CA1210242A - Production of synthetic natural gas from coal gasification liquid by-products - Google Patents

Abstract

ABSTRACT OF THE DISCLOSURE

In coal gasification processes for the production of snythetic natural gas by the reaction of coal with steam and oxygen under pressure to form a gasifier synthesis gas and a liquid hydrocarbon by-product, the liquid hydrocarbon by-product is treated for solids and metal removal and is then passed to a cata-lytic partial oxidation zone containing a monolithic platinum-palladium catalyst. The hydrocarbon by-product liquids are con-verted to secondary synthesis gas by being reacted with steam and oxygen. Optionally, the effluent from the catalytic partial oxidation zone may be passed through a second, steam reforming catalyst to react residual hydrocarbons with water to produce hydrogen and carbon oxides. The gasifier and secondary synthesis gases may be treated to remove acid gases therefrom and then methanated to provide a product synthetic natural gas.

Description

4~
., BAC~GROUND OF THE I~IVENTION

The present invention relates to a process for the production 'of a svnthetic natural gas, more particularl~, to the production of synthetic natural gas from a secondarv synthesis gas, by which ,term is meant a synthesis gas produced from the coal derivedliquid hydrocarbon by-product resulting from the g~sification of coal to produce gasifier synthetic natural gas. The secondary synthesis ,gas is methanated to provide additional synthetic natural gas which may be blended with the gasifier synthetic natural gas.
Coal gasification technology is well known and has been in ,commercial use in South Africa since about 1954 and was commonly used in the United.States prior to 1950 to make town gas. The ;most commonly employed gasifier process is that developed by Lurgi ~ohle und Mineraloeltechni~ GmbH, Frankfurt (~lain), Federal ~epublic of Germany. The Lurgi process utili~es a fixed bed gasifier in which coal of a selected particle size is fed into the top of the gasifier countercurrently to a stream of steam and , oxygen fed from the bottom of the gasifier. A synthesis gas (herein and in the claims called gasifier synthesis gas) and a ' hydrocarbon liquid by-product are produced from the coal and with-, drawn from,thetop of the gasifier. Solid ash residue is withdrawn thro~gh a rotating grate at the bottom of the gasifier. Up to ~ about one fourth of the coal fed to the process will emerge as thc ! liquid hydrocarbon by-product, rather than as the desired gasifier , synthesis gas. Thus, a gasification plant using 8 million tons l of coal per year may produce as much as about 2 million tons of ' llquid hydrocarbon by-product.
Such liquid hydrocarbon by-product essentially comprise three , major fractions classified as oil, tar and phenolics. The oil ~Zl(~2 fraction has ~ boiling range ofabout 200-600~ (93-316C) and requires treatmcnt if it is to be cmployed as a petroleum product substitute. The tar contains substantial quantities of oxyger.
and nitrogen and about .01 percent ash. The phenolics are some-~hat similar to cresols and have a boiling range of about 29~-~noo.-(1~3-~05C). Generally these by-products are not particularly valuable and do not command a high price even in those areas where a market exists for them. The liquid hydrocarbon by-product is also carcinogenic and toxic. Disposition of the by-product l~hen no market exists for it presents significant environmental and economic problems.
The present invention enables the conversion of such liquid hydrocarbon by-product into additional synthesis gas (herein and in the claims called secondary synthesis gas). This "secondary synthesis gas" (sometimes herein abbreviated to "secondary SG") is to be distinguished from the "gasifier synthesis gas" (someLimss herein abbreviated to "gasifier SG") which is obtained in the coal gasification step. Methanation of the gasifier SG and secondary SG is carried out to provide product synthetic natural gas (some-2v times herein abbreviated to "SNG"). As explained below the con-version of the liquid hydrocarbon by-product to secondary SG is carried out by a catalytic partial oxidation process in which steam reforming and hydrocrac~ing reactions are believed to also take place and to provide an efficient and economical means of 2i converting the liquid hydrocarbon by-product.
Steam reforming is a well known method for treating hydro-carbons to generate hydrogen therefrom. It is usually carried out by supplying hea~ to a mixture of steam and a hydrocarbon feed while contacting the mixture with a suitable catalyst usually z nic~el. Steam reforming is gcnerally limitcd to paraffinic naphtha and lightcr fceds which have been de-sulfurizcd and .reatDd to remove ~iLro~cn compo~lnds, because of difficulties in attempting to steam reform heavier hydrocarbons and the poisoning of steam reforming catalysts by sulfur and nitro~en compounds.
Another known method of obtaining hydrogen from a hydrocarbon feed is partial oxidation, in which the hydrocarbon feed is intro-duced into an oxidation zone maintained in a fuel rich mode so that only a portion of the feed is oxidized.
It is known that steam may also be injected into the partial o~idation reactor ~ressel to react with the feed and with products of the partial oxidation reaction. The process is not catalytic and requires high ~emperatures to carry the reactions to comple-tion, resulting in a relatively high oxygen consumption. On the other hand; the partial oxidation process has the advantage that it is able to readily handle hydrocarbon liquids heavier than paraffinic naphthas and can even utilize coal as the source of the hydrocarbon feed.
Catalytic autothermal reforming of hydrocarbon liquids is

2~ also known in the art, as evidenced by a paper Catalytic Autothermal Reforming of Hydrocarbon Liquids by ~laria Flytzani-Stephanopoulos and Gerald E. Voecks, presented at the American ; Institute of ChemicalEngineers' 90th National Meeting, Hous-or, Texas, April 5-9, 1981. Autothermal reforming is defined thcre-in as the utilization of catalytic parital oxidation in the presence of added steam, which is said to increase the hydrogen yield becaus^ of simultaneous (with the catalytic partial

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oxidation) stcam reforming being attained. The paper discloses ,utilization of a particulate bed of three different nicXel catalysts into ~hic:~ steam, air and a hydrocarbon fuel supply comprising a ~lo. 2 fuel oil are iniected. The resulting product ,gases contain hydrogen and carbon o~ides.
In Brennstoff-Chemie 46, ~'o. 4, p. 23 (1965), a German publication, Von P. Schmulder describes a Badische Anilin and Soda ~Fabrik (B~SF) process for autothermal reforming of gasoline. The 'process utilizes a first, pelletized, i.e., particulate, platinum catalyst zone followed by a second, pelletized nickel catalyst zone. A portion of the product gas is recycled to the process.
Disclosure of the utilization ofa noble metal catalyzed mono-lith to carry out catalytic partial oxidation to convert more than half of the hydrocarbon feed stock upstream of a steam re-l~ forming zone is disclosed in an abstract entitled Evaluation of ¦' Steam Reformin~ Catalvst for use in the_Auto-Thermal Reforming li of Hydrocarbon Feed Stocks by R. M. Yarrington, I. R. Feins, and ~ S. ~Iwang (~ational Fuel Cell Seminar, July 14-16, 19~0, San ¦I Diego). The abs-~ract noted the unique ability of rhodium to I steam reform light olefins with little coke formation and noted I that results were obtained for a series of platinum-rhodium ¦ catalysts with various ratios of platinum to total metal in which ¦¦ the total metal content was held constant.
¦1 U. S. Patent 4,054,407, assigned to the assignee of this ¦¦ application, discloses two-stage catalytic o~idation using platinum group metal catalytic components dispersed on a mono-lithic body. At least the stoichiometric amount of air is ¦ supplied over the two stages and steam is not employed.

-4-U.S. Patent 3,481,722, assigned to the assignee of this application, discloses a two stage process for steam reforming normally liquid hydrocarbons using a platinum group metal catalyst in the first stage. Steam and hydro-gen, the latter of which may be obtained by partially cracking the hydrocarbon feed, are combined with the feed to the process.

SUMMARY OF THE INVENTION
.
In accordance with the present invention there is provided in a coal gasification process in which coal is reacted with steam and oxygen to produce (i) a gasifier synthesis gas which is methanated to produce a synthetic natural gas, and (ii) a liquid hydrocarbon by-product, the improvement of preparing a secondary synthesis gas from the liquid hydrocarbon by-product, which secondary ; synthesis gas may be methanated to form additional synthetic i! natural gas, by the following steps: preheating an inlet stream comprising the liquid hydrocarbon by-product, H2O

and oxygen to a preheat temperature which is preferably at least 800F (427C) but in any case sufficiently high to initiate catalytic oxidation of the hydrocarbon by-product as defined below, but less than about 1200F (649C)~
introducing the preheated inlet stream into a first catalyst zone comprising a monolithic body having a plurality of gas flow passages extending therethrough and comprising palladium and platinum catalytic components and op-tionally rhodium catalytic component dispersed therein, the amounts of hydrocarbon by-product, H2O and oxygen introduced into ~2~ 2 the first catalyst zone being controlled to maintain an H2O to C ratio of from about 0.5 to 5 and an 2 to C ratio from about 0.15 to 0.4 in the inlet stream;
contacting the preheated inlet stream within the first catalyst zone with the aforesaid catalytic component to initiate and sustain therein catalytic oxidation of a quantity, substantially less than all, of the hydrocarbon by-product sufficient to attain a temperature within the `' first catalyst zone at least high enough to crac~ substan-tially all unoxidized C4 or heavier hydrocarbons in the by-product tD Cl to C4 hydrocarbons, the temperature of at least a portion of said monolithic body being at least 250F (139C~ higher than the ignition temperature of said inlet stream, but not more than about 2000F (1093C), whereby to produce a first catalyst zone effluent comprising ~A methane, hydrogen, carbon monoxide, carbon dioxide and . .
H2O; passing the effluent to a treatment zone for the removal of carbon dioxide and water therefrom and with-drawing the treated first catalyst zone effluent as second-ary synthesis gas product; and methanating the gasifier synthesis gas and the secondary synthesis gas to provide synthetic natural gas therefrom.
In one aspect of the present invention, there is included the step of passing the first catalyst zone eff--.;i luent, while still at an elevated temperature, from the first catalyst zone to a second catalyst zone containing a steam reforming catalyst, preferably comprising platinum and rhodium catalytic components, and contacting the first "

~2~ 2 zone effluent in the second catalyst zone with the steam reforming catalyst to react hydrocarbons therein ~,Jith H2O to produce hydrogen and carbon oxides therefrom, and then passing the effluent of the second catalys'c zone as the effluent to the aforesaid treatment zone for the re-moval of carbon dioxide and water therefrom.
In preferred aspects of the inven~ion, at least about 50% by weight of the hydrcarbon by-product is con-verted to Cl hydrocarbons in the first catalyst zone and/
or a total of at least about 98% by weight of hydrocarbon by-product is converted to Cl hydrocarbons in the first and second catalyst zones.
In another aspect of the invention, the effluent of the first catalyst zone may be substantially entirely depleted of oxygen and the secondary synthesis gas may be combined with the gasifier synthesis gas to provide a combined product synthesis gas.
In another aspect of the invention, the first catalyst zone is maintained at a temperature of from about 1400F
to 2000F (760C to 1093C) and the first zone effluent is introduced into the second catalyst zone at substantially the same temperature. A volumetric hourly rate of at least 100,000 volumes of throughput per volume of catalyst may be maintained in the first catalyst zone and a volumetric hourly rate of from about 2,000 to 20,000 volumes of throughput per volume of catalyst may be maintained in the second catalyst zone, and the process of making the second-ary SG may be carried out at a pressure of from about 50 psig to 1500 psig.

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In one aspect of the invention, the liquid h~dro-carbon by-product is treated to remove ash and metals, if any, and the heaviest portion of the tars therein, prior to being passed to the first catalyst zone.

BRIEF DESCRIPTION OF THE DRAWINGS
Figure 1 is a schematic elevation view in cross section of a laboratory or p:ilot plant size embodiment of an autothermal reformer apparatus utilizable in accordance with the present invention; and Figure 2 is a schematic flow sheet diagram of a coal gasification plant, including an autothermal reforming section for conve~ting liquid hydrocarbon by-product from the coal gasifier to secondary synthetic natural gas.

DESCRIPTION OF THE PREFERRED EMBODIMENTS
In a preferred embodiment of the present invention, a coal gasifier plant includes a section for making a secondary SG~ which section includes an autothermal reformer.
By incorporating the secondary SG section into the coal gasificatlon plant, a greater quantity of product SG can be obtained with the same size coal gasifier plant or a smaller plant can be utilized to produce a given total product synthetic natural gas (SNG) while consuming an environmentally damaging substance.
In a preferred embodiment of the present invention, at least a first catalst zone is provided for carrying out catalytic partial oxidation, an exothermic reaction, and a second, catalytic zone is optionally provided for carrying ..~

out steam reforming, an endothermic reaction. Steam reforming, as well as hydrocracking ofheavier hydrocarbons, , also appears to take place in the first catalyst zone so that under certain conditions a second catalyst zone specifically for steam reforming is not require. Such steam reforming as takes place in the the first catalyst zone absorbs some of the heat generated by the partial oxidation step and tends to moderate the operating temperature attained. The net reaction in the first catalyst zone is however exothermic. The exothermic, first catalyst zone comprises a monolithic catalyst carrier on which a platinum group metal catalyst is dispersed. Such catalyst can effectively catalyze the partial oxidation and steam reforming of the liquid hydrocarbon by-product, resulting in hydrogen being formed. As compared to a non-catalytic combustion process, such as conventional non-catalytic i partial oxidation, catalytic partial oxidation enables the utilization of lesser amounts of oxygen and lower temper-ature levels to both oxidize andhydrocrack the liquid hydrocarbon by-product to lighter hydrocarbon fractions.
Nonetheless, the temperature of the reactant mass is sufficiently elevated for the optional subsequent steam reforming step, in those cases where steam reforming is required. At the temperatures maintained in the catalytic oxidation zone, and in the presence of the product hydrogen and catalyst utilized in the first zone, hydrocracking of the unoxidized C2 and heavier hydrocarbons takes place to .~

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form primarily Cl hydrocarbons, with minor amounts of C2 and C3 compounds. The effluent gas from the first catalyst . .
zone contains primarily H2, H2O, CH4, CO and CO2 and, , depending upon the sulfur content of the liquid hydrocarbon by-product, H2S and COS.
The endothermic, second catalyst zone may contain any suitable steam reforming platinum group metal catalyst.
Usually, the steam reforming catalyst will be utilized in the form of a particulate bed comprised of spheres, extrud-~ 10 ates, granules, configured packing material, e.g., rings, saddles or the like, ox any suitable shape. Obviously, a combination of different types of particulate materials may be utilized as the steam reforming catalyst. Further, a monolithic catalyst carrier may also be used in the second catalyst zone, as is used in the first catalyst zone.
The process of the present invention provides a simpler and less expensive means of converting the liquid hydrocarbon by-product of a coal gasification process, e.g., the Lurgi coal gasification process than the conventional partial oxidation or steam reforming processes. The combination of features provided by the present invention provides a highly efficient and flexible method of effectuating the conversion. The low pressure drop and high volumetric rate throughput of a monolithic body platinum group metal catalyst i provides a reduced size and volume of catalyst. The use of platinum group metals as the catalyst metal re~uires a very low catalytic metal loading as compared to use of base metal catalyst. This provides good overall economies ~, -- 1 0 in reduced equipment size and enhanced throughput rates despite the much higher unit weight cost of platinum group metals as compared to base metals. The combination of the monolithic platinum group metal partial oxidation catalyst ¦ with a platinum group metal steam reforming catalyst enables operations at relatively very low 2 to C ratios without carbon deposition fouling the catalyst, which is important in attaining enhanced CH4 production in the process. Use of platinum group metal catalysts also enhances resistance of the catalyst to poisoning by sulfur compounds and enhances the ability of the catalyst to treat the aromatics in the liquid hydrocarbon by-product.

The Monolithic Partlal Oxidation Catalyst The partial oxidation catalyst is provided on monolithic carrier, that is, a carrier of the type comprising one or more monolithic bodies having a plurality of finely divided gas flow passages extending therethough. Such monolithic carrier members are often referred to as "honeycomb" type carriers and are well known in the art. A preferred from of such carrier is made of a refractory, substantially inert rigid material which is capable of maintaining its shape and a sufficient degree of mechanical strength at high temperatures, for example, up to about 3,272F (1,800C). Typically, a material is selected for the support which exhibits a low thermal coefficient of expansion, good thermal shock resistance i and, though not always, low thermal conductivity. Two general types of material of construction for such carriers are known.

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One is a ceramic-like porous material comprised of one or more metal oxides, for example, alumina, alumina-silica, alumina-silica-titania, mullite~ cordierite, zirconia, zirconia-spinel, zirconia-mullite, silicon carbide, etc.
A particularly preferred and commercially available material of construction is cordierite, which is an alumina-magnesia-silica material and well suited for operations - below about 2,000F (1,093C). Honeycomb monolithic supports are commercially available in various sizes and configurations. Typically, the monolithic carrier would comprise, e.g., a cordierite member of generally cylindrical configuration (either round or oval in cross section) and having a plurality of parallel gas flow passages of regular polygonal cross sectional extending therethrough. The gas flow passages are typically sized to provide from about 50 to 1,200, preferrably, 200-600 gas flow channels per square inch of face area.
The second major type of preferred material of construction for the carrier is a heat- and oxidation-resistant metal, such as a stainless steel or the like.
Monolithic supports are typically made from such materials by placing a flat and a corrugated metal sheet one over the other and rolling the stacked sheets into a tubular configuration about an axis parallel to the corrugations, :d to provide a cylindrical-shaped body having a plurality of fine, parallel gas flow passages extending therethrough.
The sheets and corrugations are sized to provide the desired number of gas flow passages, which may range, typically, ~2~ Z
from about 200 to 1,200 per square inch of end face area of the tubular roll.
Although the ceramic-like metal oxide materials such as cordierite are somewhat porous and rough-textured, they nonetheless have a relatively low surface area with respect to catalyst support re~uirements and, of course, a stainless steel or other metal support is essentially smooth. Accord-ingly, a suitable high surface area refractory metal oxide support layer is deposited on the carrier to serve as a support upon which finely dispersed catalytic metal may be distended. As is known in the art, generally, oxides of one or more of the metals of Groups II, III and IV of the Periodic Table of Elements having atomic numbers not greater than 40 are satisfactory as the support layer. Preferred high surface area support coatings are alumina, berylia, zirconia, baria-alumina, magnesia, silica and combinations of two or more of the foregoing.
The most preferred support coating is alumina, most preferably a stabilized, high-surface area transition alumina.
As used herein and in the claims, "transition alumina" includes gamma, chi, eta, kappa, theta and delta froms and mixtures thereof. An alumina comprising or predominating in gamma alumina is the most preferred support layer. It is known that certain additives such as, e.g., one or more rare earth metal oxides and/or alkaline earth metal oxides may be included in the transition alumina (usually in amounts comprising from 2 to 10 weight percent of the stabilized coating) to stabilize it against the generally undesirable high temperature phase transition to alpha alumina, which is of a relatively low ~,~

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surface area. For example, oxides of one or more of lanthanum, cerium, praseodymium, calcium, barium, strontium and magnesium may be used as a stabilizer. The specific combination of oxides of lanthanum and barium is a preferred stabilizer for transition alumina.
As used herein and in the claims, the term "platinum group metal" means platinum, palladium, rhodium, iridium, osmium and ruthenium. The platinum group metal used may optionally be supplemented with one or more base metals, ; 10 particularly base metals of Group VII and metals of Groups VB, VIB and VIIB of the Periodic Table of Elements. Preferably, one or more of chronium, copper, vanadium, cobalt, nickel and iron may be employed.
Desirable catalysts for partial oxidation should have the following properties: They shouId be able to operate effectively under conditions varying from oxidizing at the inlet to reducing at the exit; they should operate effect-ively and without significant temperature degradation over a temperature range of about 800F to about 2400F (427C to 1315C); they should operate effectively in the presence of carbon monoxide, olefins and sulfur compounds; they should provide for low levels of coking such as by preferentially catalyzing the reaction of carbon with H2O to form carbon monoxide and hydrogen thereby permittlng only a low level of carbon on the catalyst surface; they must be able to resist poisoning from such common poisons as sulfur and halogen compounds; further, all of these requirements must be satisfied simultaneously. For example, in some otherwise ~ 14 -~Z~24Z
suitable catalysts, carbon monoxide may be retained by the catalyst metal at low temperatures thereby decreasing or modifying its activity. The combined platinum and palladium is a highly efficient oxidation catalyst for the purposes j of the present invention. Generally, the catalyst activity of platinum-palladium combination catalysts is not simply an arithmetic combination of their respective catalytic activities; the disclosed range of proportions of platinum and palladium have been found to posses the previously described desirable properties and, in particular, provide efficient and effective catalytic activity in treating a rather wide range of hydrocarbonaceous, particularly hydro-carbon, feeds with good resistance to high -temperature operation and catalyst poisons.
` The following data compare the effectiveness of pall-adium, rhodium and platinum, respectively, for the oxidation of methane and further compares the efficacy of, respectively, palladium-platinum, palladium-rhodium and platinum-rhodium combined catalyst for oxidation of methane.
The catalyst of Table I-A comprise a lanthia-chromia-alumina frit impregnated with the platinum group metals by techniques as described aboveO The frit has the following composition:
Component Weight Percent 23 3.8 Cr203 1.8 The lanthia-chromia stabilized alumina is then impreg-nated with the platinum group metal and calcined in air for -~ -r 1 5 four hours at 230F and for an additional four hours at 1600F. Three catalysts of different platinum metal loadings were prepared as follows:

Weight Percent Sample No.PdPt Rh Total PGM
~063U-1 3.425.95 - 9.37 4063R-1 4.58- 4.52 9.10 4063V-1 -5.62 3.14 8.76 The resultant platinum group metal (PGM) impregnated alumina frit was deposited on alumina beads and the thus-coated beads were placed in a shallow bed and tested by passing a 1% (volume) methane 99% (volume) air feed at about atmospheric pressure through the catalyst. An electric heater was used to cyclically heat the test gas stream fed to the catalyst, and conversion results at the indicated temperatures were obtained on both the heating and cooling phases of each heat cycle.
The results are shown in the following Table I-A.
TABLE I-A

PGM Weight Percent of Original Sample (Mole Ignition Content Converted at Indicated No. Ratio) Temp. F Temperature (F) 4063U-1Pd,Pt(l.l) 610 - 3 10 26 60 80 4063R-1Pd,Rh(l-l) 710 - - 2 5 9 12 4063V-1Pt,Rh(l:l) 730 - - 1 1 3 5 :' .
These data demonstrate the ability of platinum-palladium to promote catalytic oxidation over a wide range of temperatures.
Rhodium may optionally be included with the platinum and palladium. Under certain conditions, rhodium is an ~` ~2~

effective oxidation as well as a steam re~orming catalyst, particularly for light olefins. The combined platinum group metal catalysts of the invention also have a signif-icant advantage in the ability to catalyze the autothermal reactions at quite low ratios of H2O to carbon (atoms of carbon in the ~eed) and oxygen to carbon, without signif-icant carbon deposition on the catalyst. This important feature provides flexibility in selecting H2O to C and 2 - to C ratios in the inlet streams to be processed.
The platinum group metals employed in the catalysts ,of the present invention may be present in the catalyst ;.~ composition in any suitable form, such as the elemental metals, as alloys or intermetallic compounds with the other - platinum group metal or metals present, or as compounds such as an oxide of the platinum group metal. As used in the . claims, the terms palladium, platinum and/or rhodium "catalytic component" or "catalytic components" is intended to embrace the specified platinum group metal or metals present in any suitable form. Generally, reference in the claims or herein to platinum group metal or metals catalytic component or compounds embraces one or more platinum group metals in any suitable catalytic form. Table I-A demonstrates that the palladium-rhodium and platinum-rhodium combinations are rather ineffective for methane oxidation. The effect-iveness of rhodium as a methane oxidation catalyst is attenuated by the relatively high calcination temperature of 1600F. At a lower calcination temperature used in preparation of the catalyst, say 1100F, rhodium retains good methane oxidation characteristics. Howe~er, the catalytic partial oxidation catalyst of the present inven-tion may operate at ranges well above 1100F, which wculd probabl~
also reduce the effectiveness of rhodium for methane oxid-ation.
The tests in which the results of Table I-A were developed used a bed of the pla-tinum group metal-impregnated frit dispersed on alumina beads, rather than a monolithic body on which the frit is dispersed. The bed of frit-coated beads was of shallow depth to avoid excessive pressure drop.
The geometric confirguration of a 400 cell/in2 monolithic body provides more geometric surface area exposed to the reactant gas than does a bed of coated beads. Since the catalytic partial oxidation reactions of this invention are extremely rapid at the temperatures involved, the catalytic metals which reside on or near the surface of ~he catalyst body are predominatly involved in the reactions. The results of the tests wlth coated beads are indicative of results with monolithic bodies, but lower catalytic metal loading can be used with the latter as compared to metal loadings on beads, to attain equivalent results.
Table I-B shows the results of testing a monolithic body-supported catalyst on which a ceria-stabilized alumina frit impregnated with the indicated platinum group metals was dispersed upon a monolithic support. The alumina frit comprised 5~ by weight CeO2, balance A12O3, impregnated with one or two platinum group metals to provide the PGM
loadings indicated in Table I-B. The catalyst was calcined r " -- 18 ~

in air at 500C for two hours and then was aged 24 hou~s at 1800F in air.
Two different test gases, A and B, having the following composition were passed through the catalyst:

PARTS PER MILLION (VOL) OR
COMPOSITION VOLUME PER CENT
B
2 3~ 3%
CO lg6 1%

C2 10% 10%
H O 10% 10%
NO 50Oppm 50Oppm C2H4 300ppm C3H8 - 300ppm N2 balance balance Table I-B indicates the temperature in degrees centigrade for conversion of 50% by weight of the original amount of the component present, indicated under the column heading T50, and the temperature required for 75~ by weight - conversion, under the heading T75. A lower temperature accordingly indicates a more active catalyst. The results obtained are as follows; the platinum group metal (PGM) loading on the monlithic support is shown as grams of platinum group metal per cubic inch of monlithic catalyst.

TABLE I-B
PGM

Catalyst Weight Ratio PGM Loading Total PGM
Sample No. Pt:Pd (Pt/Pd (g/in3) Loading (g/in3) 1 100:0 .051/- .051 2 82:18 .044/.010 .054 3 58:42 .027/.019 .046 4 25:75 .011/.031 .042 0:100 - /.039 .039 3p 6 11:89 .003/.025 .028 `~ 7 100:0 .035/- .035 8 70:30 .034/.014 .048 ~ Iq_ Test Gas A Test Gas B
_ Component CO C2H4 CO C3H8 ;. Percent Conversion T50 T7s T50 75 T50 T7s T50 75 ~ Catalyst Sample No. C C C C
.
1. 325 335325 335265 275470 565 2. 270 275280 290280 285545 615 3. 235 250260 305260 265495 640 4. 235 245260 320260 270465 585

5. 230 235245 270245 255440 510 ; 6. 270 275275 315245 255430 555 7. 345 355350 365320 330495 550 8. 255 265265 290245 250485 585 The data of Table I-B demonstrate the lower temperatures at which a palladium containing catalyst will attain, respectively, 50% to 75% conversion of ethylene as compared to a platinum only catalyst. As mentioned above, the presence of platinum in addition to palladium provides effective catalyization of other species as well as providing enhanced poison resistance.

An exemplary mode of preparation of partial oxidation catalyst compositions utilizable in accordance with the present invention is set forth in the following Example 1.
Example 1 (a) To 229g of 2.5 wt % lanthia, 2.5 wt % baria -95 wt % A12O3 powder (a predominantly gamma alumina which has been stabilized by incorporation of lanthia and baria therein) is added a solution containing 20g Pt as H2Pt(OH)6 solubilized in monoethanolamine so as to give total volume of 229 ml. After mixing for 5 minutes, 25 ml of glacial ,~

acetic acid is added and the material is mixed a~ additio-nal 5 mintues before being dried and then calcined for one and one-half hours at 350C in air -to form a free flowing powder.
! (b) Similarly, to 229g of 2.5 wt % lanthia, 2.5 wt %
! baria 95 wt % A12O3 powder there is added 21 g Pd as Pd(NO3)3.
The material is mixed and reduced with 16 ml of N2H4H2O
so-lution with constant mixing. The impregnated powder is dried and then calcined for one and one-half hours at 375C in air.
(c) Two hundred gram of each of powder (a) and (b) is added to a 1/2 gallon size ball mill with appropriate amount of grinding media. To the powder is added 20 ml of glacial acetic acid and 550 ml of H2O. The sample is ball milled for 16 hours. The resulting slurry has a solids content of 43%, a pH of 4.0 and a viscosity of 337 cps and is used to coat a Corning cordierite monolith having a I diameter of 3.66", a length of 3" and 400 gas flow passages (of square cross section) per square inch of end face area.
The coating is accomplished by dipping the monolith in the slurry for 2 minutes, draining excess slurry and blowing the excess slurry from the gas flow passages with high pressure air. The resultant slurry-coated monolith is dried at 110C and calcined at 500C in air for 30 minutes.
The finished catalyst body contains 238g of platinum group metal per cubic foot of catalyst body volume at a weight `j ratio of paltinum to palladium of 1:1, with the platinum group metal dispersed on a ceria-stabilized alumina "washcoat" support layer. The catalyst body contains 1.64 grams per cubic inch of catalyst body of s~abilized alumina washcoat.

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A series of partial oxidation catalyst compositions utilizeable in accordance with the present invention were prepared by substantially the procedure described in Example 1, with appropriate modifications to obtain the reported loadings of different catalyst metals.
Each of the below described materials is a monolithic catalyst compositionO Except for the catalyst identified as CPO-5, in each case the honeycomb carrier is a C-400 cordierite carrier (400 gas flow passages per square inch of end face area) manufactured by Corning. The CPO-5 catalyst is on an alpha alumina monolith body, sold under the trademark TORVEX by DuPont~ and having 64 gas flow channels per square inch of end face area. The Corning cordierite monoliths have gas flow channels which are ~' square in cross section; those of the ~ORVEX monolith are hexagonal in cross section. The amount of platinum group metal on the catalyst is given in grams of elemental platinum group metal per cubic foot of monolith catalyst and the amount of refractory metal oxide coating is given in grams per cubit inch of monolith-catalyst~ The weight ratio of the platinum group metals in the order listed is given in parentheses~ Thus~ catalyst CPO-l in Table I, -~ for example, contains platinum and palladium in a weight -~ ratio of one part platinum to one part palladium. In eachcase, the refractory metal oxide coating is alumina, predominantly comprising gamma alumina stabilized as indicated, the respective weight percents of stabilizer 4~

being indicated, the balance comprising substantially alumina.
- TABLE I

Weight % and Alumina Supp~rt PG Metal PG Metal Stabilizer in Coating g/in Catalyst - ~omponent~ g/ft3 `Support Coating (%Stabilizer) CPO-l Pt, Pd (1:1) 219 5~ ceria 1.27 CPO-2 Pt, Pd (1~6 5~ ceria 1.64 CPO-3 Pt, Pd (1:4) 275 5% ceria 1.79 CPO-4 Pt, Pd (1:1) 310 5% ceria 2.32 CPO-5(*) Pt, Pd (1:1) 200 5% ceria 1 26 CPO-6 Pt, Pd Rh (9.5:(.51) 230 5% ceria 1.47 CPO-7 Pt, Pd (1:1) 186 2.5~ lanthia 2.5% baria 1.64 _ (*) TORVEX alpha alumina monolith; all others are cordierite ~ monoliths.
9 Genexally, the most preferred catalysts comprise platinum and palladium catalyst components and combinations thereof, with other platinum group metal catalytic components, preferably, combinations comprising 10-90% by weight palladium, preferably 25-75%, more preferably 40 to 60%, by weight palladium, and 90 to 10% by weight platinum, prefer-ably 75 to 25%, more preferably 60 to 40%, by weight platinum.
Generally, as the sulfur content of the hydrocarbon feed being treated in the first catalyst zone increases, a higher proporation of platinum to palladium is preferred. On the other hand, for feeds which have a relatively high methane content, an increasing proporation of palladium is preferred.

The monolithic configuration of the catalytic partial oxidation catlyst of the first catalyst zone affords a relatively low pressure drop across it as compared to the packed bed of a particulate support catalyst. This is particularly important in view of the increase in gas volume occasioned by the reactions taking place in the first catalyst zonel The total moles of product produced in the first catalyst zone is higher than the total moles of the inlet stream introduced therein.
The individual gas flow passages of the monolith also serve, in effect, as individual adiabatic chambersr thus helping to reduce heat loss and promote hydrocracking.
This is particularly so when the monolithic carrier comprises a ceramic-like material such as cordierite which has generally better heat insulating properties than do the metal substrates and, to this extent, the ceramic-type monolithic carriers are preferred over the metal substrate monolithic carriers~ Further~ as the monolith body becomes heated during operation, the gas in the up-stream portion of the monolith is preheated by the heat whlch is transferred back from the down-stream catalytic partial oxidation to the inlet thus facilitating desired hydrocracking and oxidation reactions~
i Steam Reforming Catalyst The steam reforming catalyst utilized in the second catalyst zone in accordance with the present invention may utilize a monolithic carrier as described _ ~4 -~L21~Z

above in connection with the partial oxidation catalyst or it may comprise a particulate support such as spheres, extrudates, granules, shaped members (such as rings or saddles) or the like. As used herein and in the claims, the term "particulate catalyst" or the like means catalysts of regularly or irregularly shaped particles or shaped members or combinations thereof. A preferred particulate support is alumina pellets or extrudate having a BET

(Brunnauer-Emmett-Teller) surface area of from about 10 to 200 square meters per gram. Alumina or alumiina stabilized with rare earth metal and/or alkaline earth metal oxides as described above, may be utilized as the pellets or extrudate. An alumina particulate support stabilized with lanthanum and barium oxides as described above is preferred.
The catalytically active metals for the optional s~eam reforming catalyst comprise the well known base metals (e.g., nickel) as well as platinum group metals, as stated above. A preferred platinum group metal steam reforming catalyst is a combination of platinum plus rhodium catalytic components with the rhodium comprising, on an elemental metal basis, from about 10 to 90% by weight, preferably 20 to 40% by weight, and the platinum comprising 90 to 10% by weight, preferably 80 to 60% by weight. The proportion of platinum and rhodium utilized will depend on the type of hydrocarbon feed to be treated in the processO Other platinum group metals may be utilized.

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Example 2 ~ a) A barium nitrate solution is prepared by dissolving 15909g Ba(NO ~2 in 1,650 ml of H2O. ~anthanum nitrate, in the amount of 264~9g La~NO312.6H2O is dissolved in the barium nitrate solution by mixing vigorously to yield a barium-lanthanum solution, to which is added to 3,000g of high surface area gamma alumina powder. The solution and powder are throughly mixed in a sigma blade mixer for 30 minutes~
~b~ The impregnated alumina resulting from step Ca~ was extruded through 1/16" diameter dies so as to give 1/16" diameter extrudate in lengths from 1/4" to 3/8".
rc) The extrudates from step (b) were dried at lloQC for 16 hours and then calcined 2 hours at 1,050QC
in air.
(d) ~ platinum-rhodium solution was prepared by dissol~ing 42.6g as H2Pt(OH)6 in monoethanolamine and 18g Rh as Rh(NO3)3.2H2) and combining the materials in H2O
to provide a solution having a volume of 1,186 ml and a pH of 0.7 after adjustment with concentrated HNO3.
(e) The platinum-rhodium solution of step (d) is added to the extrudate obtained in step (c) in a tumbling coater and mixed for 30 minutes. The impregnated extrudate is dried at 120~C for 4 hours and then-calcined for 30 minutes at 500QC ln air.
The resultant particulate steam reforming catalyst, designated SR-l, comprises 1.4 wt % platinum and 0.6 wt %
rhodium on La2O3-BaO stabilized gamma alumina extrudate.

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The catalysts of Examples 1 and 2 were utilized in test runs. Before describing these test runs, however, preferred embodiments of the apparatus of the present invention are descri~ed in some detail below.

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The ~eactor Vessel Preferably, the reactor utilized in the auto-thermal reforming process of the invention comprises a fixed ~ed, adiabatic reactor. Figure 1 shows a somewhat schematïc rendition of a preferred laboratory or pilot plant size reactor comprising a unitary vessel 1 within which a monolithic carrier paxtial oxidation catalyst 2 is disposed in flow communication via a passageway 3 with a bed of steam reforming catalyst 4. The vessel is suitably insulated by thermal insulating material 5 to reduce heat losses and to provide essentially a fixed bed, adiabatic reactor. Inlet lines 6, 7 and 8 feed a mixer 9 with, respectively, a hydrocarbon feed, steam and oxygen. The admixed reactants are introducted through an inlet line A into partial oxidation catalyst 2, thence via passage 3 into steam reforming bed 4 from which the contacted material is with-drawn through outlet line B. Valves, flow meters and heat exchange units, utilized in a manner known to those skilled in the art, are not shown in the schematic illustration of Figure 1.
Typically, in the apparatus of Figure 1, the monolithic carrier catalyst 2 is of cylindrical configuration, three quarters of an inch (1.9 cm) in diameter and nine inches (22.9 cm) long. The steam reforming bed is a cylindrical bed of particulate catalyst three inches (7.62 cm~ in diameter by nine and a quarter inches 123.5 cm) long. In operation, reactants are preheated with the oxidant stream being preheated separately from the hydrocarbon feed as a safety measure. After preheating, the streams were intimately mixed and immediately fed into the partial oxidation catalyst 2 of vessel l. Generally, all the oxygen present in the feed reacts within monolithic catalyst bed 2 to oxidize a portion, but not all, of the hydrocarbon feed, resulting in an increase in temperature due to the exothermic oxldation reaction. The heavier hydrocarbons are hydrocrakced in catalyst bed 2 to lighter, predominantly Cl hydrocarbons. The heated , partially oxidized and hydrocracked effluent from catalyst bed 2 is then passed through steam reforming catalyst bed 4 wherein the steam reforming reaction takes place. The product gases with-drawn via outlet B are cooled and unreacted water as well as any unreacted hydrocarbon feed, if any, is condensed and removed therefrom. The dry gas composition may be monitored by gas chromatography. The same principles of operation are followed in commercial embodiments of the apparatus.
Referring now to Figure 2, there is shown a schematic flow sheet illustration of a coal gasifier plant including an autothermal reforming section utilized to convert the liquid hydrocarbon by-product to secondary SG. A typical coal gasification plant, such as one according to the Lurgi design, includes a coal crushing and screening zone 10 to which coal is conveyed by suitable means for crushing and screening to segregate the coal particies by size, and any other treatments such as washing, etc., which may be required. Finely crushed coal is transmitted by means 12 to a power plant 14 to which water and air is supplied and in which the coal is burned to generate steam and electric power required in operation of the plant.
A coarse coal stream is fed via means 16 to a coal gasifier 18. Coal gasifier 18 may be of any suitable design, such as a Lurgi fixed bed reactor with rotating bottom grate, as briefly described a~ove. Steam is transmitted via lines 20, 22 from power plant 14 to coal gasifier 18.
An air separation plant 24 is supplied with air and, by any suitable technique, separates an oxygen stream from the air. Nitrogen is removed via line 26 and oxygen transmitted via lines 28, 30 to coal gasifier 18. It will be understood that in the case of the Lurgi coal gasifier design, as mentioned above, the steam and oxygen are trans-mitted through the gasifier counter-currently to descending stream of coarse coal particles. Under the conditions of temperature and pressure maintained in coal gasifier 18, gasifier SG is generated in coal gasifier 18, together with a liqued hydrocarbon by-product, both of which are removed from coal gasifier 18 via line 32. Ash is removed fro~
coal gasifier 18 via line 34. The gasifier SG and liquid hydrocarbon by-product are quenched in quench zone 36 from which the liquid hydrocarbon by-product is removed via line 38. The gasifier SG is transmitted via line 40 to a gas purification zone 42 in which carbon dioxide and hydrogen sulfide are removed therefrom by any suitable, known treatment.
The liquid hydrocarbon by-product is transmitted via line 38 to a gas-liquor separation zone 46 in which a gaseous fraction largely comprising ammonia and gasified phenolics is separated and removed via line 48 to a phenols separation zone 50, wherein off-gases including ammonia are separated and removed via line 52. The phenolics are - transmitted via line 54 to a blending zone 56 wherein the recovered phenolics are combined with the liquor separated from gas-liquor separation zone 46 via line 58.
A typical composition for the liquid hydrocarbon by-product, comprising the re-combined phenolics and liquor - 20 in line 60, is given in Table II.
TABLE II
Composition of Typical Hydrocarbon Liquid By-Product.
Component Average Formula Average Molecular Weight 1) Oil 13.5 18 180 2) Tar CllHloO 158 3) Phenolics C7H8O 108 4) Blend of 1), 2) and 3) 11 14.3 162 4~

The hydrocarbon by-product is transmitted via line 60 to a solids removal zone 62 wherein ash and heavy tar components are separated and removed via line 64. The hydrocarbons by-product may be treated by any suitable technique such as filtration and/or distillation for solids removal in zone 62.
In such process, metals and residual ash in the hydrocarbon ; by-products together with the heaviest tar fractions thereof are removed.
While any suitable method or combination of methods for removing solids and metals and the heavy tar fraction j may be utilized, a particularly useful and efficient method is the ARTSM treatment developed by Engelhard Corporation, the assignee of this application. This process utilizes ; an ARTCATTM material to carry out an asphalt residual treat-ment process which is highly effective and efficient in treating heavy petroleum or o-ther hydrocarbon containing : fractions to render them suitable for processing to more valuable materials.
The ash and heavy tar removed via line 64 may be cycled to power plant 14 for combustion of the combustible values therein to supplement the coal supplied as fuel thereto.
The thus-treated liquid hydrocarbon by-product is passed via line 66 through a heat exchanger 68 in which it is heated by indirect heat exchange as described below, and thence into a mixer 70.

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An oxygen stream is transmitted from air separation zone 24 via lines 28, 72 and steam is trans-. mitted from power plant 14 via lines 20, 74 through a heat exchanger 7~ for indirect heat exchange as explained below, thence to mixer 70 wherein the oxygen, steam and treated liquid hydrocarbon by-product are admixed for transmission ; via line 78 as the inlet stream to an autothermal reformer 80.
In reformer 80 the mixture of hydrocarbon by-product, steam and oxygen is passed through a catalytic partial oxidation catalyst supported on a monolithic honeycomb carrier disposed within neck portion 80a of reformer 80. Some, but not all of the hydrocarbon is catalytically oxidized within the first catalyst zone contained within neck portion 80a and the heavier unoxidized hydrocarbons are hydrocracKed to lighter constituents, mostly Cl hydrocarbons, with a very minor amount o~ C2 and C3 hydrocarbons. Methane is the pre-. dominant hydrocarbon product attained by the hydrocracking.
Depending on the specific nature of the hydrocarbon by-product fed to autothermal reformer 80 and the specific operating conditions utilized therin, the hydrocarbon by-product may be substantially entirely converted to a gaseous product containing H2, H2O, CH4, CO and CO2. However, under conditions in which a significant amount of heavier hydro-carbons would remain in the effluent from the first catalyst zone, a second catalyst zone may be disposed within main body 32, 3Z4;~

portion 80b of reformer 80. In the second catalyst zone a steam reforming reaction is catalyzed to convert hydrocarbons to hydrogen and carbon oxides.
It is of course desired to provide a high methane content in the secondary S~. Therefore, the exit temperature of the gases exiting via line 82 is preferably controlled to reduce the amount of C2 and C3 compounds formed within reformer 800 For example, an outlet temperature at line 82 of about 1,400F (760DC) has been found to be satisfactory to hydrocrack the material within the first catalyst zone to mostly Cl hydrocarbons, i.e., CO, CO2 and CH4. A methanation step as described below, for reacting carbon monoxide and hydrogen to methane, is conveniently utilized in the process.
It is therefore desirable to also control the conditions with-in reformer 80 to provide ln the gas obtained therein a molar ratio of hydrogen to carbon monoxide of slightly more than 3:1, the molar ratio in which the two gases react to form methane and H2O.
The effluent gas from reformer 80 is passed via line 82 through heat exchanger 76 to heat therein the oxygen and steam being transmitted via, respectively, lines 72 and 74 to mixer 70. The reformer effluent is then passed through heat exchanger 68 to heat the incoming liquid hydrocarbon by-product in line 66. The cooled reformer effluent gas is passed via line 82 to a quench zone 84 wherein the gas is cooled and water is separated therefrom via line 86. The z~

cooled secondary SG is passed via line 88 to be introduced into llne 40 to be admixed with the gasifier SG and the combined gases are then passed to gas purification zone 42 wherein acid gases, e.g., CO2 and H2S, are removed by known techniques and withdrawn via line 45. The purified combined gases are then passed to methanation zone 90 in which carbon monoxide contained therein is reacted with hydrogen to methane and H20, thereby increasing the overall methane content of the product synethesis gas which is withdrawn therefrom via line 92. It is this product, after drying to remove water, ! which is usually called "synthetic natural gas" or "SNG".
The following example exemplifies operating conditions and results obtained in treatlng a liquid hydrocarbon by-product which has been treated, as described above, for the removal of ash and heavy tars therefrom.
Example 3 Inlet Streamlb.-moles 2 -to C H20 to C
(78 in FIG. 2 Composition) per hour Ratio ~atio 1) Feed-item (4) of Table II 828 0.176 2.00 2) Steam 18,216 --3) Oxygen 1,602 _________________________ _______________ Temperature of Inlet Stream (78 in FI5. 2) = 800F ( a 27C) Pressure - 31 atmospheres.
h Effluent Gas (82 in FIG. 2) lb.-moles Composition per hour H2 8,325 11,212 CH4 2,300 CO 2,679 C2 ~,229 3 10 Temperature o Effluent Gas (82 in FIG. 2) - 1,400F (7~0C) Pressure of Effluent Gas = 30 atmospheres It will be seen that upon removal of H2O and CO2, and conversion of CO and H2 to CH4 by methanation~ a pre-dominantly hydroyen and methane containing synthetic natural gas is obtained from the secondary SG to supplement the synthesis natural gas obtained from the gasifier SNG.
While the invention has been described in detail with respect to specific preferred embodiments thereof, it will be appreciated that those skilled in the art, upon 20 reading and understanding of the foregoing will readily envision modifications and variations to the preferred embodi-3 ments which are nonetheless within the spirit and scope of the invention and of the claims.

~' .3.

Claims (31)

THE EMBODIMENTS OF THE INVENTION IN WHICH AN EXCLUSIVE
PROPERTY OR PRIVILEGE IS CLAIMED ARE DEFINED AS FOLLOWS:

1. In a coal gasification process in which coal is reacted with steam and oxygen to produce (i) a gasifier synthesis gas which is methanated to produce a synthetic natural gas, and (ii) a liquid hydrocarbon by-product, the improvement comprising preparing a secondary synthesis gas from said liquid hydrocarbon by-product and methanating the secondary synthesis gas, by the steps of:
(a) preheating an inlet stream comprising said liquid hydrocarbon by-product, H2O, and oxygen to a pre-heat temperature at least sufficiently high to initiate catalytic oxidation of said hydrocarbon by-product as defined below, but less than about 1200°F (649°C);
(b) introducing the preheated inlet stream into a first catalyst zone comprising a monolithic body having a plurality of gas flow passages extending therethrough and having a catalytically effective amount of a platinum and palladium catalytic component dispersed therein, the amounts of hydrocarbon by-product, H2O and oxygen introduced into said first catalyst zone being controlled to maintain an H2O to C ratio of from about 0.5 to 5 and an O2 to C ratio from about 0.15 to 0.4 in said inlet stream;
(c) contacting the preheated inlet stream within said first catalyst zone with said platinum and palladium catalytic component to initiate and sustain therein catalytic oxidation of a quantity, less than all, of said hydrocarbon by-product sufficient to attain a temperature within said first catalyst zone at least high enough to crack substan-tially all unoxidized C5 or heavier hydrocarbons in said by-product to C1 to C4 hydrocarbons, the temperature of at least a portion of said monolith being at least about 250 F (139 C) greater than the ignition temperature of said inlet stream, but not more than about 2000 F (1093 C), whereby to produce a first catalyst zone effluent comprising primarily methane, hydrogen, carbon monoxide, carbon dioxide and H2O and C2-C4 hydrocarbons;
(d) passing said effluent to a treatment zone for the removal of carbon dioxide and water therefrom;
(e) withdrawing the treated first catalyst zone effluent as a secondary synthesis gas; and (f) methanating said gasifier synthesis gas and said secondary synthesis gas to provide therefrom synthetic natural gas.

2. The process of claim 1 further including the step of passing said first catalyst zone effluent, while still at an elevated temperature, from said first catalyst zone to a second catalyst zone containing a steam reforming catalyst therein, and contacting the first zone effluent in said second catalyst zone with said steam reforming catalyst to react hydrocarbons therein with H2O to produce hydrogen and carbon oxides therefrom, and then passing the effluent of said second catalyst zone as said effluent to said treatment zone of step (d).

3. The process of claim 1 wherein at least about 50% by weight of said hydrocarbon by-product is converted to C1 hydrocarbons in said first catalyst zone.

4. The process of claim 2 wherein a total of at least about 98% by weight of said hydrocarbon by-product is converted in said first and second catalyst zones.

5. The process of claim 2 wherein said steam reforming catalyst comprises a platinum group metal component.

6. The process of claim 1 or claim 2 wherein the temperature of said first catalyst zone effluent is at least about 1,400°F (760°C).

7. The process of claim 1 or claim 2 wherein said secondary synthetic natural gas is combined with said gasifier synthetic natural gas to provide a combined product synthetic natural gas.

8. The process of claim 1 or claim 2 wherein the preheat temperature is from about 800°F to 1200°F (427°C to 649°C).

9. The process of claim 2 wherein said first catalyst zone is maintained at a temperature of from about 1400°F to 2000°F (760°C to 1093°C) and the first zone effluent is introduced into said second catalyst zone at substantially the same temperature.

10. The process of claim 1 or claim 2 wherein a volumetric hourly rate of at least 100,000 volumes of throughput per volume of catalyst is maintained in said first catalyst zone.

11. The process of claim 2 wherein a volumetric hourly rate of at least 100,000 volumes of throughput per volume of catalyst is maintained in said first catalyst zone and a volumetric hourly rate of from about 2,000 to 20,000 volumes of throughput per volume of catalyst is maintained in said second catalyst zone.

12. The process of claim 1 or claim 2 wherein said platinum group metal component of said first catalyst zone comprises a catalytically effective amount of palladium and platinum catalytic component and, optionally, rhodium catalytic component distended upon a refractory metal oxide support layer carried on said monolithic body.

13. The process of claim 1 wherein said catalytic components of said first catalyst zone comprise, on an elemental metal basis, about 10 to 90% by weight palladium and about 90 to 10% by weight platinum.

14. The process of claim 13 wherein said catalytic components of said first catalyst zone comprise about 25 to 75% by weight palladium, and about 75 to 25% by weight platinum.

15. The process of claim 2 wherein said steam reforming catalyst comprises a catalytically effective amount of rhodium and platinum catalytic components distended upon a refractory metal oxide support.

16. The process of claim 15 wherein said catalytic components of said steam reforming catalyst comprises, on an elemental metal basis, about 10 to 90% by weight rhodium, and about 90 to 10 by weight platinum.

17. The process of claim 16 wherein said catalytic components of said steam reforming zone comprise about 20 to 40% by weight rhodium and about 80 to 60% by weight platinum.

18. The process of claim 1 or claim 2 or claim 3 carried out at a pressure of from about 50 to 1500 psig.

19. In a coal gasification process in which coal is reacted with steam and oxygen to produce a gasifier synthesis gas and liquid hydrocarbon by-product, the improve-ment comprising preparing a secondary synthesis gas from said liquid hydrocarbon by-product by the steps of:
(a) preheating an inlet stream comprising said liquid hydrocarbon by-product, H2O and oxygen to a preheat temperature of from about 800°F to 1200°F (427°C to 649°C);
(b) introducing the preheated inlet stream into a first catalyst zone comprising a monolithic body having a plurality of gas flow passages extending therethrough and comprising a catalytically effective amount of palladium and platinum catalytic components dispersed therein, the amounts of hydrocarbon by-product, H2O and oxygen introduced into said first catalyst zone being controlled to maintain an H2O to C ratio from about 0.5 to 5 and O2 to C ratio of from about 0.15 to 0.4 in said inlet stream;
(c) contacting the preheated inlet stream within said first catalyst zone with said palladium and platinum catalytic components at a volumetric hourly rate of at least about 100,000 volumes of throughput per volume of catalyst per hour to initiate and sustain therein catalytic oxidation of a quantity, less than all, of said hydrocarbon by-product sufficient to attain a temperature within said first catalyst zone of from about 1400°F to 2000°F (760°C to 1093°C) and cracking substantially all unoxidized C5 or heavier hydrocarbons in said by-product to a mixture of C1 to C4 hydrocarbons and predominating in C1 hydrocarbons, whereby to produce a first catalyst zone effluent comprising methane, hydrogen, carbon monoxide, carbon dioxide and H2O;
(d) passing said effluent to a treatment zone for the removal of carbon dioxide and water therefrom; and (e) withdrawing the treated first catalyst zone effluent as a secondary synthesis gas.

20. The process of claim 19 further including the step of passing said first catalyst zone effluent, while still at a temperature of from about 1400°F to 2000°F (760°C to 1093°C) from said first catalyst zone to a second catalyst zone containing a platinum group metal steam reforming catalyst therein, and contacting the first zone effluent in said second catalyst zone with said steam reforming catalyst at a volumetric hourly rate of from about 2,000 to 20,000 volumes of throughput per volume of catalyst per hour to react hydro-carbons therein with H2O to produce hydrogen and carbon oxides therefrom, and then passing the effluent of said second catalyst zone as said effluent to said treatment zone of step (c).

21. The process of claim 19 wherein at least about 50% by weight of said hydrocarbon by-product is converted to C1 hydrocarbons in said first catalyst zone.

22. The process of claim 20 wherein a total of at least about 98% by weight of said hydrocarbon by-product is converted to C1 hydrocarbons in said first and second catalyst zones.

23. The process of claim 19 or claim 20 wherein said first catalyst zone comprises palladium and platinum and, optionally, rhodium catalytic components distended upon a refractory metal oxide support layer carried on said mono-lithic body.

24. The process of claim 19 wherein said catalytic components of said first catalyst zone comprise, on an elemental metal basis, about 10 to 90% by weight palladium, and 90 to 10% by weight platinum.

25. The process of claim 24 wherein said catalytic components of said first catalyst comprise about 25 to 75%
by weigh palladium and 75 to 25% by weight platinum.

26. The process of claim 20 wherein said steam reforming catalyst comprises platinum and rhodium catalytic components distended upon a refractory metal oxide support.

27. The process of claim 26 wherein said catalytic components of said second catalyst zone comprise, on an elemental basis, about 10 to 90% by weight rhodium and 90 to 10% by weight platinum.

28. The process of claim 27 wherein said catalytic components of said second catalyst zone comprise about 20 to 40% by weight rhodium and 80 to 60% by weight platinum.

29. The process of claim 19 or claim 20 further including treating said liquid hydrocarbon by-product to remove solids, metals, and heavy tar therefrom prior to passing said by-product to said first catalyst zone.

30. The process of claim 20 further including methanating said secondary synthesis gas and said gasifier synthesis gas.

31. The process of claim 14 or claim 25 wherein said catalytic components of said first catalyst zone comprise about 40 to 60% by weight palladium, and about 60 to 40% by weight platinum.