Process for the production of high BTU methane-containing gas

Abstract

A process for the production of high Btu methane containing gases by dividing the effluent, having a substantial CO content, from fossil fuel gasification (preferably at a much higher temperature than 500.degree.C) from which C and S have been removed into a plurality of fractions; subjecting one effluent fraction to shift conversion and CO.sub.2 removal; subjecting another effluent fraction to a methanation step; joining the effluents from the shift conversion and methanation steps and subjecting them to a second methanation step, utilizing therein the N.sub.2 --H.sub.2 --CH.sub.4 mixture from the shift conversion; and separating the excess nitrogen from the product methane gas by cryogenic removal.

Description

BACKGROUND OF THE INVENTION

The present invention generally relates to the production of high Btu methane-containing gases having a heating value in the range of about 800 to about 1000 Btu/SCF, and more particularly to a process for producing such gases from the gaseous effluent resulting from the gasification of fossil fuels such as gas, oil, etc. By means of this invention, the effluent from gasification, upon removal of carbon and sulfur therefrom, can be divided into a plurality of fractions or streams, whereby, in a first fraction or stream, the hydrogen contained therein is used to methanate part of the CO and CO.sub.2 ; whereby, another or second effluent fraction or stream is subjected to shift conversion, followed by CO.sub.2 removal, to result in a H.sub.2 --N.sub.2 --CH.sub.4 mixture; and whereby the effluents of the first and second fractions or streams can be joined, with the hydrogen in said mixture serving to methanate the remaining carbon oxides from said first fraction stream, and the excess nitrogen removed from the resulting methane product by cryogenic separation.

Heretofore, those attempting to produce a methane gas product such as substitute natural gas from the effluent of oil or coal gasification have been confronted with the problem of having to convert the rich CO content thereof (e.g., 50% CO) into methane in the face of the fact that said conversion is a highly exothermic one. Traditional solutions to this problem have employed alternate heating and cooling treatments; used extensive recycle streams as a diluent to absorb some of the exothermic heat evolved; and employed other conventional methods, etc.

However, these previous methods have been costly, essentially because the methanation is carried out under conditions which require a high recycle gas rate and/or very large heat removal, which is not the case with respect to the first and second methanation steps of the present invention. Furthermore, in order to get a high Btu value in the final gas, the previous processes have used oxygen in their gasification step, produced in a costly air separation unit, whereas air or oxygen enriched air can be used in the gasification step of the present invention.

In another area of substitute natural gas production, several known processes are based on the reaction of steam with a light hydrocarbon, in a catalytic reactor operating adiabatically around 500.degree.C. The operating temperature is such that enough methane is synthesized in the reactor (exothermic reaction) to supply heat for the endothermic steam reforming reaction. The main drawback of such processes is that the feedstock must not be heavier than naphtha. In order to gasify feedstocks heavier than naphtha, it is recognized that much higher temperatures are needed, even if a catalyst is used. At such higher temperature, however, much less methanation is taking place in the reactor, and consequently an additional source of heat is required to supply the heat needed for the steam reforming reaction. In the present invention, however, this can be accomplished by the injection of air or oxygen enriched air in the reactor, which is also operating adiabatically.

Methods relating generally to the production of substitute natural gas are well known. For example, U.S. Pat. No. 3,347,647 discloses a process for the conversion of solid fossil fuel to high Btu pipeline gas which includes a two-stage hydrogasification reactor as well as a hydrogen plant for producing hydrogen from char residue by the so-called Texaco partial oxidation process. In this system, coal supplies all process feedstock requirements for the product gas and the reaction hydrogen, and countercurrent flow of the char and hydrogen in combination with dual reaction temperature zones enable the raw feed coal to enter the low temperature end of the reactor and mix with an atmosphere with a relatively high concentration of methane while hydrogen enters the high temperature end of the reactor and reacts with the partially gasified char where the methane concentration is relatively low, thereby providing for continuous production of high Btu methane. U.S. Pat. No. 3,347,647, however, is not cognizant of splitting its gasification effluent into a plurality of fractions each of which is subjected to a different treatment (methanation as opposed to shift conversion), whereby the effluents of the different treatments are combined and methanated, and the excess nitrogen removed from the product methane or substitute natural gas. Nor is U.S. Pat. No. 3,347,647 aware of the benefits to be derived from methanation of its carbon oxides in two separate steps.

Catalytic methanation is disclosed in U.S. Pat. No. 3,511,624 wherein a mixture of carbon oxides, hydrogen, steam and at least 25% volume of methane is passed, in a first stage, over a methanation catalyst after which the steam is partially removed from the mixture and the resulting mixture is passed, in a second stage, over a methanation catalyst which is at a temperature within a range lower than the temperature of the mixture leaving the first stage. However, while U.S. Pat. No. 3,511,624 discloses a two-stage catalytic methanation process, it is to be noted that such stages involve taking a feed gas already containing from 25-50% methane and 5-20% CO.sub.2, with little CO (0-5%) present and subjecting it to two consecutive methanation stages. Subjecting such a feed gas to such methanation steps is the means this patent teaches to avoid the vast amount of heat released that accompanies the exothermic reaction between hydrogen and the carbon oxides of more conventional feed gases which are rich in both of these gases such as the effluents of oil or coal gasification. Thus, the process of U.S. Pat. No. 3,511,624 is totally unlike that of the present invention which is capable of utilizing various feedstocks and which has two non-consecutive methanation steps involving feedstocks of different composition and other distinguishing process features.

Although individual process features or steps of the present invention are perhaps known such as, e.g., the cryogenic separation of natural gas, as shown, for example, in U.S. Pat. No. 3,616,652, none of the foregoing patents or the prior art have heretofore been aware of or appreciated the significant advance in the art to be associated with the process of the present invention and its combination and sequential arrangement of process steps.

SUMMARY OF THE INVENTION

The present invention relates generally to a process for producing a methane-rich gas or substitute natural gas from any conventional carbon or hydrocarbon-containing feedstock such as oil or coal comprising generating a gas effluent from said feedstock and removing the carbon solids and sulfur impurities (e.g., H.sub.2 S and COS) therefrom; dividing the gas effluent into a plurality of fractions, and methanating one fraction while concurrently shift converting another and removing the CO.sub.2 therefrom; combining or mixing the resulting effluents from methanation and shift conversion and subjecting the consequent mixture to methanation; and separating the excess nitrogen present from the methanerich, or substitute natural gas product.

BRIEF DESCRIPTION OF THE DRAWINGS

The process of the present invention is illustrated schematically in the drawings by means of flow diagrams, showing each step or stage of said process, wherein:

FIG. 1 is a block diagram showing the process steps of the present invention in their proper sequence;

FIG. 2 is a flow diagram of the gas generation step;

FIG. 3 is a flow diagram of the sulfur removal step;

FIG. 4 is a flow diagram of the first methanation step;

FIG. 5 is a flow diagram of the shift conversion step;

FIG. 6 is a flow diagram of the carbon dioxide removal step;

FIG. 7 is a flow diagram of the second methanation step; and

FIG. 8 is a flow diagram of the cryogenic separation step.

DESCRIPTION OF THE PREFERRED EMBODIMENTS

Shown in FIG. 2 are the essential features of the gas generation step, utilizing, as a feedstock, a liquid hydrocarbon. However, any hydrocarbon- or carbo-containing material can be used as a feedstock in the present gas generation step provided that such material can react with oxygen or a source of oxygen at high temperatures of upwards of 500.degree.C, preferably in the range of 650.degree.C to 1000.degree.C, to produce a gas mixture containing hydrogen, carbon oxides and methane. Thus, utilizable in the gas generation step are such feedstocks as coal; synthetic polymers such as waste polymeric materials, shale oil; naphtha or heavier hydrocarbon feedstocks, etc.

As an oxidant for use in the present gas generation, suitable oxygen sources such as air per se, air enriched with oxygen, steam, and/or mixtures thereof can be employed. Introduction of air, of course, involves addition of nitrogen to the process. The process step of "gas generation" should be understood to be applicable to a wide diversity of types, including, e.g., non-catalytic adiabatic gas generation, catalytic adiabatic gas generation, etc., and is intended to refer to any process that is able to convert a carboncontaining feedstock into a gas mixture, capable of being processed eventually into a methane-rich gas, preferably of high Btu content, in an adiabatic reactor, with the aid of air or oxygen-enriched air. A conventional catalyst can be utilized in the gas generation step, if desired, and such use enables the outlet temperature and corresponding nitrogen content of the exiting gaseous effluent to be lowered.

The gas generation step may also be denoted as "partial oxidation"; however, should a conventional catalyst be utilized therein, such step would then be generally denoted as "air reforming".

The hydrocarbon- or carbon-containing feedstock can be completely gasified, if desired; however, where heavier feedstocks are used, such as those heavier than gas oil, the feedstock may only be partially gasified; consequently the liquid hydrocarbon mixture would have to be separated by condensation before purification thereof by the subsequent sulfur removal step.

In the present gasification step, the hydrocarbon feedstock is pumped to the desired operating pressure through pump 1, and then is mixed with steam and preheated in the heater 2 to a temperature as high as possible, preferably in the range indicated above.

Air or oxygen-enriched air, for partial oxidation, or, if a conventional catalyst is used, air reforming, of the carbon or hydrocarbon feedstock, is compressed in a compressor 3 and preheated by means of heating apparatus 4 to a very high temperature that may range between 500.degree.C and 1000.degree.C. Alternative means for achieving this high preheat temperature of the air stream include such apparatus as fired heaters, fixed or moving bed pebble type heat exchangers, or heat exchange with the gas effluent, etc.

In an alternative embodiment of gas generation according to the present process, some steam may also be added to the air stream before preheating.

The preheated air or air-oxygen mixture and the preheated steam feedstock mixture are then fed into reactor 5, which is adapted to operate adiabatically, and reacted. Reactor 5 is entirely lined with refractory material and designed to operate at high temperatures, preferably ranging between 650.degree. and 1000.degree. C. This reactor can be an empty vessel such as that used in the so-called Texaco and Shell partial oxidation processes, or the reactor can be filled with a fixed or fluid bed catalyst.

It is preferred that the operating pressure in the gas generator be as high as is technologically feasible. Preferred pressures are those in the range of 30 to 100 atmospheres.

The effluent gas mixture leaving reactor 5 contains a substantial amount of CO, as is known in the art, e.g., amounts generally in the range of 15 to 50%, and is ordinarily at a temperature between 650.degree.C and 1000.degree.C; hence, it is directed to a waste heat boiler 6, so as to enable the recovery of its evolved heat as high pressure steam for use elsewhere in the present process, and for driving the various plant apparatus. Alternatively, part or all of this heat can be used to preheat the steam feedstock mixture introduced into the gas generator.

In the event that the feedstock has been completely gasified in reactor 5, the effluent gas is further cooled in heat exchanger 7 (for recovery of its heat) and then in final cooler 8.

Since in some gasification processes, however, the feedstock is only partially gasified, in such cases, after heat recovery in waste heat boiler 6, it is usually preferable to achieve a final cooling in a tower having coutercurrent water and liquid hydrocarbon circulation.

Since the gaseous effluent produced by gasification includes impurities, e.g., hydrogen sulfide and carbonyl sulfide, these impurities must be removed therefrom in a purification step denoted herein as sulfur removal. A number of feasible alternative means are suitable for this purpose, among these being the use of an absorption stripping system, which is preferred means. As shown in FIG. 3, a "sour" or impure gaseous effluent obtained from the gasification step shown in FIG. 2 is contacted with a conventional liquid solvent such as NaOH, KOH, diethanolamine, etc., in an absorption tower 9. Hydrogen sulfide and other sulfur bearing gases, together with part of the carbon dioxide, are absorbed selectively or non-selectively by the solvent. Acid gases are recovered from the solvent in a stripping tower 10.

The regenerated liquid solvent is then recycled to the absorption tower 9 through pump 11. The regeneration heat, if required, is supplied by a reboiler (not shown).

Alternatively, regeneration can be effected by flashing the solvent from the high pressure of absorption tower 9 to the low pressure of stripping tower 10, or by using an inert gas for stripping in stripping tower 10.

In the event the acid gases leave the stripping tower 10 at a temperature appreciably higher than the absorption temperature in absorption tower 9, it is usually preferable to cool these gases in cooler 12 and condense the liquid in condenser or heat exchanger 13 for reflux to stripping tower 10, thus obtaining the final concentrated acid gas that can be subsequently treated in a conventional sulfur plant for sulfur recovery.

In accordance with the composition of the effluent gas leaving the sulfur removal step, the gas can be split into a plurality of fractions or streams following the criteria previously described. In a preferred embodiment of this invention, the effluent gas is split into two fractions or streams of about 50% each. (The material balance, set forth in Example I below, it will be noted, showed a 45-55% split).

A first stream undergoes a methanation reaction shown in a schematic flow diagram in FIG. 4. The main purpose of this reaction is to methanate, as much as possible, the carbon oxides contained in the desulfurized gas effluent, by consuming virtually all the hydrogen present, which represents 37% in the particular case of the aforesaid material balance. In view of the fact that the residual carbon oxides content of the effluent from the first methanation step has no appreciable effect on the following step of the process, it is likely that one catalyst bed is sufficient to perform this reaction, but due to the high content of hydrogen and carbon oxides in the methanation feed, it is preferred that this goal be achieved by recycling an appreciable part of the gas effluent from the first methanation, after cooling. The amount of recycled gas can be adjusted in accordance with the gas composition from the sulfur removal step in such a way so as to maintain the maximum temperature in the methanation reactor at a level consistent with a total consumption of hydrogen in the reaction to form methane.

In one embodiment of this invention, some amount of steam may be mixed with the feed gas to the first methanation step in order to achieve some degree of shift conversion at the same time as the methanation reaction is proceeding. This mode of operation may be desirable in some cases, especially when the catalyst used does not have the proper inhibitors to avoid carbon formation due to the high CO partial pressure. In fact, it is possible to adjust this steam injection to the rate required for substantially eliminating the presence of carbon monoxide in the effluent of the fist methantation step, thus minimizing the risk of carbon formation. This process variation, although not represented in the aforesaid material balance, is within the spirit of the present invention, and should be taken into consideration for the proper split of the gas fractions after the sulfur removal step.

As shown in FIG. 4, the feed gas from the sulfur removal step is first mixed with the gas recycled from compressor 24, and the mixture is preheated in heat exchanger 25 and then fed into the methanation reactor 26. The preferred inlet temperature to the reactor is in the range of 240 to 300.degree.C, while the preferred outlet temperature is in the range of 450 to 700.degree.C.

The gas leaving the reactor 26 is first cooled in heat exchanger 27 so as to recover the high temperature heat either in generating steam, or in any other way compatible with the other sections of the plant. The gas is then cooled in heat exchanger 25 to preheat the reactor feed and finally is further cooled in heat exchanger 28 for heat recovery to a lower level and in cooler 29 to about ambient temperature. The water produced in the methanation reaction is condensed and separated before proceeding to the following stage of the process.

The second desulfurized effluent gas fraction is subjected to shift conversion, so as to convert CO contained in said effluent gas fraction to CO.sub.2. This reaction requires an appreciable amount of steam, whereas the gas effluent from the sulfur removal step is usually at a temperature close to ambient temperature and contains very little or no steam. Thus, for optimum heat recovery around the shift reactor, it is preferred to use a humidifier-dehumidifier system to convey the heat to the shift reactor, in the form of steam, flowing from downstream to upstream.

As shown in FIG. 5, the cold, second desulfurized effluent gas fraction is first contacted countercurrently in the humidifier tower 30 with hot water coming from the dehumidifier tower 31 and then is mixed with an additional amount of steam, and then preheated in heat exchanger 32 to a temperature of between 350 and 400.degree.C.

The gas then enters the first shift converter (reactor 33) which preferably uses a conventional high temperature shift catalyst to expedite or otherwise render more efficient the shift conversion operation. In one embodiment of this invention, shift conversion can be conducted with two or more catalyst beds within reactor 33, with a quench or heat exchanger in between the beds to achieve a greater part of the conversion duty in said reactor 33. The gas leaving the reactor 33 is cooled in exchanger 34 or quenched to bring down the temperature to a lower level in a second shift converter 35. The catalyst used in the latter can be any conventional low temperature shift catalyst that is operable within the range of 200.degree. to 250.degree.C. However, in those instances where the gasification pressure is very high, it would not be possible to use a low temperature shift reactor because of the risk of condensation over such a conventional catalyst, unless a new catalyst were developed to withstand these operating conditions.

The carbon monoxide content of the gas leaving the low temperature shift reactor 35 is in the range of about 0.2 to about 0.8% dry basis, whereas the gas leaving the high temperature shift converter has a residual CO content of between about 2 and about 5%.

The gas leaving the final shift converter can be cooled, if necessary, in heat exchanger 36 before entering the dehumidifier tower 31. In this tower, the cold water flow coming from humidifier tower 30 is preheated countercurrent to the hot gas, and the heat thus gained by the water is later released in humidifier tower 30 to the feed gas.

The gas leaving the dehumidifier tower 31 is further cooled in heat exchanger 37 before being conveyed to the CO.sub.2 removal step.

The heat recovery in the various heat exchangers 32, 34, 36, 37 can preferably be optimized in each particular case for best efficiency by combining the streams either in the shift conversion step itself, or by combining such streams with gas or liquid streams from other process steps.

The effluent gas resulting from the shift conversion of the second gas fraction is subjected to a CO.sub.2 removal step in order to eliminate virtually all the CO.sub.2 present and thus obtain a hydrogen-rich stream with which to methanate the residual carbon oxides in the first effluent gas stream resulting from first methanation upon combination of the hydrogen-rich stream with said first methanation effluent stream. The hydrogen-rich stream is typically, e.g., a hydrogen-nitrogen-methane mixture containing small amounts of CO and CO.sub.2, much of the nitrogen having been introduced during gasification.

While there are several conventional means that can be used for the removal of CO.sub.2 from this gas stream, such as by scrubbing with appropriate solvents such as various amines, potassium carbonate, methanol, various other organic solvents, etc., the means most preferred for the present process would be the one most suited to:

(a) take maximum advantage of the high partial pressure of CO.sub.2 available; (b) utilize minimum amounts of heat or other forms of energy for regeneration of the solution; (c) allow, if necessary, a small amount of CO.sub.2 leakage in the gas leaving the scrubber, in view of the fact that it will be further methanated downstream in the second methanation step; and to (d) avoid use of any chemical harmful to the methanation catalyst downstream.

Generally, those carbon dioxide removal processes using absorption and regeneration above ambient temperature operate along the lines indicated in the schematic flow diagram illustrated in FIG. 6.

As shown in FIG. 6, the gas leaving the shift conversion step is scrubbed countercurrent to the regenerated solution in the scrubber 40, and the purified gas leaving the top of the scrubber goes directly to the second methanation step after being mixed with the effluent gas fraction from the first methanation step.

The rich solvent leaving the bottom of scrubber 40 is flashed to near atmospheric pressure into the regeneration tower 41, preferably after having been preheated by the hot regenerated solvent extracted from the bottom of regeneration tower 41 and pumped by pump 42.

Reboiling heat can be supplied at the bottom of regeneration tower 41 through heat exchanger 43. The gas leaving the top of the regenerator tower 41 is cooled in cooler 44, and the condensed solvent is separated in separator 45 and recycled to the system through pump 46, whereas the concentrated CO.sub.2 stream is evacuated from separator 45 to the atmosphere or for any downstream use desired.

In the second methanation step of the present process, the effluent stream exiting from the CO.sub.2 removal step is combined with the effluent stream exiting from the first methanation step. Thus, the combined gas stream to be subjected to methanation in the second methanation step of the present process has approximately the stoichiometric amount of hydrogen required to methanate all the CO and CO.sub.2 contained therein. It should be recognized, however, that a very slight excess of hydrogen, above the stoichiometric ratio, may be desirable in the combined gas feed to the second methanation, in order to achieve a complete methanation of the carbon oxides. On the other hand, a deficiency of hydrogen in the second methanation feed will lead to a small amount of CO.sub.2 and only traces of CO in the second methanation effluent. Either deviation, whether above or below the stoichiometric ratio of hydrogen to carbon oxides, may be justified under certain economic conditions, at the option of the process operator, and although such deviations have not been set forth in the subsequent typical material balance, they are nevertheless within the spirit and scope of the present invention, and should be taken into consideration for the proper split of the gas fractions after the sulfur removal step.

It is preferred that this second step stage methanation be effected in a manner slightly more elaborate than in the first methanation, since the complete elimination of carbon oxides in the product gas is intended. To achieve this goal, it is preferred that the second methanation step be effected in two stages, as illustrated by the flow diagram contained in FIG. 7, wherein the first stage is conducted in reactor 50 at a very high temperature, with a correspondingly high leakage of carbon oxides; whereas, in the second stage, conducted in reactor 51, the reaction proceeds at a very moderate temperature so as to result in a complete conversion to methane.

Since the content of the carbon oxides in the gas mixture is quite appreciable, it is important that there be a gas recycle through compressor 52 in order to maintain the reaction temperature in reactor 50 within reasonable limits.

The gas mixture from the first methanation and CO.sub.2 removal steps is first mixed with the recycled gas, then preheated in feed product exchanger 53 by the effluent gas from the second reactor 51, and then fed into the first stage of reactor 50 at a temperature of from about 240.degree. to about 300.degree.C. The outlet temperature from reactor 50 is kept between about 450.degree. and about 700.degree.C. The gas effluent from the first reactor 50 is first cooled in heat exchanger 54, where high temperature heat can be recovered either to produce high pressure steam or to supply heat to other parts or stages of the process.

If desired, the inlet temperature to the second bed can be adjusted by injecting some cold fresh gas by-passing the first reactor 50. This procedure has the advantage of requiring less recycle gas to achieve a given temperature rise in the first reactor. The inlet temperature to the second reactor 53 is preferably adjusted between about 240.degree. and about 300.degree.C and the resulting outlet temperature is preferably in the range of about 270.degree. to about 400.degree.C.

The gas effluent from reactor 51 is first cooled in the feed product exchanger 53 and in another heat exchanger 55 and then in the final cooler 56. Part of the gas produced is recycled through compressor 52 and the rest is conveyed to the cryogenic stage or step of the present process.

It is preferred that cryogenic separation take place at a very low temperature in order to obtain a liquid methane fraction that is separable from the nitrogen gas stream. Therefore, it is important to separate the traces of water vapor and carbon dioxide that are still present in the gas. This can be done in several ways. For example, a switch exchanger system can be employed at the inlet of the cryogenic separation system. When the deposition of ice and dry ice becomes important in the first feed exchanger to the point where it becomes inoperable, the feed gas is switched to the other parallel exchanger, and the fouled exchanger can then be rendered utilizable during the time the other exchanger is in service. Alternatively, the gas can be treated with molecular sieves which would act as absorbents to remove the traces of water and CO.sub.2 from the gas. As an additional alternative, a chemical treatment can also be employed, whereby the final traces of H.sub.2 O and CO.sub.2 can be removed by scrubbing with a suitable solvent such as ethylene glycol.

The drying of the feed gas to the present cryogenic separation step is not shown in the overall schematic flow diagram illustrated in FIG. 8.

As utilized herein, the basic concept underlying the present cryogenic separation is to obtain a methane-rich mixture, while leaving a minimum of methane in the nitrogen stream to be vented to the atmosphere. It should be realized that the particular scheme to be followed in this cryogenic separation can vary appreciably depending on the following variables:

a. the nitrogen content of the feed coming from the second methanation stage;

b. the nitrogen content desired in the final methane-rich gas or substitute natural gas;

c. the efficiencies required in terms of the methane content of the nitrogen stream vented to the atmosphere; and

d. the available pressure of the feed gas from the second methanation stage.

In any event, it is believed preferable to attain a methane concentration in the waste nitrogen stream that is vented to the atmosphere, of between about 0.4 and about 1.0% by volume, and a nitrogen concentration in the methane-rich product gas of between about 1.0 and about 8.0% by volume.

In FIG. 8, there is shown a typical cryogenic separation contemplated for application in the present process. In this figure, the feed gas from the second methanation step is fed into a heat exchanger at the full operating pressure of the gas production train. The feed gas is cooled by the waste nitrogen stream to the atmosphere and also by the methane-rich gas. Part of the methane contained in the feed gas is condensed in heat exchanger 60. The feed gas and the condensed liquid, together or separately, are then held at a pressure slightly below the partial pressure of the methane in the feed gas, whereby the resultant gas at such pressure in tower 61 undergoes a separation such that there results a liquid methane stream in the bottom and a gas nitrogen stream in the top of tower 61. The latter is again "throttled" (through valve or expander) into a flash drum 62, wherein the extra liquid methane that has condensed is collected and is pumped through pump 63 to the same pressure as in tower 61, and mixed with the liquid extracted therefrom. The pressure in flash drum 62 is usually that required to meet the requirements of the whole system, depending upon the nitrogen content of the feed gas from the second methanation. Whenever possible, the waste nitrogen stream is recovered at a pressure above atmospheric and used to supply power wherever needed in the process, through a gas expander.

The cold nitrogen stream leaving the top of tower 62 is first reheated in a reflux condenser at the top of tower 61 and then in a feed-product exchanger not shown.

Any excess hydrogen or CO present in the gas after the second methanation step, as mentioned above, will be evacuated with the waste nitrogen stream from the cryogenic separation.

In Example I below, there is set forth a typical example of the present invention, illustrated by a material balance, obtained by carrying out the present process, as illustrated in FIG. 1. The block diagram of FIG. 1 represents the main process steps, using a gas oil feedstock and air in the gas generator. The material balance of Example I specifically applies to the process streams as numbered on FIG. 1 on the block diagram.

__________________________________________________________________________ EXAMPLE I MATERIAL BALANCE PRODUCTION OF SNG BY GASIFICATION OF GAS OIL C15.86 H 29.68 S 0.0344 BASED ON 100 MOLES OF __________________________________________________________________________ FEED MOLES GAS OIL STEAM AIR GAS EFFL. TOTAL GAS GAS FRACTION GAS FRACTION TO TO TO FROM AFTER TO HTS* TO FIRST GENERATOR GENERATOR GENERATOR GENERATOR SULFUR AND LTS** METHANATION REMOVAL 1 2 3 4 5 6 7 __________________________________________________________________________ GAS OIL 100 -- -- -- -- -- -- H.sub.2 -- -- -- 1084.40 1084.40 490.95 593.45 CO -- -- -- 667.54 667.54 302.22 365.32 CO.sub.2 -- -- -- 245.70 122.85 55.62 67.23 CH.sub.4 -- -- -- 672.76 672.76 304.58 368.18 H.sub.2 S -- -- -- 3.44 -- -- -- O.sub.2 -- -- 104.79 -- -- -- -- N.sub.2 + A -- -- 395.22 395.22 395.22 178.93 216.29 H.sub.2 O -- 1376.43 -- 427.07 -TOTAL 100.00 -- 500.01 3069.06 2942.77 1332.30 1610.47 MOLES DRY MOLES GAS FRACTION GAS FRACTION GAS FRACTION COMBINED TOTAL GAS AFTER 1st AFTER AFTER CO.sub.2 GAS TO 2nd TO CRYOGENIC NITROGEN FINAL METHANATION *HTS LTS** REMOVAL METHANATION SEPARATION STREAM SNG 8 9 10 11 12 13 14 __________________________________________________________________________ GAS OIL -- -- -- -- -- -- -- H.sub.2 -- 787.73 787.73 787.73 -- -- -- CO 177.24 5.43 5.43 182.68 -- -- -- CO.sub.2 59.93 352.40 -- 59.93 -- -- -- CH.sub.4 563.56 304.58 304.58 868.14 1110.75 1.5 1103.25 H.sub.2 S -- -- -- -- -- -- -- O.sub.2 -- -- -- -- -- -- -- N.sub.2 +A 216.29 178.93 178.93 395.22 395.22 360.92 34.30 H.sub. 2 O -- -- -- TOTAL 1017.02 1629.07 1276.67 2293.7 1505.97 362.42 1148.55 MOLES DRY __________________________________________________________________________ *HTS= high temperature shift conversion **LTS = low temperature shift conversion

It should be noted that the split between the two streams, after sulfur removal, is such that the combined gas to the second methanation step contains the stoichiometric ratio of hydrogen to carbon oxides.

While particular embodiments of the present invention have been described, it will be understood, of course, that this invention is not limited thereto since many modifications may be made, and it is, therefore, contemplated to cover by the appended claims any and all such modifications as may fall within the true spirit and scope of this invention.

Claims

What is claimed is:

1. A process for producing a high Btu, methane-rich gas from a suitable carbon- or hydrocarbon-containing feedstock, comprising (a) generating a gaseous effluent having a substantial CO content from said feedstock and purifying said gaseous effluent to remove the impurities therefrom comprising the residual carbon and sulfur impurities; (b) dividing the resulting gaseous effluent into a plurality of fractions; (c) subjecting a first effluent fraction to methanation; (d) subjecting a second effluent fraction to the successive steps of shift conversion and CO.sub.2 removal; (e) mixing the resulting effluents from steps (c) and (d) and subjecting the resulting mixture to methanation; and (f) separating the excess nitrogen present from the methane-rich product.

2. A process according to claim 1, wherein the feedstock is a fossil fuel.

3. A process according to claim 1, wherein the generation of said gaseous effluent is conducted in the presence of an oxidant selected from the group consisting of air and oxygen enriched air and the resulting nitrogen contained in the methane-rich final gas stream is separated therefrom by cryogenic means.

4. A process according to claim 1, wherein the feedstock is one heavier than naphtha and the gaseous effluent from said feedstock is generated at a temperature between about 650.degree.C and 1000.degree.C under adiabatic conditions.

5. A process according to claim 4, wherein the pressure is between 30 and 100 atmospheres.

6. A process according to claim 1 wherein the gas generation step is conducted under adiabatic conditions in the presence of a catalyst.

7. A process according to claim 1, wherein at least substantially all the hydrogen in the first effluent fraction is reacted with the carbon oxides present therein.

8. A process according to claim 2, wherein a humidifier-dehumidifier system is used to supply at least a part of the steam required for shift conversion.

9. A process according to claim 1, wherein the mixture of effluents from steps (c) and (d) contains a ratio of hydrogen to carbon oxides substantially stoichiometric to that required for substantially complete methanation of said carbon oxides.