CA1210567A - Method of ammonia production - Google Patents

Abstract

ABSTRACT OF THE DISCLOSURE
A process for ammonia production includes preparing a hydrogen and nitrogen synthesis gas by autothermally reforming a hydrocarbonaceous feed, such as a hydrocarbon feed, with oxygen enriched air and steam in an autothermal reformer utilizing a first, monolithic catalyst having a palladium and platinum containing catalyst therein, and a second, platinum group metal steam reforming catalyst. The autothermal reformer provides a relatively simple and compact reactor within which a wide variety of hydrocarboneaceous feeds, from heavy hydrocarbon feedstocks to natural gas, may be utilized for the synthesis process, which is advantageously carried out at an elevated pressure. By controlling H2O to carbon, O2 to carbon, and N2 to carbon feed ratios, a given feed can be autothermally reformed to yield a synthesis gas of desired ammonia synthesis ratio of about 3:1 H2:N2.

Description

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11 ~
li l BACRGROUND O~ THE INVENTION

¦ The present invention is concerned with the production of ammonia from a ~ hydrocarbonaceous feed, including the preparation of an ammonia synthesis gas ¦ containing hydrogen, more particularly, with the preparation of a synthesis gas ¦ comprising hydrogen and nitrogen in selècted molar ratios and the reacting of these S ¦ constituents to form ammonia.
¦ As used herein and in the claims, the term "hydrocarbonaceous feed" is intended ¦ to include, without limitation, hydrocarbon feeds of all types, as well as alcohols ¦¦ such as ethanol, methanol and mixtures thereof, and biomass-derived feeds which I¦ normally contain compounds of carbon, hydrogen and oxygen, and sometimes ¦ compounds of nitrogen and sulfur, as well as one or more of the foregoing elements ¦ in elemental form. Such biomass-derived feeds may be obtained by any suitable ¦ process such as fermentation of grains or other materials including food products I generally, from treatment of agricultural by~products and ~aste products, or by ¦ distillation or combustion (with insufficient oxygen for stoichiometric reaction) of wood and/or other cellulosic products and by-products.
Reference hereinbelow and in the claims to a "synthesis gas" or "synthesis gases" will be understood to mean a gas mixture comprising hydrogen and nitrogen, possibly with other constituents, such as H2O, carbon dioxide, sulfur compounds, ¦ inert gases and the like. Such other constituents are often removed prior to or ¦ during the ammonia synthesis operation in which the synthesis gas is reacted.
Ammonia synthesis g~s may be prepared by the partial oxidation of hydrocarbonaceous feeds, such as the heavier hydrocarbons, e.g., fuel oil and coal, and by steam reforming of hydrocarbonaceous feeds, such as the lighter hydrocarbons, e.g., natural gas and naphthas. Processes to derive synthesis gases from methanol or coal derived hydrocarbons are also known.

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jGenerally, difficulties associated with the preparation of synthesis gases from ¦ heavier feedstocks favor the use of light naphthas or natural gas uhen a hydrocarbon is the source of the hydrocarbonaceous feed. However, shortages of such light ¦ hydrocarbon feeds indicate the need for an economical process for generating a j synthesis gas from heavier hydrocarbon feedstocks, such as normally liquid hydro-carbons.
jSteam reforming is a well known method for generating synthesis gas from light hydrocarbon feeds and is carried out by supplying heat to a mixture of steam and a hydrocarbon feed while contacting the mixture with a suitable catalyst, usually nickel. HoweYer, steam reforming is generally limited to paraffinic naphtha and lighter feeds which have been de-sulfurized and treated to remove nitrogen compounds, because of difficulties in attempting to steam reform heavier hydrocarbons and the poisoning of steam reforming catalysts by sulfur and nitrogen compounds. Further9 I steam reforming for ammonia production generally must be carried out in two stages, ! a primary steam reforming stage and a secondary steam reforming stage. The former ! requires a tubular furnace containing catalysts disposed within the tubes and means ¦ to supply heat to the tubes in order to sustain the endothermic steam reforming reaction. Secondary reforming is usually carried out in a separate vessel to which ¦ oxygen is added to carry out a partial combustion to supply additional heat for the I endothermic steam reforming step. Such processes are well known to the art, as indicated by the article "Checklist for High Pressure Reforming" by Orlando J.
¦ Quartulli appearing at pages 151-162 of Hydrocarbon Processin~, April, 196j, Vol. 44, ¦ No. 4.
!l Another known method of obtaining hydrogen from a hydrocarbon feed is the ~ partial oxidation process in which the feed is introduced into an oxidation zone ¦ maintained in a fuel rich mode so that only a portion of the feed is oxidized. Steam may be injected into the prrtial oxi~ on reactor vessel to react with the feed nnd with products of the partial oxidation reaction. The process is not catalytic and requires high temperatures to carry the reactions to completion, resul~ing in a relatively high oxygen consumption. On the other hand, the partial o~idation process has the advantage that it is able to readily handle hydrocarbon liquids hea~ier than paraffinic naphthas and can even utilize coal as the source of the hydrocarbon feed.
Catalytic autothermal reforming of hydrocarbon liquids is also known in the art, as evidenced by a paper Catalytic Autothermal Reforming of Hydrocarbon Liquids by Maria Flytzani-Stephanopoulos and Gerald E. Voecks~ presented at the AmericanInstitute of Chemica~ Engineers' 90th National Meeting, Houston, Texas, April 5-9, 1981. Autothermal reforming is defined therein as the utilization of catalytic partial oxidation in the presence of added steam, which is said to increase the hydrogenyield because of simultaneous (with catalytic partial oxidation) steam reforming being attained. Steam, air and a No. 2 fuel oil are injected through three different nickel particulate ~atalysts and the resulting product gases contained nitrogen, hydrogen and carbon oxides.
In Brennstoff-Chemie 46, No. 4, p. 23 (1965), a German publication, Von P.
Schmulder describes a Badische Anilin and Soda Fabrik (BASF) process for autothermal reforming of gasoline utilizing a first, pelletized, platinum catalyst zone followed by a second, pelletized nickel catalyst zone. A portion of the product gas is recycled to the process.
Dis~losure of the utilization of a noble metal catalyzed monolith to carry out a catalytic partial oxidation to convert more than half of the hydrocarbon feedstock upstream of a steam reforming zone is discussed in abstract entitled "Evaluation of Steam Reforming Catalyst for use in the Autothermal Reforming OI Hydrocarbon Feed Stocks" by R.M. Yarrington, I.R. Feins, and H.S. Hwang (National Fuel Cell Seminar, July 14-16, 1980, San Diego, California.) The abstract noted the uniqueability of rhodium to steam reform light olefins with little col;e formntion nnd noted ! . 12~567 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.
U.S. Patent 4,054,407, assigned to the assignee of this application, discloses two-stage catalytic oxidation using platinum group metal catalytic components dispersed on a monolithic body. At least the stoichiometric amount of air is supplied over the two stages and steam is not employed.
U.S. Patent 3,481,722, assigned to the assignee of this application, discloses atwo-stage process for steam reforming normally liquid hydrocarbons using a platinum group metal catalyst in the first stage. Steam and hydrogen, the latter of whichmay be obtained by partially cracking the hydrocarbon feed, are combined with the feed to the process.
The present invention provides a highly efficient method for producing synthesisgases from hydrocarbonaceous feeds, including hydrocarbons, which attains excellent yields in a relatively compact and simple apparatus.
SUMMA~Y OF THE INVENTION
In accordance with the present invention there is provided a process for the production of ammonia from a hydrocarbonaceous feed, the process comprising the following steps. Preheating an inlet stream comprising a hydrocarbonaceous feed,H2O, air and oxygen to a preheat temperature sufficiently high to initiate catalytic oxidation of the feed as defined below. 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 palladium and platinum catalytic components dispersed therein, the amounts of feed, H20 and oxygen introduced into said first catalyst zone being controlled to maintain in said inlet stream an H2O to C ratio of from about 0.5 to 5 and 2 to C ratio ~f from about 0.35 to l).dS. Contrc ~ the preherted inlet strerm within the first ~ ~21~567 li catalyst zone with the palladium and platinum catalytic components to initiate and sustain therein catalytic oxidation of the feed to produce hydrogen and carbon oxides therefrom, the temperature of at least a portion OI said monolithic body being at least about 250F~139 C) greater than the ignition temperature of said inlet stream and oxidizing in the first catalyst zone a quantity, less than all, of the feed, which quantity is sufficient to heat such first zone effluent to an elevated temperature high enough to catalytically steam reform, within a second catalyst zone definedbelow, hydrocarbon remaining in such first zone effluent without supplying external heat thereto, passing the first zone effluent, while still at Hn elevated temperature, from the first catalyst zone to a second catalyst zone containing a platinum group ¦ metal steam reforming catalyst therein, and contacting the first zone effluent in the second catalyst zone with the reforming catalyst to react hydrocarbons therein with H2O to produce hydrogen and carbon oxides therefrom, and withdrawing the effluent of the second catalyst zone as a hydrogen-containing synthesis gas and removing heat therefrom to cool the synthesis gas. ~eacting cerbon monoxide in the synthesis gas with H2O to produce hydrogen. Removing sulfur containing ¦ compounds and H28 from the synthesis gas. Passing the synthesis gas into an ¦, ammonia synthesis loop to react the hydrogen with nitrogen thereof over an ammonia synthesis catalyst at ammonia synthesis conditions, and withdrawing ammonia as product from the ammonia synthesis loop.
In one aspect of the invention, air is introduced as part of the inlet stream ¦ with the proportion of air to oxygen being such as to provide oxygen enrichment of ¦ the air in the inlet stream to at least about 33 volume percent oxygen, preferably I about 33 to 50 volume percent oxygen, whereby to provide a nitrogen-containing~ hydrogen-rich synthesis gas.
Preferred aspects of the invention provide one or more of the following individual features. The hydrocarbonaceous feed may be a hydrocarbon feed. The preheat I ~z~ s67 .. l temperature may be from about 8000F to 14000F (~270C to 7600C) and the first catalyst zone may be maintained at a temperature of from about 1750F to 2400F
(954C to 1316C), with the first zone effluent being introduced into the secondcatalyst zone at substantially the same temperature. A volumetric hourly rate ofat 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. The first catalyst zone comprises palladium, platinum, and, optionally, rhodium catalytic components distended upon a refractory metal oxide support layer carried on the monolithic body. The steam reforming catalyst may comprise one or both of platinum and rhodium catalytic components, preferably also distended upon a refractory metal oxide support.
In other preferred aspects of the invention, the process may be carried out at a pressure of from about 100 to 1500 psia, more preferably 100 to 1100 psia. Thefeed may be a normally liquid hydrocarbon or a normally gaseous hydrocarbon.
In another aspect of the invention, the feed is a hydrocarbon and there is included the additional step of treating the synthesis gas withdrawn from the second catalyst zone to convert carbon monoxide therein to carbon dioxide and then removing carbon dioxide from the synthesis gas to provide a carbon oxides-depleted, nitrogen containing synthesis gas. The step of converting cflrbon monoxide to carbon dioxide may comprise adding additional steam (H2O) to the synthesis gas withdrawn from the second catalyst zone and contacting the resulting mixture with a catalyst effective for the water gas shift reaction to convert carbon monoxide to hydrogen ~nd carbon dioxide.
Any residual carbon monoxide remaining after the shift conversion ~nay be converted to carbon dioxide by adding oxygen, preferably as air, to the synthesis gas withdrawn from the low temperature shift reactor and COntRcting the resulting ll 12~567 1~

mixture with a catalyst effective for the selective oxidation of carbon monoxide to carbon dioxide.
Preferably, the final molar ratio of hydrogen to nitrogen in the synthesis gas is from about 2.9 to 3.1:1.
Other aspects of the invention provide that the amounts of feed, H2O, and oxygen, and air introduced into the first catalyst zone are controlled to maintain in the inlet stream an H2O to C ratio of from about 1 to 4, and an 2 to C ratio offrom about 0.5 to 0.6.

BRIEF DESCRIPTION OF THE DRAWINGS
I
¦ Figure 1 is a schematic elevation view in cross section of a laboratory or pilot ¦ 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 an ammonia synthesis plant, 15 ¦ including a synthesis gas making section in accordance with one embodiment of the present invention, the synthesis gas section including an autothermal reformer.

DESCRIPrl~ON OF THE PRE~FE~RED E~BODIMENTS
In a preferred embodiment of the present invention, a synthesis gas making plant includes an autothermal reformer for generation of hydrogen from a hydrocar-bonaceous feed. The autothermal reformer includes a first catalyst zone for carrying out catalytic partial oxidation, an exothermic reaction, and a second catalyst zone for carrying out steam reforming, an endothermic reaction. Some steam reforming also appears $o take place in the first catalyst zone and thereby moderates somewhat the temperatures attained in the first catalyst zone inasmuch as the endothermicsteam reforming absorbs some of the heat generated by the partial oxidation step.
The net reaction in the first catalyst zone is however exothermic and the zone is I

¦ therefore referred to as an exothermic catalyst zone. The exotherm~c, first catalyst I zone comprises a monolithic catalyst carrier on which a platinum group metal catalyst j is dispersed. Such catalyst can effectively catalyze the parti~ oxidation of not only ¦ gaseous and light hydrocarbon liquids such as natural gas or paraffinic naphtha, but ¦ of heavier hydrocarbons such as diesel oil, number 2 fuel oil, and coal derived liquids.
As compared to a non-catalytic combustion process such as conventional non-catalytic ¦l partial oxidation, catalytic partial oxidation enables the utilziation of lesser amounts ~¦ of oxygen and lower temperature levels to both oxidize a portion of the feed and I crack hea~rier hydrocarbon feedstocks to lighter hydrocarbon fractions, while raising ¦ the temperature of the reactant mass for subsequent treatment. The ability to use I less oxygen in the autothermal reformer is advantageous in making an ammonia ¦ synthesis gas in that a portion of the total air to be fed to the process to attain ¦ the preselected nitrogen to hydrogen ratio in the product synthesis gas may thus be ! diverted from the reformer to a downstream selective oxidation process. This has I advantages in increasing the overall yield of the process.
Generally, at least about half the hydrocarbon feedstock is partially oxidized in the catalytic partial oxidation zone to produce primarily carbon oxides and hydrogen and the heat required for the endothermic steam reforming reaction, which takes place in the second catalyst zone. Substantially all of the limited amount of oxygen ¦ introduced into the first catalyst zone may be consumed in the catalytic partial ¦ oxidation step. At the temperatures maintained in the catalytic oxidation zone, and ¦ in the presence of the product hydrogen and catalyst utilized in the first zone, a degree of hydrocracking of unoxidized (~5 and heavier hydrocarbon feed, if any is ¦ present, takes place to form C4 and lighter compounds. The effluent gas from the ¦ first catalyst zone thus contains primarily CO, H2 and H2O and, ~vhen a hydrocarbon ¦ feed is used, C2 to C4 and other lighter hydrocarbons, including olefins, and, depending upon the sulfur content of the feedstock, H2S and CQS. If air, as opposed to oxygen, ~ 567 is used RS the oxidant gas, the effluent will of course also contain N2.
The endothermic, second catalyst zone may contain any suitable platinum group metal steam reforming catalystO Usually, the steam reforming catalyst will be utilized in the form of a particulste bed comprised of spheres, extrudates, granules, configured packing material, e.g., rings, saddles or the like, or 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 be used in the second catalyst zone, as is used in the first catalyst ~one. The catalystmetal for the steam reforming zone may comprise any platinum group metals such as, e.g., platinum or rhodium or combinations thereof.
The combination of features provided by the process of the present invention provides a highly efficient process of manufacturing a synthesis gas for ammoniaproduction by converting various types of hydrocarbonaceous feeds, including hydrocarbon feeds, to a hydrogen and nitrogen-rich ammonia synthesis gas. The 15 ¦ combination of the monolithic platinum group metal partial oxidation catalyst with a platinum group metal steam reforming catalyst provides a great flexibility in handling diverse feedstocks, including heavy hydrocarbon feedstocks not normallywell-suited for generating a synthesis gas. For example, a uride range of 2 to carbon (atoms of carbon in the feed) and ~2O to carbon ratios may be used.

The Monolithic Partial Oxidation Catalyst The partial oxidation catalyst is provided on a 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 therethrough. Such monolithic carriermembers are often referred to as "honeycomb" type carriers and are well known inthe art. A preferred form of such carrier is made of a refractory, substantiallyinert rigid material which is capable of maintaining its shape and a sufficient degree of mechanical strength at high temperatures, for example, up to about 3272F

'I ~21(~5~7 ¦¦ (1800C). Typically, a material is selected for ~he support which exhibits a low thermal coefficient of expansion, good thermal shock resistance and, though not always, low thermal conductivity. Two general types of material of construction for such carriers are known. One is a ceramic-like porous material comprised of S one or more metal oxides, for example; alumina, alumina-silica, alumina-silica-titania, mullite, cordierite, zirconia, ~irconia-spinel, zirconia-mullite, silicon carbide, etc. A
particularly preferred and commercially available material of construction for operations below about 2000F (1093C) is cordierite, which is an alumina-magnesia-silica material. For applications involving operations above about 2000F (1093C~, ¦ an alumina-silica-titania material is preferred. 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, preferably 200-600 gas flow channel 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, 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, from about 200 to 1,200 per square ~nch ofend face area of the tubular roll.
Although the ceramic-like metal oxide materials such as cordierite or alumina-silica-titania are somewhat porous and rough-textured, they nonetheless have 121~567 a relatively low surface area wi~h respect to catalyst support requirements and, of course, a stainless steel or other metal support is essentially smoothO Accordingly, a suitable high surface area refractory metal oxide support layer is deposited on the carrier to serve as a support upon which finely dispersed cataly$ic metal may bedistended. 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, beryllia, 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 claims9 "transition alumina" includes gamma, chi, eta, kappa, theta and delta forms and mixtures thereof.
An alumina comprising or predominating in garnma 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 phastransition to alpha alumina, which has a relatively low 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 oflanthanum and barium is a preferred stabilizer for transition alumina.
The catalytic metal of the catalytic partial oxidation catalyst comprises platinum and palladium and, optionally, rhodium. The platinum group metal may optionally be supplemented with one or more base metals, particularly base metals of Group vn and metals of Groups VB, VIB, and VIIB of the Perioclic Table of Elements.
Preferably, one or more of chromium, copper, vanadium, cobalt, nîckel and iron may be employed.

~L21~567 Desir~ble catalysts for p~rti81 o~idation should h~ve the following properties:
¦ They should be able to operate effectively under conditions varying from oxidizing at the inlet to reducing at the exit; they should operate effectively 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 H20 to form carbon monoxide and hydrogen thereby permitting only a low level of carbon ~ on the catalyst surface; they must be able to resist poisoning from such common I poisons as sulfur and halogen compounds; further! all these requirements must be satisfied simultaneously. For example, in some otherwise suitable catalysts, carbon monoxide may be retained by the catalyst metal at low temperatures thereby decreasing or modifying its activity. The combination of platinum and palladium is a highly efficient oxidation catalyst for the purposes of the present invention.
Genera~ly, the catalytic activity of platinum-palladium combination catalysts is not simply an arithmetic combination of their respective catalytic activities; the disclosed range of proportions of platinum md palladium have been found to provide efficient and effective catalytic activity in treating a rather wide range of hydrocarbonaceous, particularly hydrocarbon, feeds with good resistance to high temperature operation

2 0 and catalyst poisons.
The following data compares the effectiveness of palladium, rhodium and platinum, rcspectively, for the oxidation of methane and further compares the efficacy of, respectively, palladium-platinum, palladium-rhodium and platinum-rhodium combined catalysts for oxidation of methane.
The catalysts of Table I-A comprise a lanthia-chromia-alumina frit impregnated with the platinum group metals by techniques as described above. The frit has the following composition:
Composition WeiFht Percent La2O3 3.8 Cr23 1.8 ~123 94 4 The lanthia-chromia stabilized alumina is then impregnnted with the plntinllm grollp metal and calcined in air for four hollrs nt 230~ nlld for nn ndditionnl four hours at 1600F. Three catalysts of different plntillum metal londin~s were prepnred ns follows:

l! I

, ei~ht Percent Sample No. Pd Pt RhTotal PGM

4063U-1 3.425.95 - 9.37 4063 R-l 4.58 - 4.52 9.10 1 40~3V-1 - 5.62 3.14 ~.76 1-.

The resultant platinum group metal (YGM) impregnated alumina frit was deposited on alumina beads and the thus-coated beads were placed in a shallow bed and tested I by passing a 1% (volume) methane 99% (volume~ air feed at about atmospheric ¦ pressure through the catalyst. An electric heater was used to cvclically 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 cycle.
The results are shown in the following Table I-A.

l l I
l 5 ~ TABLE I-A

PGM
li Sample (Mole Ignition Weight Percent of Original Methane Con-¦ No. Ratio Temp.F tent Converted at Indicated Temperature (F) I

! ! 4053U-1 Pd,Pt(l:l) 610 - 3 10 26 60 80 Il 4063R-l Pd,Rh(l:l) 710 - - 2 5 9 12 20 ~1 4063V-1 Pt,Rh(l:l) 730 - - 1 1 3 5 I!
1~ These data demonstrate the ability of platinum-palladium catalyst to promote catalytic oxidation of m ethane o~er r wide rrnge of temperature.
I

-i3 -}567 , Il Rhodium may optionally be included with the platinum and palldium. The combined platinum group metal catalysts of the invention also have a significant advantage in thc ability to catalyze the autothermal reactions at quite low ratios of H20 to ~l carbon (atoms of carbon in the feed) and oxygen to carbon, without significant carbon 1 deposition on the catalyst. This important feature provides flexibility in selecting 'I H20 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 metalI or metals present, or as compounds such as an oxide of the platinum group metal.
As used in the claims, the terms palladium, platinum andlor 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 components !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 effectiveness 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 , llOQF, rhodium retains good methane oxidation characteristics. However, the catalytie partial oxidation catalyst of the present invention may operate at ranges weU above 1100F, which would probably also reduce the effectiveness of rhodium ` for methane oxidation.

The tests in which the results of Table I-A were developed used a bed of the ~ ~IPtimlm group metal-impregnated frit dispersed on alumina beflds, rnther thnn a mcnolithic bodyon which the frit is dispersed. The bed of frit-coated beads was of shalion depth to avoid excessive pressure drop. The geometric configuration of R400 cell/in2 monolithic body provides more geometric surface e~posed to the renctnnt gas than does a bed of coatecl beads. The catalytic pnrtirl ).~idntioll reaction~ cf !!

,1 ,j -14-12~1~567 ~
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. 1.

this invention are extremely rapid at the temperatures involved. Therefore, the I eatalytie metals on the surfaee of the eatalyst body are predominantly involved in !
j the reaetions. The results of the tests with coated beads are indicative of results ¦
¦ with monolithic bodies, but lower catalytic metal loading can be used with the latter j ¦ as eompared to metal loadings on beads, to attain equivalent results.
Table I-B shows the results of testing a monolithic body-supported catalyst on I
whieh a ceria-stabilized alumina frit impregnated with the indieated platinum group metals was dispersed upon a monolithic support. The alumina frit comprised 5% by weight CeO2, balanee A12O3, impregnated with one or two platinum group metals ¦ to provide the loadings indieated in Table I-B. The catalyst was calcined in air at 500C for two hours and then was aged 24 hours at 1800F in air.
Two different test gases, A and B, having the following eompositions were passed through the catalyst: I

l l PARTS PER MILLION (VOL) I! COMPOSITION O~ VOLUME PERCENT
¦l A B

2 3% 3%
11 CO 1% 1%
CO2 10% 10~ 1, H20 10% 10%
I NO 500 ppm 500 ppm ¦ C,~H4 300 ppm _ _ C3H8 __ 300 ppm N2 Balanee Balance Table I-B indicates the temperature in degrees centri~rade necessary for `
conversion of 50% by weight of the original amount of the eomponent 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 aetive eatalyst. The results obtained are as follows: the platinum group metal (PGM) loading on the monolithic support is shown as grams of platinum group metal per cubie ineh of monolithie eatalyst~

l ll iZl~}567 I
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TABLE I-B

PGM
CatPlystWeight Ratio PGM Loadi~g Total PGM 3 Sam~le No. Pt:Pd Pt/Pd (F/in )Loading (g/in ) 892-68-S~P 100:00 .051/- .051 892-69-S~P 82:18 .044/.010 .054 1892-70-SSP 58:42 .027/.Olg .046 892-71-SSP 25:75 .~11/.031 .042 892-72-SSP 0:100 -/.039 .039 892-76-SSP 11:89 .003/.025 .028 P-PX 100:00 .035/- .035 P-PXIIB 70:30 .034/.014 .048 Test Gas A Test Gas B

~omponent CO C2H4 CO C3H8 Percent Conversion T50 75 I T50 T75T50 T75 , T50 T75 ~,lmple No. l C I C l C I C

Pt) 892-68-SSP 325 335 325 335 265 275 470 565 Pd/Pt) 892-69-SSP 270 275 280 290 280 285 545 615 Pd/Pt) 892-70-SSP 235 250 260 305 260 265 495 640 Pd/Pt) 892-71-SSP 235 245 260 320 260 270 465 585 Pd) 892-72-SSP 230 235 245 270 245 255 440 510 Pd/Pt) 892-76-SSP 270 275 275 315 245 255 430 555 ~Pt) P-PX 345 355 350 365 320 330 495 550 IIB 12ss26512652901245250 1 ~2~(~5~i7 The data of Table I-B demonstrates the lower temperatures at which a palladium-containing catalyst will attain, respectively, 50% and 75% conversion of ethylene as comparecl to a platinum only catalyst. The platinum provides effective catalyzation of other species as well as providing enhanced poison resistence?
particularly to sulfur and sulfur compounds.
An exemplary mode of preparation of partial oxidation catalyst compositions utilizable in accordance with the present invention is set forth in the following Example I.

Example 1 (a) To 229 g of 5 wt % CeO2 - 95 wt % A12O3 powder (a predominantly gamma alumina which has been stabilized by incorporation of ceria therein~ is added a solution containing 21g Pt as H2Pt(OH36 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 an additional 5 minutes 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 229 g of 5 wt % CeO2 - 95 wt % A12O3 powder there is added 21 g Pd as Pd(nO3)3. The material is mixed and reduced with 16 ml of N2H4 H2O solution with constant mixing. The impregnated powder is dried and thencalcined for one and one-half hours at 375C in air.
~c) Two hundred grams 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 po-vder is added 20 ml of glacial acetic acid and 550 ml of H20. The sample is ball milled for 16 hours. The resulting slurry has a solids content of 4396, a pH of 4.0 and a viscosity of 337 cps and is used to coat a Corning cordierite monolith havin~ a diameter of ~ J21~567 ~3.66", a length of 3" and 401) g~s flow passages (of square oross seotion) per square inch of end face area. The coating is acomplished by dipping the monolith in theslurry 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 ratio of platinum to palladium of 1:1, with the platinum groups metal dispersed on a ceria-stabilized alumina "washcoat" support layer. The catalyst body contains 1.4 grams per cubic inch of catalyst body of stabilized alumina washcoat.
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 ofjdifferent catalyst metals. Each of the below described materials is a monolithic ¦catalyst composition. Except for the catalyst identified as CPO-5, in each case the Ihoneycomb 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 ~ITORVEX 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. The weight ratio of the platinum group metals in the iorder listed is given in parentheses. Thus, catalyst CPO-l in Table 1, for example, contains platinum and palladium in a weigllt ratio of one part platinum to one part palladium. In each case, the refractory metal oxide coating is alumina, predominantly comprising gamma alumina stabilized as indicated, the respective eight percents of stabilizer being indicated, the balance comprising substantia~ly alumina.

~1 ~LZ~567 TABLE I

Weight % and Alumina Sup~ort PG Metal PG M~tal Stabilizer in coating g/in Catalvst Component g/ft SupDort Co~ (% Stabilizer) ~ CPO-l Pt,Pd(l:l) 219 5% ceria 1.27 ! ' CPO-2 Pt,Pd(l:l) 186 5% ceria 1.64 CPO-3 Pt,Pd(1:4) 275 5% ceria 1.79 CPO-4 Pt,Pd(1:0 310 5% ceria 2.32 ! CPO-5( ) Pt,Pd(l:l) 200 5% ceria 1.26 10 ~ CPO-6 Pt,Pd,Rh (9.5:9.5:1) 230 5~ ceria 1.47 CPG-7 Pt,Pd(l:l) 186 2.5% lanthia 2.5% baria 1.64 ( ) TQRVEX alpha alumina monolith; all others are cordierite monoliths.

Preerred catalyst metals are platinum and palladium and combinations thereof, , preferably, combinations comprising about 10-90.~o by weight palladium, preferablv 25-75%, more preferably 60 to 40~ by weight palladium, and about 90 to 10% by weight platinum, preferably 75 to 25%, more preferably 40 to 60~ by weight platinum.
Generally, as the sulfur content of the hydrocarbon feed being treated in the first I catalyst zone increases, a higher proportion of the platinum to palladium is preferred.
i; On the other hand, for feeds which have a relatively high methane content, an increasing proportion of palladium is preferred.
The monolithic configuration of the catalytic partial o~iidntion catalyst of the~I first catalyst zone affords a relatively low pressure drop across it as compared to Il the packed bed of a particulate support catalyst. This is particulnrly important in view of the increase in gas volume occasioned by the reactions tnl;ing plnce in the I first catalyst zone. The total moles of product produced in the first catnlyst zone is higher than the total moles of H2O, oxidane gns and hydroc-lrlon feecl introduced therein. I`he individuRl gns flow pnssnges of the monolith nlso serve, in et`fect, ns .1 I
ll I

individual adiabatic chambers, 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, it serves as an ¦excellent heat source, radiating heat back towards the incoming gas stream which is pr~heated thereby, thus facilitating desired hydrocracking and oxidation reactions.

Steam Reforming Catalyst The steam reorming catalyst utilized in the second catalyst zone in accordance with the present invention may utilize a monolithic carrier as described above in ¦ connection with the partial oxidation catalyst or it may comprise a particulate Isupport 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"
I or the like means catalysts of regularly or irregularly shaped particles or shaped members or combinations thereof. A preferre d particulate support is alumina pellets or extrudate having a BFT (Brunnauer-Emmett -Teller) surface area of from about 10 to 200 s~uare meters per gram. Alumina or a lumina stabilized with rar earth metal ¦and/or alkaline earth metal oxides as desc ribed above, may be utilized as the pellets lor extrudate. An alumina particulate suppo rt stabilized with lanthanum and barium ¦ oxides as described above is pref erred.
I The catalytically active metals for the s team reforming catalyst may comprise ~any of the catalytic metals known for such purpose, for example, nickel, cobalt and 2~ ! mixtures thereof are well suited for use as steam reforming catalysts. Platinum group metals such as platinum and rhodium o r both may also be utilized for steam l reforming, as is known in the art. A preferred platinum group metal steam reforming ¦¦catalyst is a combination of platinum plus rhodium with the rhodium comprising from ~ . 1~

356~

¦about 10 to 90% by weight, preferably 20 to 40% by weight, of the total platinum Igroup metal present and the platinum comprises 90 to 10%, preferably 80 to 60%.¦The proportion of platinum and rhodium utilized will depend on the type of hydrocarbon ¦feed to be treated in the process. Other platinum group metals may also be utilized.
5IFor example, as disclosed in U.S. Patent 3,481,722, assigned to the assignee of this application, one or more of platinum, palladium, rhodium, iridium, osmium and ruthenium may be util;zed as the steam reforming catalyst.

Example 2 10(a) A barium nitrate solution is prepared by dissolving 159.9g Ba(NO3)2 in 1,650 ml of H2O. Lanthanum nitrate, in the amount of 264.9g La(NO3)2-6H2O is dissolvedin 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 thoroughly mixed in a sigma blade mixer for 30 minutes.
15~b) The impregnated alumina resulting from step (a) was extruded through 1/16"diameter dies so as to give 1/16" diameter extrudate in lengths from 1~4" to 3/8".
(c~ The extrudates from step (b) were dried at 110C for 16 hours and then calcined 2 hours at 1,050 C in air.
(d) A platium-rhodium solution was prepared by dissolving 42.0g Pt as H2Pt(OH)6 20in monoethanolamine and 18.0g Rh as Rh(NO3)-2H2O 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 obtainedin step (c) in a tumbling coater mixed for 30 minutes. The impregnated extrudate25is dried at 120C for 4 hours and then calcined for 30 mimltes at 500~ in air.
The resultant particulate steam reforming catalyst, designated SR-l, comprises 1.4 wt % platinum and 0.6 wt 96 rhodium on a La203 - BaO stabilized gamma alumina extrudate.

Il 12~567 ,i '.1 ¦ The catalysts of Examples 1 and 2 were utiiized in test runs. Before describing I these test runs, however, preferred embodiments of the apparatus of the present ¦ invention are described in some detail below.

¦ The Reactor Vessel Preferably, the reactor utilized in the autothermal reforming process of the invention comprises a fixed bed, adiabatic reactor. Figure 1 shows a somewhat schematic rendition of a preferred laboratory or pilot plant size reactor comprising a unitary vessel 1 within which a monolithic carrier partial oxid&tion catalyst 2 is disposed in flow communication via a passageway 3 with a bed of steam reforming catalysts 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 1 oxygen. The latter may be introduced as an oxygen containing gas, preferably air.

I The admixed reactants are introduced 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 withdrawn 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 j I schematic iliustration of Figure. 1.
In order to exemplify operation of the autothermal reforming process, test runs I¦ were carried out in an apparatus substantially in accordance with that schematically il illustrated in Figure 1, in which the monolithic carrier catalyst 2 was of cylindrical j configuration, three quarters of an inch in diameter and nine inches long. The steam l~ reforming bed was a cylindrical bed of particulate catalyst three inches in diameter 1¦ by nine and a quarter inches long. The following test rlms were carried out and Il the indicated results obtained. In operation, the reactants were preheated with thc ¦ oxidant stream being preheated separately from the hydrocarbon feed as a safety measure. After preheating, the streams were intimately mixed and immediately fed I

-2:~--I
I' .j into the partial oxidation catalyst 2 of vessel 1. Generally, all the oxygen present in the eed reacts within monolithic catalyst bed 2 to oxidize a poriton, but not all, of the hydrocarbon feed, resulting in an increase in temperature due to the exothermic oxidation reaction. At least some of the C5 and heavier hydrocarbon is hydrocracked in catalyst bed 2 to lighter, Cl to C4 hydrocarbon fractions. The heated, parti~ly 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 withdrawn via outlet B are cooled and unreacted water as well as any ¦ unreacted hydrocarbon feed is condensed and removed therefrom. The dry gas crmposi irn wes monitored by ga~ chr~maiography.

i l ~2~567 ¦ Example 3 (a) A monolithic oxidation catalyst made in accordance with Example 1 has ¦ the following composition:

i 186g of platinum group metal (PGM) per cubic foot of catalyst volume, the RGM c~mprising platinun~ and palladium in a 1:1 weight ratio. The PGM is distended upon a lanthia-barïa stabilized predominantly gamma alumina washcoat .~ dispersed on a Corning cordierite monolith 3/4 inch (1.9 cm~ in diameter and 9 inches (22.9 cm~ in length, and having ~00 gas flow passages per square inch of end face area. The monolith is loaded with 1.64g of washcoat per cubic ¦ inch of catalyst volume.
(b) A PGM steam reforming catalyst is provided by 1,075 ml of catalyst SR-l ¦ of Example 2, in a packed bed measuring 3 inches (7.62 cm~ in diameter and 9 1/4 I inches (23.5 cm) in depth.
¦ (c) The hydrocarbon feed in a No. 2 fuel oil having the following properties:
¦ API Gravity: 34.7 Distillation Range: 374-664F
Sulfur Content: 1200 parts per million (weight) hydrocarbon classes per ASTM D1319:
I Aromatics: 22.0%
Olefins: 5 7%
Saturates: 72.3%
(d) The reactor vessel is a fixed bed, adiabatic pressure vessel reactor of the type schematically illustrated in Figure 1. For safety considerations, the oxidant I stream, comprising oxygen enriched air, is preheated separately from the hydrocarbon ~ stream in a preheater (not shown in Figure 1). The steam is separated into two streams, one of which is blended with the oxidant stream and one with the hydrocarbon feed. The preheated streams are intimately mixed within a mixer, schematically ~ a ~ D t~ ft f~ ~

~ .
Ii 1 ~2~5~7 I .

illustrated in 9 in Figure 1, and the combined inlet stream at a pressure of 1 atmosphere and a preheat temperature of 1200F (649C), is immediately fed to ¦ the partial oxidation catalyst. The partial oxidation catalyst (2 in Figure 1) in the first catalyst zone contains the monolithic catalyst of (a), above, and the steam reforming catalyst (4 in Figure 1) in the second catalyst zone contains the SR-lcatalyst of (b), above.
¦ All of the oxygen contaîned in the inlet feed is completely reacted and a sufficient amount of the hydrocarbon is oxidized to heat the effluent reactant mass, by the exothermic catalytic oxidation reaction, to a temperature of 942C, high enough for steam reforming. The effluent from the catalytic partial oxidation catalyst 2 is immediately flowed into the steam reforming catalyst 4 and then withdrawn via the outlet opening as indicated by the arrow B in Figure 1, at an exit temperature of 1432F (778C). The volumetric throughput rate through the partial oxidation catalyst was 126,000 volumes of throughput at standard temperature and15 I pressure per volume of catalyst per hour and the volumetric throughput rate through ¦ the steam reforming catalyst (same basis) was 6500.
I The product gfls is cooled and unreacted water (and any unreacted hydrocarbon ¦ oil) is condensed therefrom. The dry gas composition is monitored by gas I chromatography and the following results were measured and the results tabulated in Table II. In this illustration air was used instead of oxygen-enriched air. The result of this is that the "equivalent hydrogen", i.e., (CO + H2) to N2 ratio is low, i.e., 1.53/l instead of 3/l.

~ 567 I

TABLE lI

A. Effluent Composition and Hydrocrrb:)n Conversion ¦ Conversion of Dried Product Gas Composition, Mole 96 I l)ried Product Gas As l~qeasured ¦ H 2 41.11 CO 11.52.
C2 12.51 N2 34.37 CH4 0.26 C2 6 0.01 Hydrocarbon ~onversion 98.8%( ) ______________________________ I
¦ ( )weight percent of quantity of hydrocarbon feed to the reactor. I
I I
The foregoing Example 3 shows the efficacy of the process for substantially complete (98.8%) conversion of a No. 2 fuel oil, which is a difficult feedstock to convert by conventional steam reforming. The process of the invention can of course readily handle lighter hydrocarbon feeds, which are easier to reform, and use higher 2 to ~ ratios.
2~ Referring now to Figure 2, there is shown a schematic illustration of an ammonia synthesis plant which includes a synthesis gar making section in accordance with one embodiment of the present invention. A source of a hydrocarbonaceous feed, in this case a hydrocarbon feed is introduced via line 10 and hydrocarbon feed compressor 12 for passage through a heater 14 and thence to a mi~er 16 for admi~ture with steam and oxygen-enriched air as described below. Hydrocarbon feed compressor 12 compresses the hydrocarbon feed to the elevated pressure at which the autothermal reforming operation is to be carried out. Heater 14 may be of any conventional design and would include a burner means (not shown) for combllsting a fuel therein to providc preheating by indirect heat trnnsfer to the strenms passing therethrollgh.
An nir compressor 18 is supplied with ntmospheric nir vin a~n inlct line 2n ~Indll I i2~567 compresses the air to the pressure at which the autothermal reforming operation is I to be carried out. A portion of the compressed air is passed via lines 22, 24 around ¦ an air separation zone 28 to a line 30 for passage to a heater 14 for preheating.
I Another portion of compressed air from air compressor 18 is passed through line 26 ¦¦ to air separation zone 28 wherein oxygen is separated from the compressed air by I¦ any suitable means known to those skilled in the art. The resultant oxygen stream i is fed via line 30, into which the air frorm line 24 is introduced, for passage of the ~¦ thus oxygen-enriched air via line 30 through heater 14 for preheating and thence to ¦ mixer 16. In air separation zone 28, a nitrogen stream 29 which also contains rare ¦ gases, including argon, found in air is removed from the process and is normally recovered as a nitrogen gas by-product. The argon and other inert gases which are ¦ removed from the process in air separation zone 28 advantageously reduce the total I inerts which will ultimately find their way into the ammonia synthesis loop via the ¦ synthesis gas and therefore reduces the amount of purge gas which must be withdrawn ¦ from the loop.
¦ Air separation plant 28 may use any suitable type of air separation process including, for example, a cryogenic separation process, a membrane diffusion process, or a pressure-swing absorption process utilizing inorganic absorbants or carbon molecular sieves. Although not shown on the flow sheet of Fiigure 2, heat available ~ from the autothermal reforming process (shown in Figure 2 as being utilized only in ¦ heat exchanger 363 may also be utilized in the air separation zone 28 if an air Il separation process is utilized which requires a heat input.
Il As will be noted from Figure 2, it is seen that complete separation of a pure ¦ oxygen stream from the air is not necessary, but only that an oxygen enriched stream must be obtained to enrich the oxygen content of the a;r fed to autothermal reformer 42. As illustrated in the schematic diagram of Figure 2l line 29 carries nitrogen~
or at least a nitrogen-enriched streaml away from the air separation plant 28. I~
may also contain a substantial proportion, if not all, of the argon, which comprises l ll 1.

S67 ~

!

about one percent of atmospheric air, which enters zone 28.
Make up water is introduced via line 32 and boiler feed water pump 34 through heat exchanger 36 wherein the water, together with condensate water recycled from l a subsequent point in the process as described below, is heated and steam is generated ! therefrom. The steam generated in heat exchanger 36 is passed via line 38 to mixer 116. A mixture of oxygen enriched air, steam and hydrocarbon feed obtained in ¦ mixer 16 is introduced via inlet line 40 into autothermal reformer 42.
In reformer 42, the inlet stream mixture o hydrocarbon feed, steam and l oxygen-enriched air is passed initially through a catalytic partial oxidation catalyst ~ supported on a monolithic honeycomb carrier disposed within the neck portion 42a of reformer 42. The effluent from the first catalyst zone passes into the second~ catalyst zone comprising a platinum group metal steam reforming catalyst contained ¦¦ within main body portion 42b of autothermal reformer 42. Generally, as mentioned ¦l above, a portion, less than all, of the hydrocarbon feed content of the inlet stream 1! is catalytically oxidi~ed within the first catalyst zone. If C5 or heavier hydrocarbons ¦ form a part of the feed, they are hydrocracked under the conditions prevailing in the first catalyst zone, to lighter, Cl to C4 constituents. The steam reforming 1~ reaction carried out in the second catalyst zone reacts H2O with unoxidized ¦ hydrocarbons to form hydrogen and carbon monoxide. Generally, the inlet streaml~ components (hydrocarbon, oxygen-enriched air, and steam) react in autothermal! reformer 42 to produce a mixture containing H2, CO, C02, N2, H20 and a small amount of residual methane.
Effluent from autothermal reformer 42 is passed via line 44 through heat Il exchanger 36 as mentioned above. The temperature of the effluent in line 44 is ¦ sufficiently high so that superheated steam may be effectively generated in heat exchanger 36. After the heat exchange in exchanger 36, the autothermal reformingeffluent is passed via line 46 to a high temperature shif t conversion zone 50. A
ll l~

567 ~--side stream of steam may be passed via line 48 to increase the proportion of steam in the effluent entering high temperature shift conversion zone 50 to improve the reaction conditions for the shift conversion.
Typically, the temperature of the effluent in line 44 from autothermal reformer 1 42 will be at a temperature of from 1600F to 1900F (871C to 1038C) and will ¦ be cooled in heat exchanger 36 to a temperature of about 800F to 900F (427Cj to 482C), which is a temperature suitable for shift conversion as described below.
Shift converters are conventionally employed in conjunction with steam ¦¦ reforming operations. In steam reforming, the hydrocarbon reacts with H2O to yield ¦1 a product gas containing primarily hydrogen and carbon monoxide, plus any unreacted ¦ hydrocarbons. In order to reduce the carbon monoxide level and enhance the hydrogen ¦ yield, the effluent of the steam reforming process may be passed into a so-called shift converter, in which the effluent is contacted with a catalyst of the knowntype over which carbon monoxide will react with H20 to yield carbon dioxide and ! hydrogen according to the following reaction: !

CO + H20 = C2 H2 This water gas shift reaction is often carried out in two stages, a first high I temperature stage, e.g., 800-900F (427-482C) in order to secure high reaction rates and a second, low temperature stage, e.g., 700-750F (371-399C) to enhance the overall yield of the reaction. Accordingly, as shown in Figure 2, the effluent from reformer 42 is passed via lines 44, 46 into high temperature shift reactor zone 50 in which it is contacted with a suitable catalyst to carry out the shift reaction.
Upon emerging from zone 50, the shift-reacted effluent may be passed via line 52to a sulfur treatment zone 54, which may comprise any suitable equipment for removing or reducing the sulfur content of the gas stream, such as the known iron ~xide or zinc oxide dry removal processes, or a known ~let method for the removrl ~ I
l l l l I
l l of sulfur and sulfur compounds. The treated effluent from the sulfur treatment zone 54 is passed via line 56 to a low temperature shift reactor zone 58 in which a second, low temperature shift reaction is carried out to convert carbon monoxide ¦I to carbon dioxide and hydrogen. The resultant hydrogen and nitrogen-rich gas stream, 51¦ containing primarily N2, H~, CO2 and H2O passes from low temperature shift converter zone 58 via line 60 for introduction into a selective oxidation zone 70.
~¦ Prior to passing to selective oxidation zone 70, the effluent from the low temperature shift reactor zone 58 is cooled in cooling and separation zone 62 and ¦¦ water is condensed therefrom. The resulting condensate water may be passed via 101 line 66 together with make up water from boiler feed water pump 34, to heat exchanger 36. The gases are cooled in cooling zone 62 to a temperature suitable for the selective oxidation reaction to be carried out in selective oxidation zone 70, to which the cooled effluents are passed via line 60. Additional oxygen is introduced !into the effluent entering selective oxidation zone 70, and this may be accomplished 15~by taking a side stream of compressed -air via line 72 and combining it with the Icooled effluent in line 60 for passage into selective oxidation zone 70. The amounts ¦ of supplemental air introduced via line 72 is calculated both to supply sufficient ¦loxygen for the selective oxidation process, and to provide a molar ratio of nitrogen l¦to hydrogen in the product synthesis gas of approximately 3:1. Selective oxidation 20llzone 70 rnay be any suitable selective oxidation process in which residual carbon !monoxide contained in the effluent is contacted with a catalyst to selectively oxidize ¦Ithe carbon monoxide to carbon dioxide in the presence of hydrogen. A highly efficient catalyst for the purpose is sold under the trademark SELECTOXO by I¦Engelhard Corporation. A selective oxidation process utilizable in an ammonia 25¦Imanufacturing operation is shown in IJ.S. Patent 4,238,468, issued December 9, 1980 to Bonacci et al, and assigned to Engelhard Corporation. After the selective oxidation treatment, the treated effluent is passed via line 74 to a carbon dioxide removal ~Z11~567 ¦¦ zone 76 wherein the residual carbon dioxide content of the effluent stream is further reduced by any suitable, known process. Remov~l or reduction to extremely low vfllues of the carbon dioxide in the synthesis gas is advantageous in that carbon dioxide, under specific temperature and pressure conditions and concentrations, can react with ammonia to form carbamates~ and this is of course undesirable in the ammonia synthesis process.
The selective oxidation process is thus utilized for removal or reduction of I trace quantities of carbon oxide and is desirable as minimizing the amount of methane ¦ formed, as discussed below, in methanation zone 82. The formation of such methane consumes product hydrogen and it is therefore preferable to reduce the carbon monoxide content as much as feasible in zone 70.
The carbon dioxide reduced effluent is passed via line 78 through a second cooling and separation zone 80 in which additional water is condensed therefrom and ¦ the water condensate may be recycled via lines 68, 66 to heat exchanger 36 for1¦ return to the process as steam.
¦ The cooled gas is then passed through a methanation zone 82 in which residual carbon oxides are contacted over a catalyst in a methanation step in which hydrogen reacts with CO and CO2 to produce methane and water. It will be appreciated that the residual amounts of carbon oxides available to react in the methanation zoneare very sma~l. The resulting methane is an inert in the ammonia synthesis process and tends to build up therein. Accordingly, the m ethane, along with other inerts, must be purged from the ammonia synthesis loop. But this is preferable to the ¦ adverse effect in the ammonia synthesis loop of either carbon monoxide, which is an ammonia synthesis catalyst poison, or carbon dioxide, which may react with ammonia as mentioned above, to form explosively decomposable solid carbamates.
The effluent from methanation zone 82 is withdrawn therefrom as product synthesis gas and may be passed through line 84 to an ammonia synthesis process. I

lZi~iS67 ¦ Figure 2 illustrates the so-called ammonia synthesis loop, in which the product ¦ synthesis gas from line 84 is passed via line 86 to a synthesis gas compressor 88 I wherein it is compressed to an elevated pressure, say 1,000 to 15,000 psi and at a ¦¦ relatively low temperature, from about 200-600C, suitable for ammonia synthesis.
5 tl The compressed synthesis gas is passed via line 90 to an ammonia synthesis reactor ¦~ 92 containing therein a catalyst suitable to react nitrogen with hydrogen to form ammonia. A small proportion only of the nitrogen and hydrogen in the synthesis gas is converted to ammonia in a single pass through the catalyst of ammonia synthesis reactor 92, and the partially reacted gas is passed through a cooling and n ¦ separating zone 94 within which the effluent is cooled sufficiently to condense ¦l ammonia as a liquid therefrom. The ammonia synthesis catalyst is composed of iron ¦¦ oxides (Fe~O3 and FeO) that have been triply promoted with K2O, SiO2, and A12O3.
l Various shapes and forms of this ammonia synthesis catalyst or any other ammonia I¦ synthesis catalyst can be used. A typical composition is: Fe2O3 64-66 wt %; FeO
15¦1 29-31 wt %; A12O3 2-3 wt %; SiO2 0-0.8 wt %; K2O 0-2 wt %. The ammonia is removed via line 96. The remaining synthesis gas is recycled via line 98, in which ¦¦ it is supplemented with fresh synthesis gas from line 84, and recycled through the ammonia synthesis loop. A purge line 100 removes a proportion of the gas circulating ¦ in the ammonia synthesis loop in order to control the build-up of inert gases in the 20¦¦ loop. The purge gas removed from line 100 may, as is known to those skil!ed in j the art, be treated to have removed therefrom argon, ammonia, methane, CO2 and¦ other impurities and be returned as synthesis gas to the process.
¦ Generally, the key operating parameters for autothermal reforming to generate Il an ammonia synthesis gas are the preheat or inlet temperature (at line ~0 of Figure 2511 2), the pressure within autothermal reformer 42 and in the inlet stream (line ~0 of ¦ Figure 2), the oxygen to carbon ratio, the H2O to carbon ratio, and the N2 to cnrbon ~ rrtio. The reference to ~rrbon is of course the crrbon content of the ll -32- ~

lZ1~567 !
~ I
hyàrocarbon feed and all such ratios are expressed as moles of 2 or H20 to atoms of carbon~ Since the process normally convert~s carbon oxides to hydrogen as part ¦ of the synthesis gas preparation, for ammonia synthesis it is desired that the molar ¦ ratio ~CO + H2)/N2 exiting the autothermal reformer be approximately 3:1.
! By utilizing the autothermal reforming process as described, a wide variety ofhydrocarbonaceous feeds may be utilized and efficiently and economically converted ¦into a nitrogen and hydrogen containing synthesis gas. In addition to petroleum and coal derived hydrocarbons, as m entioned above, biomass-derived feeds providing ,carbon-containing compo~mds such as methane and H, O and N in compounds or as elements may be used. Such feeds may also contain sulfur or sulfur compounds. If the sulfur content of the feed is sufficiently high, the sulfur treatment zone as illustrated in Figure 2 will normally be required. If oxygen and nitrogen are present in the biomass-derived feed in appreciable amounts, they can provide at least a Ip~rtionof the "air" for the process, and the atmospheric air introduced to the process ¦can be reduced accordingly.
The following examples 4 and 5 show typical operating conditions for the autothermal reformer utilized in accordance with the present invention. As indicated by the examples, generally, the H20 to C ratio is increased with increasing pressure ~n order o evoid an excessive increrse of the methene content in ~he synthesis grs.

~ ' I1 121~567 1l~ . I
I !I _ I

Pressure 35 Atmospheres 67~7 Atmospheres (525 psia) (1015 psia) Inlet Temperature 1200F (649C) 1200F (649C) Inlet Stream (line 40 in Figo 2) (Lb.-Moles/Hr.) (~lole %) (Mole %) CH4 1.000 (20~0) L000 (18~2) ¦! H20 2~500 (50~0) 3~000 (54~5) 2 1)~5722(11~4) 0~6012 (10~9) N2 0~9300 (18~6) 0~9060 (16~4) Exit Stream (line 44 ¦ in Fig~ 2) (Lb.-Moles/Hr.) n ! CH4 0~0185 (0~3) 0~0199 (0~3) H2O 2~275 (35~5) 2~8067 (40~9) i j 2 ~0~ _0_ ~¦ 2 0~9300 (14~5) 0~9060 (13~2) H2 2~1930 (34~3) 2~1535 (31~4) !, co 0~5991 ~9~4) 0~5646 (8~2) 2 Q~3851 (6~0) 0~4156 (6~0) i I~Exit Temperature 1775F (6968C) 1840F (1004C) ,I While the invention has been described in detail with respect to specific ¦I preferre~d embodiments thereof, it will be appreciated that those skiUed in the art, ¦l upon a reading and understanding of the foregoing, will readily envision modifications ¦¦ and variations to the preferred embodiments which are nonetheless within the spirit ¦ and scope of the invcntion and of the claims.

-3~-

Claims (27)

WHAT IS CLAIMED IS:

1. A process for the production of ammonia from a hydrocarbonaceous feed, comprising the steps of:
(a) preheating an inlet stream comprising a hydrocarbonaceous feed, H2O, air and oxygen to a preheat temperature sufficiently high to initiate catalytic oxidation of said feed as defined below;
(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 palladium and platinum catalytic components dispersed therein, the amounts of feed, H2O, and oxygen introduced into said first catalyst zone being controlled to maintain in said inlet stream an H2O to C ratio of from about 0.5 to 5 and an O2 to C ratio of from about 0.35 to 0.65;
(c) contacting the preheated inlet stream within said first catalyst zone with said palladium and platinum catalytic components to initiate and sustain therein catalytic oxidation of said feed to produce hydrogen and carbon oxides therefrom,the temperature of at least a portion of said monolithic body being at least about 250°F
(139°C)greater than the ignition temperature of said inlet stream, and oxidizing in said first catalyst zone a quantity, less than all, of said feed, which quantity is sufficient to heat such first zone effluent to an elevated temperature high enough to catalytically steam reform, within a second catalyst zone defined below, hydrocarbon remaining in such first zone effluent without supplying external heat thereto;
(d) passing the first zone effluent, while still at an elevated temperature, from aid first catalyst zone to a second catalyst zone containing a platinum group metal team reforming catalyst therein, and contacting the first zone effluent in said second catalyst zone with said reforming catalyst to react hydrocarbons therein with H2O
to produce hydrogen and carbon oxides therefrom;
(e) withdrawing the effluent of said second catalyst zone as a hydrogen-containing synthesis gas, and removing heat therefrom to cool said synthesis gas;

(f) reacting carbon monoxide in said synthesis gas with H2O to produce hydrogen;
(g) removing sulfur containing compounds and H2O from said synthesis gas;
(h) passing said synthesis gas into an ammonia synthesis loop to react the hydrogen with nitrogen thereof over an ammonia synthesis catalyst at ammonia synthesis conditions; and (i) withdrawing ammonia as product from said ammonia synthesis loop.

2. The process of claim 1 wherein the proportion of air to oxygen in said inlet stream is such as to provide oxygen enrichment of the air in said inlet stream to at least about 33 volume percent oxygen.

3. The process of claim 1 or claim 2 wherein said hydrocarbonaceous feed is a hydrocarbon feed.

4. The process of claim 2 wherein the preheat temperature is from about 800°F to 1400°F (427°C to 760°C).

5. The process of claim 4 wherein said first catalyst zone is maintained at a temperature of from about 1750°F to 2400°F (954°C to 1316°C) and the first zone effluent is introduced into said second catalyst zone at substantially the same temperature.

6. The process of claim 4 wherein a volumetric hourly rate of at least 100,000 volumes of throughput per volume of catalyst is maintained in said first catalyst zone.

7. The process of claim 6 wherein 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.

8. The process of claim 4 wherein said first catalyst zone comprises palladium, platinum, and, optionally, rhodium catalytic components distended upon a refractory metal oxide support layer carried on said monolithic body.

9. The process of claim 8 wherein said catalytic component of said first catalyst zone comprises, on an elemental metal basis, about 10 to 90% by weight palladium and about 90 to 10% by weight platinum.

10. The process of claim 9 wherein said catalytic component of said first catalyst zone comprises about 25 to 75% by weight palladium, and about 75 to 25%
by weight platinum.

11. The process of claim 10 wherein said catalytic component of said first catalyst zone comprises from about 60 to about 40% by weight platinum and from about 40% to about 60% by weight palladium.

12. The process of claim 8 wherein said steam reforming catalyst comprises one or both of platinum and rhodium catalytic components.

13. The process of claim 8 wherein said steam reforming catalyst comprises, an elemental metal basis from about 10 to 90% rhodium and from about 90 to 10%
by weight platinum.

14. The process of claim 13 wherein said steam reforming catalyst comprises from about 20 to 40% by weight rhodium and 80 to 60% by weight platinum.

15. The process of claim 2 wherein said feed is a hydrocarbon and including the additional step of treating the synthesis gas withdrawn from said second catalyst zone to convert carbon monoxide therein to carbon dioxide and then removing carbon dioxide from the synthesis gas to provide a carbon oxides-depleted ammonia synthesis gas.

16. The process of claim 15 wherein the step of converting carbon monoxide to carbon dioxide comprises adding oxygen to the synthesis gas withdrawn from said second catalyst zone and contacting the resulting mixture with a catalyst effective for the selective oxidation of carbon monoxide to carbon dioxide in the presence of hydrogen.

17. The process of claim 16 wherein the oxygen added to said synthesis gas is added as air, and the amount of added air is calculated to bring the nitrogen content of the synthesis gas to from about 2.9 to 3.1:1 molar ratio of hydrogen to nitrogen.

18. The process of claim 2 wherein the amounts of feed, H2O, air and oxygen introduced into said first catalyst zone are controlled to maintain in said inlet stream an H2O to C ratio of from about 1 to 4, and an O2 to C ratio of from about 0.5 to 0.6.

19. The process of claim 2 wherein the amounts of feed, H2O, air and oxygen introduced into said first catalyst zone are controlled to provide oxygen enrichment of the air in said inlet stream to about 33 to 50 volume percent oxygen.

20. The process of claim 2 carried out at a pressure of from about 100 to 1500 psia.

21. A process for the production of ammonia from a hydrocarbon feed, comprising the steps of:
(a) preheating an inlet stream comprising a hydrocarbon feed, H2O, air and oxygen to a preheat temperature of about 800°F to 1400°F (427°C to 760°C) and introducing the preheated inlet stream at a pressure of from about 100 to 1500 psia 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 palladium and platinum catalytic components dispersed therein, the amounts of hydrocarbonaceous feed, H2O and oxygen introduced into said first catalyst zone being controlled to maintain in said inlet stream an H2O to C ratio of from about 0.5 to 5, and an O2 to C ratio of from about 0.4 to 0.65;
(b) 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 said hydrocarbonaceous feed to produce hydrogen and carbon oxides therefrom and oxidizing in said first catalyst zone a quantity, less than all, of said hydrocarbonaceous feed, which quantity is sufficient to heat such first zone effluent to an elevated temperature of from about 1750°F to 2400°F (954°C to 1316°C);
(c) passing the first zone effluent, while still at said elevated temperature, 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 an hourly volumetric rate of from about 2,000 to 20,000 volumes of throughput per volume of catalyst to react hydrocarbons therein with H2O to produce hydrogen and carbon oxides therefrom; and (d) withdrawing the effluent of said second catalyst zone as a hydrogen-containing synthesis gas, and removing heat therefrom to cool said synthesis gas;
(e) withdrawing the effluent of said second catalyst zone as a hydrogen-containing synthesis gas, and removing heat therefrom to cool said synthesis gas;
(f) reacting carbon monoxide in said synthesis with H2O to produce hydrogen;
(g) removing sulfur containing compounds and H2O from said synthesis gas;
(h) passing said synthesis gas into an ammonia synthesis loop to react the hydrogen with nitrogen thereof over an ammonia synthesis catalyst at ammonia synthesis conditions.

22. The process of claim 21 wherein the proportion of air to oxygen is such as to provide oxygen enrichment of the air in said inlet stream from about 33 to 50 volume percent oxygen, whereby to provide a nitrogen-containing hydrogen-rich synthesis gas .

23. The process of claim 22 further including converting carbon monoxide in the effluent of said second catalyst zone to carbon dioxide and then removing carbon dioxide from said effluent of said second catalyst zone.

24. The process of claim 23 wherein the effluent withdrawn from said second catalyst zone is further treated to remove sulfur and sulfur compounds therefrom.

25. The process of claim 21 or claim 22 wherein said platinum group metal catalyst of said first catalyst zone comprises palladium, platinum, and, optionally, rhodium catalytic components and said steam reforming catalyst comprises platinum and rhodium catalytic components.

26. The process of claim 22 wherein said catalytic components of said first catalyst zone comprise, on an elemental metal basis, about 10 to 90% by weight palladium, 90 to 10% by weight platinum and 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.

27. The process of claim 26 wherein said catalytic components of said first catalyst zone comprise about 25 to 75% by weight palladium, 75 to 25% by weight platinum and said catalytic components of said second catalyst comprise about 20 to 40% by weight rhodium, and 80 to 60% by weight platinum.