的样品制备可能造成 催化剂组分配方组成略有差异）。 催化剂的详细组成列于下表 5。 其 中氧化态的催化剂样品 Comparison-8作为本发明催化剂的对比样。 表 5不同还原条件的催化剂样品 样品代号 组成， wt% 还原条件
0.18%Pd/3.55%MgO/13.41%Ce-Zr-Ox/82.86% 10%H2-90%N2混合气体
Example-14
于 450°C下还原 4小时
0.18%Pd/3.28%MgO/13.12%Ce-Zr-Ox/83.42%室温（25Ό ) 条件下 3%
Example- 15
水合肼还原 24小时
0.18%Pd/2.89%MgO/12.76%Ce-Zr-Ox/84.17%窒温 （25°C ) 条件下 3%
Example-16
水合肼还原 2小时
0.18%Pd/3.54%MgO/12.76%Ce-Zr-Ox/83.52%
Comparison-8 不还原
堇青石
对上述催化剂样品进行的催化脱氧性能测试结果表明，在本发明 权利要求所述催化剂配方组成范围内的催化剂样品 Example-14至 Example-16,在脱氧反应过程中催化剂的床层温度均较为平稳，表明 燃烧过程稳定；而本实施例中的对比样 Comparison-8，则在反应过程 中则出现了较大程度的温度波动 （图 7)。 上述现象可以解释为， 经 过 还原或水合肼还原的催化剂样品， 其 PdO和 Pd之间的分解转 换温度更高， 并且重新还原后的 Pd能够快速氧化为 PdO, 从而使得 燃烧过程更加稳定。
同样地，对上述催化剂样品分别进行了催化脱氧点火起动性能测 试。实验结果表明，低温下不同还原状态催化剂的点火起动由易到难 顺序为 ¾还原 >肼还原 24小时 >肼还原 2小时 >氧化态。
实施例 8
本实施例给出的是本发明催化剂在催化脱氧反应中的长期稳定 性实验结果。实验采用具体实施例 2中催化剂 Example-2的平行样品 Example-2-l ,在实验室固定床反应器内进行。催化剂于室温下（25°C ) 在摩尔百分比组成为 45%C ， 6%02, 6%H2, N2平衡（干基组成）， 全部气体的 GHSV为 SOOOhr-1 (干基空速）的气氛下点火起动， 并于 燃烧稳定时切换为反应原料气组成。 原料气的摩尔百分比组成为 50.5%CH4, 2.83%02, N2平衡（干基组成）；原料气中的水蒸汽（H20) 摩尔含量为 9.1%;原料气的 GHSV为 OOOhr'1 (干基空速）。在 3000 小时以上的反应过程中，通过 PROLINE®过程质谱仪在线检测的产品 气中 02浓度始终维持在 0.1%以内， 即 02的转化率在％%以上。 由 气相色谱分析的典型产品气体组成如下： 49.13%C ， 1.56%C02, 0.18%¾， 0.15%CO， N2平衡。 见图 8。 本发明催化剂的上述优良性 能表明该催化剂特别适合于在以煤层气脱氧净化为目的的甲烷催化 燃烧过程中应用。
实施例 9
本实施例给出的是本发明催化剂在高 02浓度下 （即采用非循环 脱氧工艺）的催化脱氧反应实验结果。实验采用具体实施例 2中催化 剂 Example-2的平行样品 Example-2-2， 在实验室固定床反应器内进 行。 催化剂于室温下 （25°C )在摩尔百分比组成为 45%C¾， 6%02, 6%H2， N2平衡（干基组成）， 全部气体的 GHSV为 SOOOhr'1 (干基空 速）的气氛下点火起动， 并于燃烧稳定时切换为反应原料气组成。原 料气的摩尔百分比组成为 39.15%C ， 12.60%O2, N2平衡 （干基组 成）； 原料气中不含水蒸汽（H20); 原料气的 GHSV为 OOOh 1 (干 基空速)。 在近 240小时的反应过程中 （图 9)·， 燃烧过程稳定， 通过 PROLINE®过程质谱仪在线检测的产品气中 02浓度始终维持在 0.1% 以内， 即 02的转化率在 96%以上。 上述实验结果表明本发明催化剂 还可应用于更高 02浓度的富燃贫氧还原性气氛下并进一步拓展至 CO及低碳烃类的催化燃烧过程。
实施例 10—实施例 17
实施例 10—实施例 17给出了本发明工艺中不同的煤层气产品气 循环方式以及不同煤层气原料气 02浓度、 反应床层入口温度、 出口 温度、入口压力、循环比 R、反应空速等操作工艺参数对脱氧煤层气 产品气组成的影响，其中实施例 15和实施例 17为比较例，不在本发 明保护之列。
实施例 10—实施例 17中实验所用催化剂均为重量百分比组成为 0.2%Pd/l 5%Ce02-5%La203/79.8%堇青石的蜂窝陶瓷整体催化剂。 原 料气和产品气中的 C 、 N2、 C02、 CO以及 ¾通过气相色谱热导检 测器检测； 原料气和产物中的 02通过 PROLINE®过程质谱仪在线检 测。
实验时首先向煤层气原料气中引入占上述原料气总流量 6%的 ¾， 在全部气体的 GHSV为 SOOOhf1 (干基空速）和反应器入口温度 为 25°C的条件下， ¾和煤层气中的 02在催化剂上开始发生反应， 燃 烧放热预热催化剂床层达到 C 催化燃烧的起燃温度使整个脱氧系 统顺利起动， 系统稳定运行时停止供入 ¾。 各个不同条件下稳定运 行实验数据列于下表 6。 其中， 实施例 10—实施例 15的产品气循环 方式为低温循环， 实施例 16和实施例 17的循环方式为高温循环。
由表 6可见，除实施例 15 (比较例 15)和实施例 17 (比较例 17) 夕卜，按照本发明所述的工艺参数操作的实施例均获得了较好的脱氧效 果， 煤层气产品气中的 02含量小于 lOOOppm, 即脱 02转化率大于 98.5%； 同时产品气中的 ¾和 CO含量较低， C 量损失较小， 接近 按照 c¾和 o2完全转化计算得到的理论回收率， 从而保证了较高的
5 C¾回收率。对比比较例 15和比较例 17可见， 虽然煤层气产品气中 的 02含量也小于 lOOOppm, 但由于脱氧反应器的操作温度较高， 使
^ ^ o 6 3- 得 C¾裂解积碳和水蒸汽重整等副 o反应加剧，导致了煤层气产品气中 较高的 1¾和 CO含量，从而增加了后续 § 5煤层气液化工艺过程的难度。 由此可见，通过调节煤层气产品气的循环比从而将脱氧反应器的操作
10 温度控制在 650°C以内， 是本发明工艺的关键所在 o。 表 6实施例 10-实施例 17中各个不同条件下稳定运行 3实验数据
实施例 10 11 13 13 14 15 16 17 含氧煤层气原料气：
流量， Nm3/hr
Figure imgf000038_0001
压力 （表压）， MPa, 0.01 0.01 0.01 0.02 0.01 0.01 0.01 温度， 'C 640 530 640 600 750 575 703 干基组成， vol% CH4 33.58 40.57 33.56 33.59 28.98 40.33 30.69
vol% N2 48.00 50.00 48.00 48.00 48.00 50.00 48.00
Figure imgf000039_0001
压力 （表压）， MPa 0.03 0.03 0.03 0.03 0.03 0.03 温度， °C 444 481 485 480 570 689 循环比 R 2 3 3 1 6 干基组成， vol% CH4 33.58 40.57 33.56 33.59 40.33 30.69
vol%N2 48.00 50.00 48,00 48.00 50.00 48.00 vol% 02 0.05 0.09 09 0.07 0.09 0.05 vol% H2 0.14 0.08 10 0.08 0.21 1.27 vol% CO 0.11 0.14 12 0.14 0.16 1.52 vol% C02 6.12 3.12 13 6.12 3.21 6.47
d c ci «S3
^ ^ L L 5
Figure imgf000039_0002
实施例 18—实施例 21
实施例 18—实施例 21给出了不同条件下催化脱氧反应工艺系统 的点火起动性能对比。 实验所用催化剂为重量百分比组成为
0.2%Pd/15%CeO2-5%La2O3/79.8%堇青石的蜂窝陶瓷整体催化剂， 点
5 火气源的体积空速为 δ,ΟΟΟΙιτ^ 由表 7的实验数据可见， 对于含氧煤
层气原料气来（对比例 21 )说， 在本实验^:件下， 只有将其预热到
280 以上才能使脱氧反应开始进行， 需要在进入脱氧反应器之前外 加预热器预热反应原料，这无疑增加了脱氧工艺的复杂程度。通过向 煤层气原料气中引入一定量的 ¾, ¾和煤层气中的 ο2在催化剂上开
10 始发生反应， 燃烧放热预热催化剂床层达到 C 催化燃烧的起燃温
度， 可以使整个脱氧系统在较低的温度下顺利起动。 表 7不同条件下催化脱氧反应工艺系统的点火起动性能 实施例 点火起动气源条件 点火起动温度 /气源预热温度， °c
9 37%C + 47%N2 + 6%02 +10%¾ 25
10 50%CH4 + 42%N2 + 2%02 +6%H2 50
11 37%CH4 + 55%N2 + 6%02 +2%H2 78
12 40%CH4 + 48%N2 + 12%02 280

Claims

权 利 要 求 一种煤层气脱氧催化剂， 其特征在于： 该催化剂包括主要催 化活性组分、 催化助剂以及催化剂载体； 其中：
主要催化活性组分选自铂族贵金属中的一种或几种的组合；主要 催化活性组分的含量以贵金属单质计， 占催化剂总重量的 0.01-5%; 在贵金属或其组合中， Pd的含量以单质计， 占贵金属总重量的 50-100%；
催化助剂为碱金属 /碱土金属氧化物和 (^02基复合氧化物；碱金 属 /碱土金属氧化物含量占催化剂总重量的 1-10%; Ce02基复合氧化 物含量占催化剂总重量的 1-70%; 在 Ce02基复合氧化物中， Ce02的 含量占 Ce02基复合氧化物总重量的 30-100%;
催化剂载体选自堇青石蜂窝陶瓷、 莫来石蜂窝陶瓷、 A1203蜂窝 陶瓷、 金属蜂窝、 金属泡沬中整体结构载体材料的一种或多种； 上述全部催化活性组分及催化助剂以涂层的形式担载在上述催 化剂载体上制成整体催化剂。
2、 按照权利要求 1所述煤层气脱氧催化剂， 其特征在于： 所述主要催化活性组分为铂族贵金属 Pd、 Pt、 Ru、 Rh、 Ir中的一种 或几种的组合。
3、 按照权利要求 1所述煤层气脱氧催化剂， 其特征在于： 所述 碱金属 /碱土金属氧化物为 Na20、 K20、 MgO、 CaO、 SrO、 BaO中的 一种或其组合。
4、 按照权利要求 1所述煤层气脱氧催化剂， 其特征在于： 所述 Ce02基复合氧化物为 Ce02与镧系稀土元素 Pr、 Nd、 Sm、 Eu、 Gd 或 /和过渡元素 Y、 Zr、 La或 /和 γ-Α1203的双元或多元复合物。
5、 按照权利要求 2所述煤层气脱氧催化剂， 其特征在于： 所述 主要催化活性组分为 Pd、 Pd-Rh Pd-Pt. Pd-Rh-Pt中的一种。
6、 按照权利要求 3所述煤层气脱氧催化剂， 其特征在于： 所述 碱金属 /碱土金属氧化物助剂为 MgO、 K20、 CaO中的至少一种。
7、 按照权利要求 4所述煤层气脱氧催化剂， 其特征在于： 所述 Ce02基复合氧化物为 Ce-Zr、 Ce-Sm、 Ce-Zr-AU Ce-Zr-Y复合氧化物 中至少一种。
8、 按照权利要求 1所述煤层气脱氧催化剂， 其特征在于： 所述 主要催化活性组分的含量以贵金属单质计， 古催化剂总重量的 0.1-1%。
9、 按照权利要求 1所述煤层气脱氧催化剂， 其特征在于： 所述 在贵金属或其组合中， Pd的含量以单质计， 占贵金属总重量的
10、按照权利要求 1所述煤层气脱氧催化剂， 其特征在于： 所述 碱金属 /碱土金属氧化物含量占催化剂总重量的 2-5%。
11、按照权利要求 1所述煤层气脱氧催化剂， 其特征在于： 所述 Ce02基复合氧化物含量占催化剂总重量的 5-30%。
12、按照权利要求 1所述煤层气脱氧催化剂， 其特征在于： 所述 在 Ce02基复合氧化物中， Ce02的含量占 Ce(¾ 复合氧化物总重量 的 40-75%。
13、一种权利要求 1所述催化剂的制备方法， 其特征在于： 步骤 如下：
( 1 )制备 Ce02基复合氧化物助剂，将其担载到结构规整的惰性 催化剂载体上， 经干燥和焙烧， 得到催化剂前体 A;
(2)将碱金属 /碱土金属氧化物担载到上述步骤（1 )得到的催 化剂前体 A上， 并经干燥和焙烧， 得到催化剂前体 B;
(3)将铂族贵金属活性组份担载到上述步骤（2)得到的催化剂 前体 B上， 经干燥和焙烧， 制成氧化态催化剂 C;
(4)将氧化态催化剂 C进行还原， 得最终催化剂 D。
14、 按照权利要求 13所述催化剂的制备方法， 其特征在于： 所 述 Ce02基复合氧化物助剂是 Ce02与镧系稀土元素的氧化物或 /和过 渡元素的氧化物或 /和 γ-Α1203形成的粒径小于 500nm的双元或多元微 晶混合物。
15、 按照权利要求 13所述催化剂的制备方法， 其特征在于： 所 述 Ce02基复合氧化物助剂的制备是采用共沉淀法、 均相沉淀法、 反 相微乳液法、 高温水热合成法、 速分解法中的任意一种方法制备的 Ce02与镧系稀土元素的氧化物或 /和过渡元素的氧化物或 /和 γ-Α1203 形成完全复合的 Ce02基双元或多元复合物。
16、 按照权利要求 13所述催化剂的制备方法， 其特征在于： 所 述步骤 ( 1 )为将制备的粉末态 Ce02基复合氧化物分散在去离子水中， 采用湿法高能球磨制得含有 Ce02基复合氧化物助剂重量百分含量在 20-40%之间的水溶性浆料， 用硝酸调整浆料的 pH值在 3-4之间， 然 后将此桨料涂覆到惰性催化剂载体上，经过干燥和焙烧，得到催化剂 前体 A, 此步骤可重复进行直至获得所需要的担载量。
17、 按照权利要求 13所述催化剂的制备方法， 其特征在于： 所 述步骤（2)为碱金属 /碱土金属氧化物是通过含有碱金属或碱土金属 氧化物助剂组分的前驱体水溶液浸渍的方式担载在催化剂前体 A上， 经过干燥和焙烧， 得到催化剂前体 B; 此步骤可重复进行直至获得所 需要的担载量。
18、 按照权利要求 13所述催化剂的制备方法， 其特征在于： 所 述步骤（3)为铂族贵金属催化活性组分是通过含有贵金属组分的前 驱体水溶液 /混合水溶液浸渍的方式担载在催化剂前体 B上， 经过干 燥和焙烧，制成氧化态催化剂 C; 此步骤可重复进行直至获得所需要 的担载量。
19、 按照权利要求 13所述催化剂的制备方法， 其特征在于： 所 述步骤（4)中氧化态催化剂 C的还原方式为 10%¾-90%]^2气氛下于 450-550°C还原 2-4小时。
20、权利要求 1所述催化剂应用于以煤层气脱氧净化为目的的甲 烷催化燃烧过程。
21、一种权利要求 1所述催化剂应用于含氧煤层气催化脱氧的工 艺， 包括系统低温起动过程、 工艺流程及工艺操作参数；
具体如下：
通过向含氧煤层气原料气中引入预热到 25-50 的小股氢气， 在 脱氧催化剂上与氧气反应，燃烧放热预热催化剂床层达到甲烷催化燃 烧的起燃温度；稳态操作时，初始含氧煤层气和循环返回的煤层气产 品气混合进入装有贵金属整体结构催化剂的固定床绝热脱氧反应器， 煤层气中的甲烷与氧气在催化剂作用下反应生成二氧化碳和水，产品 气经过换热 /冷却以降温并脱除其所含的水分， 得到合格的煤层气产 品气；部分产品气以一定循环比返回至脱氧反应器入口与初始含氧煤 层气混合以控制脱氧反应器入口的煤层气氧浓度； 其特征在于：
(21-1 )含氧煤层气中氧气的体积百分比浓度为 1%-15%；
(21-2)合格的煤层气产品气中氧气体积百分比浓度小于 0.2%;
(21-3)脱氧反应器的操作压力为 0-10MPa， 稳态操作时催化剂 床层的入口温度为 250-450°C， 催化剂床层的出口温度为 450-650Ό， 体积反应空速为 Ι,ΟΟΟ-δΟ,ΟΟΟΙιτ-1;
(21-4) 煤层气产品气经过至少两级换热 /冷却使其温度降至 30-50Ό并脱除其所含的水分；
(21-5)循环返回的煤层气产品气与初始含氧煤层气的体积流量 之比为 0 : 1至 6 : 1。
22、 按照权利要求 21所述含氧煤层气催化脱氧工艺， 其特征在 于： 所述合格的煤层气产品气中氧气体积百分比浓度小于 0.1%。
23、 按照权利要求 21所述含氧煤层气催化脱氧工艺， 其特征在 于：所述脱氧反应器的操作压力为 0.01-0.03MPa，稳态操作时催化剂 床层的入口温度为 285-325Ό , 催化剂床层的出口温度为 550-650Ό， 体积反应空速为 30,000-50，000hf 1。
24、 按照权利要求 21所述含氧煤层气催化脱氧工艺， 其特征在 于： 所述换热 /冷却装置包括至少一个高温的气气换热器或废热锅炉， 以及至少一个低温的气液换热器。
25、 按照权利要求 21所述含氧煤层气催化脱氧工艺， 其特征在 于：所述循环返回的煤层气产品气与初始含氧煤层气的体积流量之比 为 0 : 1至 4 : 1。
26、 按照权利要求 21所述含氧煤层气催化脱氧工艺， 其特征在 于：所述低温起动过程是通过直接向初始煤层气原料气中引入小股氢 气， 煤层气中的氧气和氢气在脱氧催化剂上燃烧放热预热床层到 250-450Ό , 达到甲烷的催化燃烧的起燃温度。
27、 按照权利要求 21所述含氧煤层气催化脱氧工艺， 其特征在 于：所述低温起动过程是通过向经过加热器预热的初始煤层气原料气 中弓 I入小股氢气，煤层气中的氧气和氢气在脱氧催化剂上燃烧放热预 热床层到 250-450Ό , 达到甲烷的催化燃烧的起燃温度。
28、 按照权利要求 21所述含氧煤层气催化脱氧工艺， 其特征在 于： 所述循环返回的煤层气产品气是经过换热 /冷却脱水后的煤层气 产品气，该股气体和高温反应气体换热以进行预热，然后和常温原料 气混合进入反应器。
29、 按照权利要求 21所述含氧煤层气催化脱氧工艺， 其特征在 于:所述循环返回的煤层气产品气是脱氧反应器出口的高温气体，该 股气体和常温原料气混合进入反应器。
30、 按照权利要求 24所述含氧煤层气催化脱氧工艺， 其特征在 于：所述高温气气换热器或废热锅炉可将脱氧反应器出口气体温度冷 却至 150-500 °C。
31、 按照权利要求 24所述含氧煤层气催化脱氧工艺， 其特征在 于：所述低温气液换热器可将高温气气换热器或废热锅炉出口气体温 度冷却至 30-50Ό。
32、按照权利要求 26和 27所述含氧煤层气催化脱氧工艺，其特 征在于： 所述供入的氢气量为初始煤层气原料气体积流量的 4-10%。
33、 按照权利要求 27所述含氧煤层气催化脱氧工艺， 其特征在 于： 其特征在于： 所述初始煤层气原料气的预热温度为 30-50 C。
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CN110743545A (zh) * 2019-09-29 2020-02-04 浙江工业大学 一种环氧乙烷双活性位燃烧催化剂及其制备与应用
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