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GB2274284A - Catalytic process for producing synthesis gas - Google Patents

Catalytic process for producing synthesis gas Download PDF

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GB2274284A
GB2274284A GB9326099A GB9326099A GB2274284A GB 2274284 A GB2274284 A GB 2274284A GB 9326099 A GB9326099 A GB 9326099A GB 9326099 A GB9326099 A GB 9326099A GB 2274284 A GB2274284 A GB 2274284A
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catalytic
oxygen
methane
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Domenico Sanfilippo
Luca Basini
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SnamProgetti SpA
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    • C01B3/00—Hydrogen; Gaseous mixtures containing hydrogen; Separation of hydrogen from mixtures containing it; Purification of hydrogen
    • C01B3/02—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen
    • C01B3/32—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen by reaction of gaseous or liquid organic compounds with gasifying agents, e.g. water, carbon dioxide, air
    • C01B3/34—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen by reaction of gaseous or liquid organic compounds with gasifying agents, e.g. water, carbon dioxide, air by reaction of hydrocarbons with gasifying agents
    • C01B3/38—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen by reaction of gaseous or liquid organic compounds with gasifying agents, e.g. water, carbon dioxide, air by reaction of hydrocarbons with gasifying agents using catalysts
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    • C01B3/02—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen
    • C01B3/32—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen by reaction of gaseous or liquid organic compounds with gasifying agents, e.g. water, carbon dioxide, air
    • C01B3/34—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen by reaction of gaseous or liquid organic compounds with gasifying agents, e.g. water, carbon dioxide, air by reaction of hydrocarbons with gasifying agents
    • C01B3/38—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen by reaction of gaseous or liquid organic compounds with gasifying agents, e.g. water, carbon dioxide, air by reaction of hydrocarbons with gasifying agents using catalysts
    • C01B3/386—Catalytic partial combustion
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    • C01B3/02—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen
    • C01B3/32—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen by reaction of gaseous or liquid organic compounds with gasifying agents, e.g. water, carbon dioxide, air
    • C01B3/34—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen by reaction of gaseous or liquid organic compounds with gasifying agents, e.g. water, carbon dioxide, air by reaction of hydrocarbons with gasifying agents
    • C01B3/38—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen by reaction of gaseous or liquid organic compounds with gasifying agents, e.g. water, carbon dioxide, air by reaction of hydrocarbons with gasifying agents using catalysts
    • C01B3/40—Production of hydrogen or of gaseous mixtures containing a substantial proportion of hydrogen by reaction of gaseous or liquid organic compounds with gasifying agents, e.g. water, carbon dioxide, air by reaction of hydrocarbons with gasifying agents using catalysts characterised by the catalyst
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    • C01B2203/00—Integrated processes for the production of hydrogen or synthesis gas
    • C01B2203/02—Processes for making hydrogen or synthesis gas
    • C01B2203/0205—Processes for making hydrogen or synthesis gas containing a reforming step
    • C01B2203/0227—Processes for making hydrogen or synthesis gas containing a reforming step containing a catalytic reforming step
    • C01B2203/0238—Processes for making hydrogen or synthesis gas containing a reforming step containing a catalytic reforming step the reforming step being a carbon dioxide reforming step
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    • C01B2203/02—Processes for making hydrogen or synthesis gas
    • C01B2203/025—Processes for making hydrogen or synthesis gas containing a partial oxidation step
    • C01B2203/0261—Processes for making hydrogen or synthesis gas containing a partial oxidation step containing a catalytic partial oxidation step [CPO]
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    • C01B2203/00—Integrated processes for the production of hydrogen or synthesis gas
    • C01B2203/08—Methods of heating or cooling
    • C01B2203/0805—Methods of heating the process for making hydrogen or synthesis gas
    • C01B2203/0838—Methods of heating the process for making hydrogen or synthesis gas by heat exchange with exothermic reactions, other than by combustion of fuel
    • C01B2203/0844—Methods of heating the process for making hydrogen or synthesis gas by heat exchange with exothermic reactions, other than by combustion of fuel the non-combustive exothermic reaction being another reforming reaction as defined in groups C01B2203/02 - C01B2203/0294
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    • C01B2203/00—Integrated processes for the production of hydrogen or synthesis gas
    • C01B2203/08—Methods of heating or cooling
    • C01B2203/0872—Methods of cooling
    • C01B2203/0883—Methods of cooling by indirect heat exchange
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    • C01B2203/00—Integrated processes for the production of hydrogen or synthesis gas
    • C01B2203/10—Catalysts for performing the hydrogen forming reactions
    • C01B2203/1005—Arrangement or shape of catalyst
    • C01B2203/1011—Packed bed of catalytic structures, e.g. particles, packing elements
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    • C01B2203/10—Catalysts for performing the hydrogen forming reactions
    • C01B2203/1041—Composition of the catalyst
    • C01B2203/1047—Group VIII metal catalysts
    • C01B2203/1064—Platinum group metal catalysts
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    • C01B2203/1041—Composition of the catalyst
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    • C01B2203/00—Integrated processes for the production of hydrogen or synthesis gas
    • C01B2203/12—Feeding the process for making hydrogen or synthesis gas
    • C01B2203/1205—Composition of the feed
    • C01B2203/1211—Organic compounds or organic mixtures used in the process for making hydrogen or synthesis gas
    • C01B2203/1235—Hydrocarbons
    • C01B2203/1241—Natural gas or methane
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    • C01B2203/14—Details of the flowsheet
    • C01B2203/141—At least two reforming, decomposition or partial oxidation steps in parallel
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    • C01B2203/80—Aspect of integrated processes for the production of hydrogen or synthesis gas not covered by groups C01B2203/02 - C01B2203/1695
    • C01B2203/82—Several process steps of C01B2203/02 - C01B2203/08 integrated into a single apparatus
    • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
    • Y02P20/00—Technologies relating to chemical industry
    • Y02P20/141—Feedstock
    • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
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Abstract

Catalytic process for producing synthesis gas by starting from methane, oxygen and optionally carbon dioxide and water, in which a noble metal catalyst supported on a solid carrier is used, which catalyst is arranged as a cascade of a plurality of catalytic beds, and the process is carried out under adiabatic conditions: -- by feeding the gas reactant stream upstream of the first catalytic bed and removing heat by heat exchange between the catalytic beds arranged in cascade, or -- by introducing the gas reactant stream partially upstream from the first catalytic bed and partially, as a cold stream, between the catalytic beds arranged in cascade, with the proviso that methane is at least partially fed to the first catalytic bed and oxygen is subdivided between all of the catalytic beds.

Description

"CATALYTIC PROCESS FOR PRODUCING SYNTHESIS GAS" The present invention relates to the production of synthesis gas ("syngas") by starting from methane, oxygen and, possibily, carbon dioxide and water, which process is carried out over a plurality of catalytic beds arranged in cascade and feeding the feedstock to the process as a plurality of subdivided streams fed upstream from each catalytic bed.
The synthesis gas, also referred to as "syngas" is prevailingly constituted by a gas mixture of CO and Hz. Producing the syngas mixture is presently the key passage in the technology of production of fueLs for motor vehicles by means of Fischer-Tropsch synthesis, in the technology of production of methanol and higher alcohols, and in ammonia synthesis. The investment costs and energy consumptions for operating the production units for syngas are estimated to be approximately 60% of total costs of the above listed processes - Syngas is presently produced by means of steam reforming or auto thermal reforming or processes of partial, non-catalytic, oxidation of hydrocarbons.The reactions which constitute the base of these conversions are the following: CnH. + n/2 Oz
n CO + m/2 H2 E1 CnHa + n HaO
n CO + (m + n/2) H2 23 CnHi + n C02
2n CO + m/2 H2 C37 Cn H.
Cn + m/2 H2 43 2CO
C + COz 53 CO + H20
HZ + COz C63 In greater detail, the steam reforming processes catalytically convert hydrocarbons/steam mixtures (Hz0:C=2.5 - 3.5), yielding CO/H2 mixtures with an H2/CO ratio which typically is of round 3.The chemical reactions involved in the process are 23, C4-53 and 63.
The H20/C ratio in the reactant mixture is both determined by the temperature and pressure conditions under which the reactions are carried out, and by the need of inhibiting the coal formation reactions C4-57.
The commonly used catalysts in these processes are based on Ni supported on Al, Mg, Si oxides. These carriers display high characteristics of heat stability and mechanical strength. The reactions are carried out inside tubular reactors installed inside a combustion chamber. The pressures inside the tubes are typically comprised within the range of from 1 to 5 MPa, and the gas temperature at tube outlets typically is of round 850or (reference is made, for instance, to "Catalysis Science and Technology"; Vol. 5 (1984), chapter 1, J.R. Rostrup-Nielsen).
The non-catalytic partial oxidation processes are less widely used and are employed in order to convert mixtures of oxygen, hydrocarbons, steam and water into syngas with H2/CO ratios of typically round 2. The chemistry of the process can be represented by the equations 13, E43-63. The facilities installed heretofore by Texaco and Shell (see Hydrocarbon Processing; April 1990, page 99) use adiabatic reactors inside which the reactions are initiated at reactors inlet by means of a burner in which total hydrocarbon combustion reactions C73 take place. These reactions produce large heat, steam and COz amounts.
Heat causes reactions of cracking of unburnt hydrocarbons and favours the steam 23 and COz C37 reforming reactions.
The operating temperatures are typically comprised within the range of from 1250 to 15000C, and the pressure is allowed to range from 3 to 12 MPa.
The processes of autothermal reforming are carried out inside adiabatic reactors to which mixtures of hydrocarbons, oxygen and steam are fed. In a first reaction zone, the reactions are initiated of total combustion of hydrocarbons, represented by the equation: CnHrn + (n + m/2) O2
n C0z + m/2 HzO C73 In a second zone inside a catalytic bed, the steam C27 and COz 33 reforming reactions take place.
In the catalytic bed, nickel-based catalysts are used, the characteristics of which are analogous to those as described above for steam reforming processes. In the autothermal reforming, mixtures of Hz/CO having values ranging from those of steam reforming processes to those of noncatalytic partial oxidation, are obtained.
The temperature of the gas streams at reactor outlets is typically comprised within the range of from 950 to 1000OC, but the temperature of the zone in which the burner is installed is considerably higher. The pressure inside the reactors is comprised within the range of from 2 to 4 MPa.
One from the main drawbacks which limit the possibilities of technological innovation in the definition of new catalytic reactors and new processes routes for syngas production and use is determined by the coal formation reactions C43-C53. Coal formation is not tolerated in the catalytic processes for syngas production and is prevented from occurring by using reactants mixtures containing steam and/or oxygen.
According to the syngas production processes and the operating conditions, therefore, restraints exist as to the composition of the reactant mixture and, in particular, as to its steam and/or oxygen contents; such restraints are generally expressed in terms of HzO/C and Oz/C ratios.
Extending the threshold values of composition of the reactant mixture, would make it possible innovative solutions to be designed for syngas production processes, because one might state that the characteristics of the reactors and of the process schemes in syngas production facilities are the result of complex interactions between the chemical properties of the catalysts and mechanical constraints to the characteristics of the materials used in the reactors.
In Italian patent application No. 19,162 A/90, filed on January 26th, 1990, to the same Applicant's name, disclosed is a process for syngas production by starting from carbon dioxide and light hydrocarbons, in particular methane, over a supported catalyst based on a metal from platinum group. Furthermore, in Italian patent application No. 21,326 A/90, filed on August 29th, 1990, to same Applicant's name, disclosed is a process for syngas production by means of a first step, of non-catalytic combustion of hydrocarbons with oxygen, followed by a second step, of reforming, in which the oxidation products from the first step are brought into contact with a further amount of hydrocarbons, in the presence of a supported catalyst of a metal from platinum group.
The present Applicant found now, according to the present invention, that the use of noble metal catalysts considerably reduces the width of the regions inside which the coal formation reaction takes place and therefore makes it possible reaction mixtures with low Hz 0/C (e.g., lower than 0.5) and Oz/C ratios (e.g., lower than 0.5) to be used without that the coal formation reaction are initiated.
Such a finding makes it possible said catalysts to be used in a process for syngas production in a reaction system consisting of a plurality of adiabatic catalytic beds arranged in cascade, in which a differentiated feed of the reactant mixture is preferably provided, and in which the composition of said mixture at the inlet to said catalytic beds may even have values of H20/C and O2/C ratios, which are lower than 0.5 and 0.5, respectively. Furthermore, a catalytic process which displays such characteristics makes it possible syngas mixtures to be obtained without requiring that at its inlet a burner is installed, because the combustion reactions are catalytically initiated at low temperatures.
More particularly, the process for syngas production, carried out on a plurality of adiabatic catalytic beds in cascade, according to the present invention, enables the following advantageous effects to be accomplished: -- reduction of temperature gradients and also of the highest temperature values inside said catalytic beds, with consequent lower thermal stresses being applied to the materials; in that way, traditional building materials can be used, with consequent savings in investment costs; -- possibility of directly obtaining, at the outlet from the catalytic partial oxidation reactor, syngas with Hz/CO ratios comprised within the range of from 0.9 to 3, without that the adjustment of the value of such a ratio requires that a further reactor for water gas shift (WGS) reactions r63 is used; -- possibility of avoiding using a burner at reactor inlet, with consequent saving in reactor investment costs; -- improvement of heat efficiency of syngas production process, both as compared to the commercial processes of non-catalyzed partial oxidation processes, and as compared to autothermal reforming processes; such an improvement is made possible because the configuration of the reactor makes it possible the heat recovery rates to be optimized, by preventing the unnecessary, extremely high temperatures which occur inside the interior of the reactors (in particular at inlet regions) used in the exisiting processes; -- possibility of kinetically controlling the coal generation reactions and, therefore, of reducing the values of H2O/C (steam mols/carbon mols) and 0z IC (oxygen mols/carbon mols) ratios in the reactant mixture; -- possibility of optimizing the process conditions, with in each layer the conditions of maxima reaction speed being reached, with the catalyst amount being consequently decreased (decreasing the catalyst amount is a determinative factor when noble metal-based catalytic system are used).
In accordance therewith, the present invention relates to a catalytic process for preparing synthesis gas by starting from methane, oxygen and, possibility, carbon dioxide and water, characterized in that: -- the catalyst used is a noble metal catalyst supported on a solid carrier, arranged as a plurality of fixed catalytic beds in cascade to each other; -- the gas feed stream contains methane, oxygen, carbon dioxide and water in the following molar proportions: methane 1.0; oxygen from 0.2 to 1.0; carbon dioxide from 0 to 3.0; water from 0 to 3.0; and -- the process is carried out under adiabatic conditions; by feeding the gas reactant stream upstream from the first catalytic bed and removing heat, by heat exchange between the catalytic beds arranged in cascade, or by feeding the gas reactant stream partially upstream from the first catalytic bed and partially, as a cold stream, between the catalytic beds arranged in cascade, with said partial feeds being of same composition, or having different compositions from each other, with the proviso that methane is at least partially fed to the first catalytic bed and oxygen is subdivided between all of the catalytic beds.
The catalysts useful for the process according to the present invention are constituted by one or more metals from platinum group, selected from Rh, Ru, Ir, Pt and Pd, supported on a carrier selected from aluminum, magnesium, zirconium, silicon, cerium and/or lanthanum oxides and/or spinels.
Said carrier can also be provided with surfacegrafted silica moieties, and suitable processes for preparing such carriers with surface-grafted silica moieties are reported in the experimental examples supplied in the following in the present application, in the above mentioned Italian patent applications and in United Kingdom patent application GB 2,240,284.
Preferred carriers for such catalysts are alumina and/or magnesium oxide, possibily provided with surface-grafted silica moieties.
The catalysts of the first catalytic bed contain rhodium in association with platinum or palladium, and the catalysts of the subsequent catalytic beds preferably contain two metals selected from rhodium, ruthenium and iridium, with the overall percent contents of noble metals in the supported catalyst being comprised within the range of from 0.05 to 1.5% by weight, and preferably of from 0.1 to 1% by weight.
In order to be used as a stationary catalytic bed, the catalysts will preferably be in granular form, with particle size comprised within the range of from 1 to 20 mm.
The catalytic beds used will be at least two, with their maximal number, dictated by practical reasons, being of four or five. Preferably, the process will be carried out with either two or three catalytic beds in series to each other. These catalytic beds can be arranged inside a plurality of reactors arranged in series to each other, but preferably, one single reactor containing a plurality of catalytic beds will be used.
According to the present invention, to the catalytic beds a gas stream is fed which contain methane and oxygen, and possibly also carbon dioxide and/or water, preferably in the following molar proportions: methane 1.0; oxygen 0.4-0.6; carbon dioxide 0-1.0; and water 0-1.0.
As said hereinabove, the process is carried out under adiabatic conditions by feeding the gas reactant stream totally upstream from the first bed and removing heat, by heat exchange, from points between the catalytic beds arranged in cascade.
According to a preferred embodiment, the process is carried out under adiabatic conditions by feeding the gas reactant stream partially upstream from the first catalytic bed and partially, as a cold stream, between the catalytic beds arranged in cascade. The gas streams fed to the individual catalytic beds can have the same composition, or compositions different from one another. In the latter case, methane will be at least partially fed to the first catalytic bed and the oxygen feed stream will suitably be subdivided between all catalytic beds.
In any case, by operating according to the present invention, synthesis gas is obtained by the effect of partial methane oxidation, and, possibly, also owing to reforming phenomena, as a function of the fed reactants.
According to an embodiment of the present invention, to the first catalytic bed a gas stream is fed which contains methane, oxygen, carbon dioxide and steam, and to the subsequent cataLytic beds an oxygen stream is fed. Preferably, the process will be carried out with a molar ratio of methane, carbon dioxide and water fed to the first catalytic bed, of 1:0.5-1:0.31, and with a total oxygen amount of 0.4-0.6 mols per each methane mol, fed as subdivided streams to each of the several catalytic beds.
According to another embodiment, to the first catalytic bed a gas stream is fed which contains methane and oxygen, and to the subsequent catalytic beds a mixture is fed which contains methane, oxygen and carbon dioxide. Preferably, the process will be carried out with a molar ratio of methane to oxygen fed to the cataLytic beds of the order of 1:0.4, and with an amount of carbon dioxide of the order of 0.4 mols per each mol of methane.
According to a further embodiment, to the first catalytic bed, and to the subsequent ones, a gas stream is fed which contains methane, oxygen and carbon dioxide. The molar ratios of these reactants to each other will preferably be of the order of 1:0.6:0.7-0.8.
According to a further embodiment, to the first catalytic bed a gas stream is fed which contains methane, oxygen and carbon dioxide, and to the subsequent catalytic bed an oxygen stream will be fed.
The process will preferably be carried out with a molar ratio of methane to carbon dioxide fed to the first catalytic bed of 1:0.3-0.6, and with a total oxygen amount of 0.5-0.6 mol per each mol of methane, subdivided to the various catalytic beds.
It should be observed that according to the present disclosure, the term "oxygen" is understood to mean pure or substantially pure oxygen, or oxygen mixed with an inert gas, such as nitrogen, e.g., air.
In general, the process will be carried out with inlet temperatures to the first bed of the order of 300-4000C and with outlet temperatures from said first bed, of the order of 700-8700C. The inlet temperatures to the beds downstream from the first bed will be of the order of 450-7300C, and the outlet temperatures will be of the order of 770-8500C. The cooling between two adjacent beds will cause a decrease in temperature of from 100C, up to as high values as 4200C and will normally be of the order of 120-170 C. The pressures under which the process is carried out may generally be comprised within the range of from 0.1 to 10 MPa.
The space velocities, under the reaction conditions, may generally be comprised within the range of from 1,000 to 50,000 h-1 and will normally be of the order of 5,000-20,000 h-l.
By operating under these conditions, the mixture recovered at the outlet from the last catalytic bed, will contain hydrogen and carbon monoxide in a molar ratio to each other comprised within the range of from about 0.9 to about 3 and normally of from about 1 to about 2.3.
It should be observed that in the case of exothermic reactions like the reaction of partial hydrocarbon oxidation 13, the expected reactant conversion rates as calculated by means of equilibrium thermodynamic computations, vary as a function of temperature, according to the trend schematically shown in Figure 1 of the accompanying drawing tables.
On the other hand it is known (0. Levenspiel, "Chemical Reaction Engineering", John Wiley and Sons, Inc., New York London) that the conversion rates, the reaction temperature and the reaction speed are mutually linked parameters. For exothermic reversible reaction (like the partial oxidation reaction C13) which are catalyzed in a "Plug-Flow" reactor, a temperature increase kinetically favours the transformation of the reactants into the reaction products, but, opposite to this trend, the temperature increase decreases the maximal conversion rate which can be obtained. In these cases, the optimal temperature variation can be obtained in reactors with a plurality of adiabatic layers with intermediate coolings induced by means of heat exchanges with heat recovery, or by means of the introduction of "cold" gas streams of reactants between the Layers.In Figure 1, "isospeed" curves are reported (i.e., curves a Long which the reaction speed remains constant with varying values of temperature and of reactants conversion), according to the typical trend of exothermic processes. The peak points of isospeed lines determine pairs of values of temperature (T) and conversion (Xa). The line which connects all of these points with each other (i.e., the line which makes it possible the maximal reaction speed values to be obtained with varying temperature) describes the optimal temperature progression for a Plug-Flow reactor in which an exothermic chemical process is being carried out.
Similar considerations may be made in the case of endothermic processes. Such a curve can be experimentally followed by means of a catalytic, adiabatic-layer reactor provided with a plurality of reaction zones separated by temperature adjustment zones, as in the case of the process disclosed herein.
The following experimental examples are reported in order to better illustrate the present invention.
Example 1 A laboratory reactor is used which is provided with two reaction zones, to which two different catalysts are charged.
The reactor was so accomplished as to make it possible the reactants (mixtures of methane, oxygen, steam and carbon dioxide) to be fed both to the reactor head, directly to the first catalytic bed (first adiabatic layer), and in the separation zone between both catalytic beds (i.e., between the first and the second adiabatic layers).
The reactor is constituted by an alumina tube with an extremely low porosity and displaying high heat resistance and mechanical strength characteristics. The alumina tube was fitted into a steel jacket. Around the steel tube, in the region of both reaction zones, two resistors are installed, the function of which is of compensating for the heat losses caused by the non-perfect adiabatic character of the reactor (this is a drawback which is impossible to remove in such a type of testing in small-size laboratory reactors). Inside the alumina tube, there is fitted a thermocouple well. The steel sheath of the thermocouple well was coated with a thin gold layer in order to prevent coal from being formed on its surface. The temperatures inside both adiabatic layers were measured with the aid of two thermocouples which could be longitudinally moved along said beds.
The two catalysts used in these tests were prepared according to the following procedures.
Catalyst for the first reaction zone (first adiabatic layer).
Into a slurry constituted by a suspension of alpha-alumina in n-hexane, a solution of Rh4 (CO)iz and Pd(C5 H5 02)23 in the same solvent, was added dropwise.
The solvent was then evaporated under vacuum and, after drying, the solid powder was pressed into pellets which, by crushing, yielded a granular solid with maximal particle diameter comprised within the range of from 2 to 2.5 mm. The catalyst volume charged to the first catalytic bed is of 5 cm3, the Rh content in the catalyst is of 0.1% by weight, the palladium content is of 0.5% by weight.
Catalyst for the second reaction zone (second adiabatic layer) In this case, a typical carrier for steam reforming catalysts was prepared, which contains magnesium oxides and alumina (Mg/Al = 7/1 mol/mol), and was obtained by means of a process comprising: (i) co-precipitating aluminum and magnesium hydroxides, by increasing the pH value of an aqueous solution of Mg(NO3 )2 and Al(NO3)3.9HzO; (ii) filtering the precipitate off and washing it; (iii) drying and calcining the precipitate at 4000C, (iv) "pelletizing" the solid powder; (v) treating the pellets by further calcining them up to 10000C and, after cooling, crushing the pellets in order to obtain a granular material with a maximal particle diameter of 2-2.5 mm.
The percent sodium level in the resulting carrier is lower than 0.1%. The carrier was then dispersed in a solution of n-hexane into which a solution, in the same solvent, of Rh4 (CO)iz and Ru3(C0)1z had been added dropwise. After evaporation and vacuum drying, a granular material was obtained which contained 0.1% by weight of Rh and 0.5% by weight of Ru. The catalyst volume charged to the second catalytic bed is of 5 cm3.
Prior to the reaction, the catalysts were treated at the temperature of 5000C, with H2/N2 streams containing increasing hydrogen levels. Then, to the inlet to the first catalytic bed a stream was fed which contained CH4:CO2:02:H20 in molar ratios of 1:1:0.5:0.3. The total flowrate of feedstock fed to the first catalytic bed was of 50 Nl/hour, the gas stream inlet temperature was kept at 3000C, the inner reactor pressure was kept at 10 atm. Before entering the second adiabatic layer, the leaving stream from the first catalytic bed was mixed with a second stream of oxygen pre-heated at 3000C, fed at a flowrate of 2.3 Nl/hour.
In Table 1, the main features of this experiment are reported.
TABLE 1 Ist adiabatic layer Catalyst: -- composition: Rh (0,1%) + Pt (0,5X) on Al203 -- amount: 5 cc Inlet composition: -- CH :CO2 :02 :H20 = 1:1:0.5:0.3 (volume ratios) Feed flowrate: -- CH4 = 17.90 Nl/hour -- CO2 = 17.90 Nl/hour -- Oz = 8.70 Nl/hour -- H20 = 5.30 Nl/hour -- total = 50.00 Nl/hour Temperatures: -- inlet = 300 C -- outlet = 7450C Lind adiabatic Laver Catalyst: -- composition:Rh (0.1%)+Ru (0.5%) on MgAlOx Inlet composition: -- gas product from the Ist layer + added 02 -- O2 feed flowrate: 2.30 Nl/hour Temperatures: -- inlet = 7300C -- outlet = 8 1 0o C Composition at reactor outlet % by mol Mols/hour -- CH4 5.20 0.16 -- COz 23.46 0.73 -- H20 21.59 0.67 -- 02 --- -- H2 27.04 0.84 -- CO 22.68 0.71 Molar ratio of H2:CO at reactor outlet: 1.18:1.
Example 2 The same experimental devices and the same catalysts as disclosed in experiment 1 were used, by feeding to the inlet to the first catalytic bed a reactant stream with a total flowrate of 50 Nl/hour and having the composition CH4:CO2:0z:H20 = 1:0.5:0.4:1 and feeding, upstream from the second catalytic bed, a stream of oxygen pre-heated at 3000C, with a flowrate of 3 Nl/hour.
The main features of this second experiment are reported in Table 2.
TABLE 2 Ist adiabatic layer Catalyst: -- composition: Rh (0,1%) + Pt (0,5%) on ALz03 -- amount: 5 cc Inlet composition: -- CH4:C02:02:H20 = 1:0.5:0.4:1 (volume ratios) Feed flowrate: -- CH4 = 17.20 Nl/hour -- CO2 = 8.60 Nlihour -- 02 = 7.00 Nl/hour -- H20 = 17.20 Nl/hour -- total = 50.00 Nl/hour Temperatures: -- inlet = 3000C -- outlet = 7050C IInd adiabatic layer Catalyst: -- composition:Rh (0.12) + Ru (0.5X) on MgAlOx -- amount: 3 cc Inlet composition: -- gas product from the Ist layer + added 02 -- 0z feed flowrate: 3.00 Nl/hour Temperatures: -- inlet = 6900C -- outlet = 8050C Composition at reactor outlet: by mol Mols/hour -- CH4 5.10 0.16 -- C0z 16.60 0.52 -- H20 29.27 0.92 -- 02 -- H2 34.11 1.07 -- CO 14.93 0.47 Molar ratio of Hz :CO at reactor outlet: 2.28:1.
Example 3 In this experiment, the same exerimental devices as disclosed in Examples 1 and 2 were used, but catalysts were used which contained noble metals deposited on alumina with surface-grafted silica moieties and magnesium carriers.
Catalyst for the first reaction zone (first adiabatic layers.
A commercial alumina sypplied by AKZO, having a surface area of approximately of 200 mz /g was suspended, with stirring, in a tetraethyl silicate (TES) solution. The temperature was kept comprised within the range of from 80 to 900 C. Under these conditions, a trans-esterification reaction took place which is represented by equation C83 and ted to the development of ethanol in gas form: Si(OCzHs)4 + AL-OH
Al-O-Si(OC2Hs)3 + C2Hs-OH A gas stream of anhydrous nitrogen was fed to the reaction environment. Gas-chromatographic analyses on the leaving gas showed that ethanol had been formed.
The reaction was regarded as concluded when in the gas stream the presence of ethanol was no longer detectable. At this point, the temperature was increased up to 1800C, in order to distil off any unreacted TES. The unreacted ethoxy groups bonded to siLicon atoms which, in their turn, were anchored to the surface, were then hydrolized by feeding, at 2000C, a nitrogen stream saturated with steam The so obtained solid material was heated up to 8000C and was kept at this temperature during 10 hours. After cooling, the material was used as a .carrier, onto which rhodium and platinum were deposited. The finished catalyst contained 0.1% of rhodium and 0.5% by weight of platinum.
Catalyst for the second reaction zone (second adiabatic ic laer2 The surface silica-grafting process as disclosed above was repeated on a carrier of commercial magnesium oxide having a surface area of 150 m2/g.
Onto this magnesium oxide with surface-grafted silica moieties obtained by means of this procedure, 0.1S by weight of Rh and 0.5X by weight of Ru were then deposited according to the same procedure as disclosed in Example 1.
The catalytic test was carried out according to the same procedure as disclosed in Examples 1 and 2.
After a reducing treatment, a stream containing CH4:CO2:02:H20 in molar ratios of 1.0:1.0:0.4:1.0 was fed to the inlet to the first catalytic bed. Before entering the second catalytic bed, the stream leaving from the first catlytic bed was admixed with an oxygen stream fed at a flowrate of 1.8 Nl/hour.
The main features of this experiment are disclosed in Table 3.
TABLE 3 I~t~a~ia~atic~layer Catalyst: -- composition: Rh (0,1%) + Pt (0,52) on silica grafted alumina -- amount: 5 cc Inlet composition: -- CH4:CO2:02:H20 = 1.0:1.0:0.4:1.0 (volume ratios) Feed flowrate: -- CH4 = 14.70 Nl/hour -- C0z = 14.70 Nl/hour -- 02 = 5.90 Nl/hour -- HzO = 14.70 Nl/hour -- total = 50.00 Nl/hour Temperatures: -- inlet = 3000C -- outlet = 6980C lind adiabatic layer Catalyst: -- composition:Rh (0.1%) + Ru (0.5) on silica grafted magnesium oxide -- amount: 3 cc Inlet composition: -- gas product from the Ist layer + added 02 -- 02 feed flowrate: 1.47 Nl/hour Temperatures: -- inlet = 6850C -- outlet = 7900C Composition at reactor outlet: by mol Mols/hour -- CH4 4.41 0.13 -- C0z 21.11 0.64 -- H20 26.83 0.81 -- Oz -- Hz 29.65 0.90 -- CO 18.01 0.55 Molar ratio of HE CHO at reactor outlet: 1.64:1.
Example 4 In this experiment, to the first catalytic bed, a volume of 5 cm3 was charged of a catalyst containing 0.1 by weight of Rh and 0.5% by weight of Pd. The metals were deposited according to the same procedure as disclosed in Example 1, on a carrier constituted by magnesium and aluminum oxides (Mg:Al = 7:1 mol/mol), using a solution containing Rh4(CO)12 and Pd(CsHsOz)z3 in a hydrocarbon solvent.
To the second catalytic bed, a volume of 4 cm3 was then charged of a catalyst containing 0.5% by weight of Ru and 0.5% by weight of Ir, deposited on magnesium and aluminum mixed oxide. The deposition of these metals onto the carrier was accomplished by adding, dropwise, a solution of Ir4 (CO)12 and Ru3(C0)12 in a hydrocarbon solvent, to a suspension of the carrier in the same solvent, as disclosed in Example 1.
After a treatment in a H2-N2 stream at 5000C, a stream of CH4 and Oz (CH4:02 = 60:25 by vol/vol) was added to the first catalytic bed, and upstream from the second catalytic bed, a stream of CH4 , Oz and CO (CH4:0z:CO2 = 40:25:40 by vol/vol) was admixed to the gas stream from the first catalytic bed.
The main features obtained during the cataLytic test are reported in Table 4.
TABLE 4 Ist adiabatic layer Catalyst: -- composition: Rh (0,1%) + Pt (O,SX) on Mc7AlOx -- amount: 5 cc Inlet composition: -- CH4 :02 = 60:25 (volume ratios) Feed flowrate: -- CH4 = 15.78 Nl/hour -- 02 = 6.60 Nl/hour -- total = 22.38 Nl/hour Temperatures: -- inlet = 3000C -- outlet = 7450C Ind adiabatic Layer Catalyst: -- composition:Ir (0.5%) + Ru (0.5%) on HgAlOy -- amount: 4 cc Inlet composition: -- gas product from the Ist layer + CH4 + Oz + COz added -- feed flowrate: -- CH4 = 10.52 Nl/hour -- 02 = 6.50 Nl/hour -- COz = 10.50 Nl/hour -- total = 27.52 Nl/hour Temperatures: -- inlet = 581CC -- outlet = 8150C Composition at reactor outlet: X by mol Mols/hour -- CH4 13.95 0.43 -- COz 14.47 0.45 -- H20 14.90 0.46 -- Oz -- HZ 32.40 1.01 -- CO 24.28 0.76 Molar ratio of H2:CO at reactor outlet: 1.33:1.
Example 5 In this case, the process of catalytic partial oxidation in an adiabatic reactor with layer configuration was studied by using three Plug-Flow reactors (which are referred to in the following as "R1", "R2", "R3"), each containing one catalytic bed.
The characteristics of said three reactors are analagous to those as of the reactor disclosed in Figure 4. A mixture of CH4, 02, COz, fed with a total gas flowrate of 149 Hl/hour (CH4 :Oz :COz = 1:0.6:0.8 by vol/vol) was subdivided into three streams. The first stream (flowrate 60.1 Nl/h) was fed to the inlet to reactor R1; the second stream (flowrate 53.3 Nl/h) was fed to a point between reactor R1 and reactor R2; the third stream (flowrate 35.6 Nl/h) was fed to a point between reactor R2 and reactor R3.
The temperature of the stream fed to the inlet to the first reactor was kept at 3000C, and the inlet temperatures to the second and third reactors were kept at 4500C. The catalyst contained in reactor R1 (catalyst volume: 3 cm3) was composed by Rh (0.12 by weight) and Pd (0.5M by weight) deposited on a support constituted by a mixed magnesium and aluminum oxide, prepared by operating according to the same procedure as disclosed in Example 1.
The catalyst contained in reactor R2 (catalyst volume: 4 cm3) was composed by Rh (0.1% by weight) and Ir (0.5% by weight), deposited on the same carrier of magnesium and aluminum oxides. The catalyst was prepared according to the same procedure as disclosed in Examples 1 and 3. The catalyst contained in R3 was composed by Rh (0.1% by weight) and Ru (0.5% by weight), deposited, also in this case, onto the same magnesium and aluminum oxide. The catalyst was prepared according to the same procedures as disclosed in Example 1.
In Table 5, the main features and the results of the present experiment are reported.
TABLE 5 Ist adiabatic layer Catalyst: -- composition: Rh (0,1%) + Pt (0,5%) on MgAlOy -- amount: 3 cc Inlet composition: -- CH4:O2:CO2 = 100:60:80 (volume ratios') Feed flowrate: -- CH4 = 25.10 Nl/hour -- COz = 20.00 Nl/hour -- 02 = 15.00 Nl/hour -- total = 60.10 Nl/hour Temperatures: -- inlet = 3000C -- outlet = 8650C IInd adiabatic layer Catalyst: -- composition:Rh (0.12.)+Ir (0.5 x) on MgAlOx -- amount: 4 cc Inlet composition: -- gas product from the Ist layer + CH4 + Oz + COz added -- feed flowrate: -- CH4 = 22.6 Nl/hour -- 02 = 17.5 Nl/hour -- COz = 13.2 Nl/hour -- total = 53.3 Nl/hour Temperatures: -- inlet = 4500C -- outlet = 8250C Iliad adiabatic layer Catalyst: -- composition:Rh (0.1%) + Ru (0.5%) on MgAlOx -- amount: 5 cc Inlet composition: -- gas product from the IInd layer + CH4 + 02 + COz added -- feed flowrate: -- CH4 = 15.0 Nl/hour -- 02 = 11.9 Nl/hour -- COz = 8.7 Nl/hour -- total = 35.6 Nl/hour Temperatures: -- inlet = 4500C -- outlet = 785oC Composition at reactor outlet : x by mol Mols/hour -- CH4 5.74 0.54 -- C0z 18.23 1.82 -- HzO 16.89 1.59 -- 0z -- H2 30.33 2.87 -- CO 28.84 2.72 Molar ratio of H2 : CO at reactor outlet: 1.055:1.
Examples 6-8 The same experimental apparatus and the same catalysts as disclosed in Example 5 were used in Examples 6, 7 and 8 in order to obtain a catalytic partial oxidation process on a three-layer catalyst, to which a feedstock consisting of methane, COz and oxygen was fed. In these cases, differently from the experiment as disclosed in Example 5, the whole amounts of CH4 and COz were fed to the inlet to the first reactor R1, and the oxygen feed was subdivided into three streams which were fed to the inlet of R1, to an intermediate point between R1 and R2, and to an intermediate point between R2 and R3. Examples 6, 7 and 8 are different from each other owing to the inlet temperatures of the gas streams to the three adiabatic layers. Different inlet temperatures to the adiabatic layers have determined different temperatures and composition of the bed leaving streams.
In following Tables 6, 7 and 8, the main features and the results obtained in Examples 6, 7 and 8 are reported.
TABLE 6 1st adiabatic laye Catalyst: -- composition: Rh (0,1%) + Pt (0,5%) on MgAlOx -- amount: 4 cc Inlet composition: -- CHr:02:CO2 = 100:30:60 (volume ratios) Feed flowrate: -- CH4 = 68.30 NL/hour -- COz = 41.00 NlShour -- 02 = 20.50 Nl/hour -- total = 129.80 Nl/hour Temperatures: -- inlet = 3000C -- outlet = 710C IInd adiabatic layer Catalyst: -- composition:Rh (0.1%) + Ir (0.5x) on MgAlOy -- amount: 4 cc Inlet composition: -- gas product from the Ist layer + 02 added -- feed flowrate: -- 02 = 13.6 Nl/hour -- total = 13.6 Nl/hour Temperatures: -- inlet = 4500C -- outlet = 7750C IlIrd adiabati layer Catalyst: -- composition:Rh (0.1%) + Ru (0.5%) on MgAlOx -- amount: 5 cc Inlet composition: -- gas product from the IInd layer + Oz added -- feed flowrate: -- 02 = 6.8 Nl/hour -- total = 6.8 Nl/hour Temperatures: -- inlet = 4500C -- outlet = 7780C Composition at reactor outlet: % by mol Mols/hour -- CH4 7.2 0.69 -- COz 16.1 1.54 -- HzO 16.6 1.59 -- 0z -- H2 32.6 3.12 -- CO 27.6 2.64 Molar ratio of Hz :CO at reactor outlet: 1.1818:1.
TABLE 7 1st adiabatic layer Catalyst: -- composition: Rh (0,1%) + Pt (0,5%) on MgAlOx -- amount: 4 cc Inlet composition: -- CH4:02:CO2 = 100:30:60 (volume ratios) Feed flowrate: -- CH4 = 68.30 Nl/hour -- COz = 41.00 Nl/hour -- 02 = 20.50 Nl/hour -- total = 129.80 NL/hour Temperatures: -- inlet = 3000C -- outlet = 715 C lInd adiabatic layer Catalyst: -- composition:Rh (0.1%) + Ir (0.5%) on MgAlOx -- amount: 4 cc Inlet composition: -- gas product from the Ist layer + Oz added -- feed flowrate: -- Oz = 13.6 Nl/hour -- total = 13.6 Nl/hour Temperatures: -- inlet = 550 C -- outlet = 7970C Iliad adiabatic layer Catalyst: -- composition:Rh (0.1%) + Ru (0.5%) on MgAlOy -- amount: 5 cc Inlet composition: -- gas product from the IInd layer + 02 added -- feed flowrate: -- 02 = 6.8 Niinour -- total = 6.8 Nl/hour Temperatures: -- inlet = 5500C -- outlet = 8160C Composition at reactor outlet: % by mol Mols/hour -- CH4 4.6 0.46 -- COz 16.1 1.34 -- H20 15.6 1.56 -- O2 -- H2 35.9 3.60 -- CO 30.6 3.07 Molar ratio of Hz :CO at reactor outlet: 1.172:1.
TABLE 8 Ist adiabatic layer Catalyst: -- composition: Rh (0,1%) + Pt (0,5%) on MgAlOy -- amount: 4 cc Inlet composition: -- CH4:02:CO2 = 100:30:60 (volume ratios) Feed flowrate: -- CH4 = 68.30 Nl/hour -- COz = 41.00 Nl/hour -- 02 = 20.50 Nl/hour -- total = 129.80 Nl/hour Temperatures: -- inlet = 4000C -- outlet = 7220C I I nd adiabatic layer Catalyst: -- composition: Rh (0.1%) + Ir (0.5') on MgAlOX -- amount: 4 cc Inlet composition: -- gas product from the Ist layer 02 added -- feed flowrate: -- 02 = 13.6 Nl/hour -- total = 13.6 Nl/hour Temperatures: -- inlet = 6000C -- outlet = 8120C IlIrd adiabatic Layer Catalyst: -- composition:Rh (0.1%) + Ru (0.5%) on MgAlOx -- amount: 5 cc Inlet composition: -- gas product from the IInd layer + 02 added -- feed flowrate: -- 02 = 6.8 Nl/hour -- total = 6.8 Nl/hour Temperatures: -- inlet = 6000C -- outlet = 8410C Composition at reactor outlet- % by mol Mols/hour -- CH4 3.3 0.34 -- C0z 11.9 1.22 -- H20 15.1 1.55 -- Oz -- H2 37.6 3.87 -- CO 32.2 3.31 Molar ratio of H2:CO at reactor outlet: 1.169:1.
Example 9 The same experimental apparatus as disclosed in Examples 5-8 was used in order to study the reactions of catalytic partial oxidation of mixtures of CH4 :02 :COZ = 100:60:30 (by vol/vol). In this case, the content of COz was kept at lower values than as in the preceding examples. Also in this case, the oxygen stream was subdivided into partial streams which were fed both to the inlet to R1, and to an intermediate point between R1 and R2, as well as to an intermediate point between R2 and R3. Furthermore (by pre-heating the gas reactant streams), inlet temperatures to the catalytic beds were tested which were higher than in the preceding examples. The catalyst used in reactor R1 (Ist adiabatic layer) contained Rh (0.1% by weight) and Pt (0.5% by weight) deposited on a mixed aluminum and magnesium oxide. The preparation procedures used have already been disclosed in the preceding examples.
The catalysts contained in the second reactor (R2) and in the third reactor (R3) (i.e., the second and third adiabatic layers) were the same as used in Examples 5-8 and contained Rh and, respectively, Ir, deposited on an aluminum and magnesium oxide, and Rh and Ru deposited on the same support.
In following Table 9, the main features of the experiment are reported.
TABLE 9 I adiabatic layer Catalyst: -- composition: Rh (0,1%) + Pt (0,52) on MgALOx -- amount: 4 cc Inlet composition: -- CH4:02:CO2 = 100:30:30 (volume ratios) Feed flowrate: -- CH4 = 79.00 Nl/hour -- COz = 23.70 Nl/hour -- 02 = 23.70 Nl/hour -- total = 126.40 Nl/hour Temperatures: -- inlet = 4000C -- outlet = 7610C IInd adiabatic layer Catalyst: -- composition: Rh (0.1%) + Ir (0.5%) on MgAlOx -- amount: 4 cc Inlet composition: -- gas product from the Ist layer Oz added -- feed flourate: -- 02 = 15.8 Nl/hour -- total = 15.8 Nl/hour Temperatures: -- inlet = 6000C -- outlet = 8530C Iliad agiabatic Layer Catalyst: -- composition: Rh (0.1%) + Ru (0.5%) on MgAlOx -- amount: 5 cc Inlet composition: -- gas product from the IInd layer + 02 added -- feed flowrate: -- 0z = 7.9 Nl/hour -- total = 7.9 Nl/hour Temperatures: -- inlet = 6000C -- outlet = 8410C Composition at reactor outlet: % by mol Mols/hour -- CH4 3.1 0.34 -- C02 6.9 0.76 -- H20 12.3 1.34 -- Oz -- -- HZ 45.9 5.03 -- CO 31.8 3.48 Molar ratio of H2:CO at reactor outlet: 1.445:1.

Claims (17)

1. A catalytic process for preparing a synthesis gas from methane, oxygen, and, optionally, carbon dioxide and/or water, wherein: the catalyst used is a noble metal catalyst supported on a solid carrier, arranged as a plurality of fixed catalytic beds in cascade to each other; the gas feed stream comprises methane, oxygen, and, optionally, carbon dioxide and/or water in the following molar proportions: methane 1.0; oxygen from 0.2 to 1.0; carbon dioxide from 0 to 3.0; water 0 to 3.0; and the process is carried out under adiabatic conditions either: (a) by feeding the gas reactant stream upstream from the first catalytic bed and removing heat by heat exchange between the catalytic beds arranged in cascade, or (b) by feeding the gas reactant stream partially upstream from the first catalytic bed and partially between the catalytic beds arranged in cascade, with said partial feeds being of the same or different compositions, with the proviso that methane is at least partially fed to the first catalytic bed and oxygen is subdivided between all of the catalytic beds.
2. A process according to claim 1, wherein the gas feed stream contains the reactants in the following molar proportions: methane 1.0; oxygen from 0.4 to 0.6; carbon dioxide from 0 to 1.0; and water from 0 to 1.0.
3. A process according to claim 1 or 2; wherein the catalyst comprises one or more platinum group metals selected from Rh, Ru, Ir, Pt and Pd, supported on a carrier selected from aluminum, magnesium, zirconium, silicon, cerium and/or lanthanum oxides and/or spinels, or from silica-treated products of such carriers.
4. A process according to claim 3, wherein the catalyst of the first catalytic bed comprises rhodium in association with platinum or palladium, and the catalyst(s) of the subsequent catalytic bed contain two metals selected from rhodium, ruthenium and iridium, with the overall content of nobel metals in the supported catalyst being from 0.05 to 1.5% by weight, preferably from 0.1 to 1% by weight.
5. A process according to any of claims 1 to 4, wherein the catalysts are in granular form with a particle size of from 1 to 20 mm, and are arranged in at least two and up to five catalytic beds, preferably either two or three catalytic beds.
6. A process according to any of claims 1 to 5, wherein, to the first catalytic bed, there is fed a gas steam which comprises methane, oxygen, carbon dioxide and stream, and to the subsequent catalytic bed there is fed a stream comprising oxygen.
7. A process according to claim 6, wherein the process is carried out with a molar ratio of methane, carbon dioxide and water fed to the first catalytic bed of 1:0.5-1:0.3-1, and with a total amount of oxygen of 0.4-0.6 mol per mol of methane, subdivided between the catalytic beds.
8. A process according to any of claims 1 to 5, wherein, to the first catalytic bed, there is fed a gas stream which comprises methane and oxygen, and, to the subsequent catalytic bed(s), there is fed a mixture which comprises methane, oxygen and carbon dioxide.
9. A process according to claim 8, wherein the process is carried out with a molar ratio of methane to oxygen fed to the catalytic beds of the order of 1:0.4, and with an amount of carbon dioxide of the order of 0.4 mol per mol of methane.
10. A process according to any of claims 1 to 5, wherein, to the first catalytic bed and to the subsequent bed(s), there is fed a gas stream which comprises methane, oxygen and carbon dioxide.
11. A process according to claim 10, wherein the process is carried out with molar ratios of said reactants of the order of 1:0.6:0.7-0.8.
12. A process according to any of claims 1 to 5, wherein, to the first catalytic bed, there is fed a gas stream which comprises methane, oxygen and carbon dioxide, and, to the subsequent catalytic bed(s), there is fed a stream comprising oxygen.
13. A process according to claim 12, wherein the process is carried out with a molar ratio of methane to carbon dioxide fed to the first catalytic bed of 1:0.30.6, and with a total amount of oxygen of 0.5-0.6 mol per mol of methane, subdivided between the catalytic beds.
14. A process according to any of claims 1 to 13, wherein the process is carried out with an inlet temperature to the first bed of the order of 300-400"C and with outlet temperatures from said first bed of the order of 700-870"C, with an inlet temperature, to the bed(s) downstream from the first bed, of the order of 450-730"C and outlet temperatures of the order of 770 850"C, with the cooling between two adjacent beds causing a temperature decrease of from at least 10 C up to 420"C, and preferably of the order of 120170CC.
15. A process according to any of claims 1 to 14, wherein the process is carried out under a pressure of from 0.1 to 10 MPa and with a space velocity value, under the reaction conditions, of from 1,000 to 50,000 h-l
16. A process according to claim 1, substantially as described in any of the Examples.
17. Synthesis gas prepared by a process according to any of claims 1 to 16.
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US6911193B2 (en) * 2002-04-19 2005-06-28 Conocophillips Company Integration of mixed catalysts to maximize syngas production
US7226548B2 (en) 2002-11-11 2007-06-05 Conocophillips Company Syngas catalysts and their method of use
US7074375B2 (en) * 2002-12-03 2006-07-11 Engelhard Corporation Method of desulfurizing a hydrocarbon gas by selective partial oxidation and adsorption
US7799451B2 (en) 2004-08-05 2010-09-21 Rolls-Royce Fuel Cell Systems (Us) Inc. Post-reformer treatment of reformate gas
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US7510793B2 (en) 2004-08-05 2009-03-31 Rolls-Royce Fuel Cell Systems (Us) Inc. Post-reformer treatment of reformate gas
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US7261751B2 (en) 2004-08-06 2007-08-28 Conocophillips Company Synthesis gas process comprising partial oxidation using controlled and optimized temperature profile
US8507566B2 (en) 2006-09-08 2013-08-13 Gelato Corporation N.V. Process for the preparation of synthesis gas
WO2010020358A3 (en) * 2008-08-21 2011-02-03 Uhde Gmbh Multi-stage reactor cascade for the soot-free production of synthesis gas
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US8685282B2 (en) 2010-05-04 2014-04-01 KT—Kinetics Technology S.p.A. Process for the production of syngas and hydrogen starting from reagents comprising liquid hydrocarbons, gaseous hydrocarbons, and/or oxygenated compounds, also deriving from biomasses, by means of a non-integrated membrane reactor
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GB9326099D0 (en) 1994-02-23
CA2112519A1 (en) 1994-06-24
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CN1089232A (en) 1994-07-13
NO934736L (en) 1994-06-24

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