US2360463A - Hydrocarbon conversion - Google Patents

Description

Oct- 17" 1944. M. H. ARvEsoN HYDROCARBON CONVERSION Filed April 24, 1941 2 Sheets-Sheet 1 INN @NWN \\\N m.\N

Oct. 17, 1944. M. H. ARvEsoN HYDROCARBON CONVERSION Filed April 24, 1941 2 Sheets-Sheet 2 NNY @NN QNX..

b, www WSW Patented Oct. 17, 1944 HYDROCARBON CONVERSION Maurice H. Arveson, Flossmoor, Ill., assignor to Standard Oil Company, Chicago, Ill., a corporation of Indiana Application April 24, 1941, Serial No. 390,185

7 Claims.

This invention relates to the conversion of hydrocarbons into useful products and particularly into high quality gasolines and still more particularly aviation gasolines.l The invention relates especially to the conversion of methane into heavier high quality hydrocarbon products.

One of the outstanding problems in the oil and gasindustry is the efficient and economic utilization of methane. In the past methane has been utilized only for fuel value and in fact has often been completely wasted. Processes have heretofore been known for the conversion of methane into heavier hydrocarbons and the conversion of methane to carbon monoxide and hydrogen followed by the Fischer synthesis is an outstanding instance of this. In the Fischer process carbon monoxide and hydrogen are converted into liquid, gaseous and solid hydrocarbons. However, this Fischer synthesis is subject to very serious disadvantages since the products produced by it constitute very'inferior motor fuels and the process produces large quantities of wax and other fractions which are very diicult to utilize.

I have provided a. combination process in which methane taken from a gas well or from a high pressure well of the so-called distillate or condensate type canA be processed by steps includinto products rich in branched chain and/or aromatic hydrocarbons.

A still further object of my invention is to utilize my various processing steps in a way which makes possible the utilization of low grade byproducts from one step for the production of high quality motor fuels or motor fuel components in another step.

Other and more detailed objects, advantages and uses of my invention will become apparent as the description thereof proceeds.

My invention will now be described in detail with particular reference to the accompanying flow diagram which illustrates various embodiments of my process. It will be understood that this ilow diagram is in highly simplified form and that modifications of it can and will be made by those skilled in the-art. In particular I" have l omitted various features such as flow and teming the Fischer synthesis tand other `operations to produce a highly superior product in very high yields. Moreover, I have combined the various steps of this process in certain new and advantageous manners as will subsequently become apparent, so that each of them contributes to the improved utility and eilciency of the other. Furthermore, I have combined the processing of hydrocarbons resulting from methane with the processing of hydrocarbons heavier than methane and particularly three and four carbon atom vide for the cooperative processing of methane and heavier hydrocarbons in such manner that the processing of each facilitates the processing of the other.

It is also an object of my invention to convert hydrocarbons produced in the Fischer process perature control. Furthermore, while this viiow diagram illustrates a number of alternatives it will be understood that these alternatives will not all be used in connection with a single installation but that they have been combined into a singlek flow diagram in order to avoid needless multiplication of the drawings.

Turning now to the drawings in more detail. the hydrocarbons used in my process can suitably come from a. Well l which can be na gas well or, preferably, a well of the so-called distillate or condensate type. Since my invention in certain of its forms relates primarily to the processing of the products resulting from the Fischer synthesis, it will be apparent that the hydrocarbons used to supply the synthesis gas can come from -still other sources such as refinery gas or gas produced along with crude oil. In fact, while I prefer to make the Fischer synthesis gas, consisting of carbon monoxide and hydrogen, from hydrocarbon sources and particularly from methane produced from oil or gas wells, this synthesis gas can be made from coal and Water or in other manners known to the art.

Returning to well I, it will be understood that while a single well is shown, a number of wells will usually be employed and in fact my process can be applied vmost economically to a large gas or distillate field including a considerabe number of wells. If well lis a gas well, the gas is passed through valves 2, 3 and 4, and lines V5,- t, l and 8 to a hydrogen sulfide removal plant '9 which is shown merely as a rectangle since this sulfur removal part of the process can be carried out in various Ways and is conventional. It should: be mentioned, however, that one suitable method of removing the hydrogen sulfide is to wash it with an alkaline purification medium and the ethanolamines including the mono-, diand tri-ethanolamine are suitable for this purpose.

While my process can thus be applied to methane, usually containing small amounts of heavier hydrocarbons, produced from a gas well or to similar of! gases from a natural gasoline plant, itis particularly advantageous as applied to wells or fields of the so-called distillate or condensate type.

These distillate fields have become increasingly common in recent years and are typied by high l reservoir pressures within the retrograde condensation range with the result that the hydrocarbons present in such a reservoir are wholly or largely in a single phase which can be referred to as a vapor phase or supercritical phase. When the pressure and temperature of such iiuids are reduced in the courserof their passage upward in a well, a portion of the heavier hydrocarbons is usually precipitated as a result oi' the cooling and particularly as a result of the retrograde condensation resulting from the pressure drop. Further quantities of valuable hydrocarbons including propanes and butanes and particularly gasoline range hydrocarbons are then recovered together usually with some slight amount of heavier hydrocarbons by a distillate recovery process.

This distillate recovery process can be of the retrograde condensation type in which the recovery is effectuated by further cooling the well fluids and reducing their pressure to a point within the retrograde condensation range, for instance to a pressure ci 700 to 1500 pounds per square inch. The pressure chosen will depend on the composition of the well uids and the economics of the particular situation.

It is generally preferred, however, to recover these so-called distillate or condensate hydrocarbans by a high pressure absorption process in which the well fluids are contacted concurrently or countercurrently with an absorption medium at a high pressure such as 1000 to 3000 pounds per square inch, for instance 1500 pounds per square inch. The absorption medium chosen can be a portion o! the heavy ends of the distillate itself but superior results can be obtained by the utilization of other types of absorption media and the relatively heavy, predominantly aromatic hydrocarbons resulting from processes such as catalytic aromatization have been found to be particularly suitable.

Assuming that well I is of the distillate type, the well fluids therefrom, after separation, if so desired, of the readily separated liquids are passed to an absorber I0. Since pressure reduction is usually required, particularly when the well operates at a very high pressure such as 4000 or 5000 pounds per square inch, these well iiuids can advantageously be passed through an expansion engine II with a resultant increase of the degree of cooling and with recovery of power which can be utilized to supply part of the energy necessary to operate compressors I2 in connection with the plant itself or can be utilized in other ways.

In case the well fluids contain an undesirable quantity of moisture, these iluids can be passed through one or more driers I3 by closing valve 3 and opening valves I4 and I5, thus avoiding the formation of natural gas hydrates which would otherwise interfere in some instances with the operation of expansion engine II and/or absorber I0. These driers can be of various types but simple calcium chloride driers are suitable. Provision can, of course, be made for regenerating the calcium chloride but in the interests of simplicity this has not been shown. A separator between Well I and absorber I0 is also often desirable to remove Water and liquid hydrocarbons.

In any event all or most of the well uids pass vthrough valve I6 into absorber I0', valve 4 being closed. In absorber I 0 the well uids contact an absorption medium which is introduced from line I1 through heat exchanger I8 and cooler IS into the top of absorber tower I0 in which the absorption medium passes downward countercurrent to the rising well fluids and serves to absorb from the well iluids the desired normally liquid hydrocarbons together with substantial quantities of propane and butanes and also some lighter hydrocarbons. Absorption plants of this type are sometimes complicated and provision is often made for mist recovery and the like but again in the interests of. simplicity the process has been shown in its elemental form.

From the base of absorber tower I0 the rich oil is passed by means of line 20 through heat exchanger I8 wherein its temperature is raised and then into stripper tower ZI. This stripper is heated at the base by means of coil 224 and the lean oil is removed from the base of the stripper and passed through valve 23, line- I1, heat exchanger IB and cooler I9 into absorber tower I0 by pump 24. Any excess absorption medium is1withdrawn through valve 25 and line 28 to tank 2 Returning to absorber tower Ill, part of the o1! gases, predominantly methane, pass through line 28, valve 29 and line 30 to compressor I2 and are reiniected into a subsurface formation through line 3i, valve 32 and well 33. A plurality of input wells can and usually will be used. This reinjection of gas serves an important purpose in increasing the ultimate recovery of valuable.

hydrocarbons from, the reservoir and in preventing retrograde condensation within the reservoir. ,The remainder of the predominantly methane,- ceous gas from absorber III passes through valve 34 to hydrogen sulde removal plant 8 previously described. Since the hydrogen sulfide removal operation normally operates at moderate pressures, these gases can suitably be passed through expansion engine 35 to recover power which can be utilized as described in the case of expansion engine I I.

An additional source of gases for the hydrogen sulde removal plant is obtained by passing the overhead from stripper 2| through condenser 3l to a separating drum 31. This condenser is operatedl at such temperature as to condense substantially all of the butane and heavier hydrocarbons together with at least a substantial part of the propane coming overhead from the stripper. The uncondensed gases pass through valve 38 and lines 1 and 8 to hydrogen sulfide removal plant 9.

The liquid fraction from separating drum 31 passes to debutanizer tower 39 through valved pipe 40. The debutanizer tower is heated at the base by coil 4I and provided with reflux by means of condenser 42, separating drum 43 and pump 44 which takes a portion of the liquid phase from separating drum 43 and reintroduces it into the top of debutanizer tower 39 through valve 45 and line 46. The remainder of the liquid phase from separating drum 43 together with the vapor phase from this same drum pass to a stabilization tower 41 through valved lines 48 and 49. Tower 41 is heated at the base by coil 50 and p'rovided with reflux by means of condenser separating drum 52 and pump 53. The fixed gas from the last 5 mentioned separating drum passes through valve 54 and lines 55, 1 and 8 to the hydrogen sulde removal plant 9. The liquid phase from this ,separating drum is largely propane and passes through valved line 56 to storage tank 51. 10 Reverting to debutanizer 39, a sidestream composed predominantly of five and six carbon atom hydrocarbons can suitably be removed by means of trapout plate 58 and passed by means of pump 59 through line 90, valved line 6I and line 62 to a so-called paraffin isomerizaton process-which will later be described. The bottom from the debutanizer tower consisting largely of the heptane and heavier fraction ofthe distillate is pref-v erably passed by means of pump 63 through valved line 64 to a catalytic aromatization operation which lwill likewise be described subsequently.

I have referred to various sources of hydrocarbon gas, consisting predominantly of methane, passing to hydrogen sulfide removal plant 9. It will, of course, be understood that it is not necessary to use hydrocarbons from all these sources and it will likewise be understood that hydrocarbons from still other sources can be introduced at this point. Furthermore, any available sources of methanaceous gases which are free of sulfur may be introduced directly into the Fischer process synthesis gas manufacturing operation rather than passing through the hydrogen sulfide removal plant but it is necessary to make a very thorough clean-up of sulfur since it tends to poison the catalysts normally used in the manufacture of synthesis gas for the Fischer process.

From the hydrogen sulfide removal plant the puriiied gases pass through lines 65 and 68 to a reformer furnace 51. In line 66 they are joined with other hydrocarbon streams and also with steam introduced through valved line 68. The hydrocarbons and steamin'line 58 are preferably 4r, preheated by closing valve 89 and opening valvesA 1|i and 1|, thus passing them through apreheater 12 wherein they pick up heat from the. synthesis Sas passing out of the furnace 81.

In furnace 81 the feed gases pass through coils 13 wherein they are heated to a temperature in the neighborhood of' 1500 F. and substantially atmospheric pressure and then pass to catalyst tubes contained in reactor 14 at substantially this same temperature and pressure.

Reformer furnace 51 can be heated in various ways but the suitable sources of fuel include a portion of the hydrocarbon stream which would otherwise pass to coil 13. Thus valve 18 can be opened and fuel supplied to burner 11 through line 18. Alternatively, or in addition, fuel can be supplied from line 19 through line 80 and valved line 8|.

The reaction between methane and steam theoreticafly produces one mol of carbon monoxide to three mols of hydrogen and in practice the production is very similar. The presence of heavier hydrocarbons increases somewhat the mol ratio of carbon monoxide to hydrogen but in any event this mol ratio is lower than that required in the conversion of this gas mixture into heavier hydrocarbons. This situation can be remedied in various ways but I prefer to rectify the matter by removing Acarbon dioxide from ue gas, preferably made from sulfur-free fuel and introducing this into reformer furnace 61 where it is reduced to carbon monoxide.

Thus flue gas from flue 82 of reformer furnace 61 can be passed through lines 83 and 84 into absorber tower 85 in which carbon dioxide from these iiue gases is absorbed by passing an absorption medium such as a. dilute alkaline solution, for instance sodium carbonate or mono-ethanolamine solution, downward through the absorber tower from line 86. The remaining nue gases are vented through vent line 81. The liquid from the base of absorber tower 85 is passed through line 88 and heater 89 into stripper tower 90, heated at the base by coil 9|, and carbon dioxide removed 4in the stripping operation is passed through lines 92 and 93 and 86 back to coil 13 of reformer furnace 61. The stripped absorption medium is, of course, removed from the base of stripper and introduced by means of pump 94 through heat exchanger 89, cooler 95 and line 86 into absorber tower 85. Additional nue gas from the catalytic aromatization process, which will subsequently be described, can likewise be passed through lines' 88, 91, 98 and 84 to this carbon dioxide recovery system. Flue gas resulting from reviviflcation of aromatization catalyst can be used similarly if desired/being introduced into absorber 85 through lines 99, 91, 98 and 84. The recovered carbon dioxide plays a part in the reactions within the reformer furnace 61 and results in an increase in the ratio of carbon monoxide to hydrogen.

Synthesis gas from reformer furnace 61 passes out through preheater 12, heat exchanger |00 and line |0| into the Fischer reactor 15. In line |0| the synthesis gas is preferably supplemented by hydrogen resulting from the catalytic aromatization process and introduced through line |02 although, as will subsequently be described,

4 instead of introducing hydrogen through this line |02, it may sometimes be advantageous to utilize a portion of the synthesis gas passing through this line |0I from reformer furnace 81 in another phase of my process; namely, the s0- called parafln isomerization step. In the latter case the flow in line |02 is, of course, the reverse of that used when hydrogen is introduced `from line |02 into line IOI.

In any event the composition of the synthesis gas passing to reactor 15 is adjusted by controlling the amount of carbon dioxide introduced into coil 13 or by controlling the amount of hyldrogen, if any, introduced through line |02 or by still other means in order to bring the mol rtio of carbon monoxide to hydrogen to the optimum figure desired in connection with the synthesis of heavier hydrocarbons inthe Fischer synthesis. This optimum mol ratio is approximately 0.5 or slightly higher.

In the manufacture of Fischer liquid from carbon monoxide and hydrogen the synthesis gas is` contacted with a catalyst which can suitably be made of cobalt, thorium oxide and magnesium oxide mounted on kieselguhr although other catalysts known to the art can, of course, be used. The reaction conditions for this conversion step include a temperature of about 350 F. to about 450 F., for instance about 400 F. The temperature must be controlled within very narrow limits. The reaction is carried on at a moderate pressure which may be atmospheric pressure or considerably higher. Thus the pressure can be from about minus 2 to about. 25 pounds per square inch gauge, for instance about 2 pounds per square inch gauge. Various types of reactors .can be employed in contacting the synthesis las made up of carbon monoxide and hydrogen with the Fischer catalyst, the main necessity being provision for removal of the exothermic heat of reaction for precise temperature control of the process.

In the type of reactor shown the catalyst is granular in form and is disposed in a large number of small tubes to facilitate removal of the heat of reaction. These tubes can suitably be immersed in water maintained under pressure such that the boiling point is at the desired reaction temperature. Steam thus generated can be used in the synthesis gas manufacture step, being introduced through valved line 68. The Fischer products pass out through cooler |03 to separator |04. Cooler |03 is operated to condense the bulk of the gasoline range hydrocarbons and these are removed from the base of separator |04 by pump |05 and pass through line |06 to bubble tower |01. This bubble tower is heated at the base by coil |08 and reflux is provided at the top by coil |09. The bubble tower is so operated as to take of as a bottom product material heavier than gasoline which is preferably passed through line ||0, pump and line 62, into the so-called paraffin isomerization process which will be described later. The overhead from bubble tower |01 passes through line ||2 into fractionating tower H3.

Reverting to separator |04, the vapor phase from this separator is drawn off through compressors ||4 into absorber ||5 where it meets a descending stream of absorption medium which can be a conventional absorber oil. Unabsorbcd gas from absorber ||5 is preferably used as fuel in reformer furnace 61 via line 80 and lines 8| and 18. Alternatively, all or part of it can be vented through valved line ||6 or cycled to coil 13 in furnace 61 by means offline 80, valved line ||1 and line 66.

The rich oil from the base of absorber tower I5 is passed by means of pump I8 through heat exchanger ||9 into stripper tower |20 which is heated at the base by coil |2| and from which the previously absorbed hydrocarbons pass overhead through line |22 and condenser |23 to separator |24. The vapors from this separator can be recycled to the absorber through valved line |25 and one of compressors I|4 or vented through valved line |26. Liquids fromy this separator are passed through line |21 by pump |28 to the previously mentioned fractionating tower ||3. Lean oil from the base of stripper |20 is, of course, cycled through valve |29, heat ex-` changer ||9 and cooler |30 by means of pump |3| to the top of absorber ||5. Excess absorber oil can be removed or make-up absorber' oil added by means of valved line |32.

In fractionating column |3 the gasoline range and lighter hydrocarbons produced in the Fischer synthesis are preferably fractionated into three cuts of which the lightest is made up of two, three and four carbon atom hydrocarbons, the intermediate is made up of light naphtha and the heaviest is made up of heavy naphtha, for instance a fraction containing seven carbon atom hydrocarbons and heavier and having an end point in the general vicinity of 400 F. The first fraction is withdrawn from the'top of the column through line |33 and condenser |34 to separator |35 from which fixed gases are vented through valved line |36 and the liquid is cycled in part back to the column as reflux by the residue passes, as will subsequently be described, through valve |40 and'lines |4| and |42 to the alkylation step which forms a part of the preferred embodiment of my process.

The intermediate or light naphtha fraction from column I3 is withdrawn by means of trapout plate |43and can be'passed by means' of line |44, pump |45, vvalve |45 and line |41 into cycling all or part of this light naphtha fraction` of the Fischer synthesis products through line |44, pump |45, valve |52, and lines |4| and |42 to the alkylation reactor |53. When all of the light naphtha and the two, three and four carbon atom fractions are both to be sent to alkyla-l tion reactor |53, trapout plate |43 is, of course, unnecessary and all of these components can be taken overhead as a single fraction.

The feed to the alkylation step of my process can come from various sources. In general this feed comprises one or more sources of light olefins ranging from two to six carbonatoms per molecule and one kor more sources of branched chain paraffin hydrocarbons, particularly isobutane. As previously mentioned, one source of olefin hydrocarbons is the light naphtha from the Fischer process removed from fractionating column H3. This feed contains substantial quantities of pentenes and hexenes. The other preferred source of olenic hydrocarbons is likewise from the Fischer process; namely, the normally gaseous oleflns removed from the top of fractionating column 3 and passed into the alkylation system through lines 4| and |42. These light hydrocarbons normally predominate in three 'and four carbon atom hydrocarbons but may contain large quantities of two carbon atom hydrocarbons rich in ethylene. In general the lighter products from the Fischer synthesis are rich in olelns while the heavier fractions such as seven carbon atom hydrocarbons and heavier are made up very largely of straight chain parafilnic hydrocarbons. r

The olens from the Fischer process together, if desired, with oleflns produced by dehydrogenationor otherwise, are introduced into an alkylation reactor |53 along with hydrocarbons rich in isobutane. The latter can suitably come from the means of pump |31, valve |38 and line |39 while 76 distillate recovery operation since many distillates and natural gasolines are rich in isobutane. Such materials can be fractionated to secure a high concentration of isobutane but preferably a butane cut containingmormal as well as isobutane is utilized. This cut can, for instance, come from the base of the stabilizer column 41 through line |54, valve and lines |56 and |42 into alkylation reactor |53. Another and highly preferable source of isobutane which can, if desired, constitute the only source used in the alkylation reaction, merization process, which will later be described. Isobutane from this source passes into alkylation reactor |53 through lines |51 and |42.

In the alkylation process the ratio of isobutane to oleflns must be maintained high, for instance from 2:1 to 6:1 and preferably 3:1 or 4:1 with the result that the oil' gases from the alkylais the so-called paraffin iso tion reaction normally contain substantial amounts of isobutane and part of this material can be recycled to the alkylation reactor by means of pump |68 and line |69. 'Ihe alkylation process can be carried on by the use of various alkylation catalysts, for instance sulfuric acid or .sulfuric acid containing a promoter. I prefer, however, particularly when the olenic gases contain substantial amounts of ethylene or ilve and six carbon atom olens from the Fischer synthesis, to use aluminum chloride or an aluminum chloride complex as the catalyst in the alkylation reaction. Thecomplexes formed bythe reaction of aluminum chloride with hydrocarbons are decidedly preferable to the use of aluminum chloride alone'and these complexes differ radically in their efficiencies as alkylation catalysts. I prefer in particular to use a complex formed by the reaction of aluminum chloride with parain hydrocarbons although olefin hydrocarbon complexes can likewise be used. Complexes formed with aromatic hydrocarbons are much less desirable. One parn 5 cut withdrawn'i'rom trapout plate |83 of frac' tionating column |69 by selection of appropriate operating conditions for this column.

As shown, two product fractions are withdrawn from column |88, one from trapout plate |83 and the other from trapout plate |84. The operation of the column can be readily controlled, as will be apparent to those skilled in the art, in order to control the character and nature of the two fractions thus withdrawn. However, the fracalkylation process or passing through the alkylation process can be included in either fraction or l can be separately withdrawn. However, if light I Fischer naphtha from column ||3 is` sent to the ticularly satisfactory type of complex is that formed between isooctane and aluminum chloride.

The alkylation reactor |53 is shown in highly simplified form as a vessel containing a set of agitators |60 to promote intimate contactbetween the liquid complex, aluminum chloride or other catalyst on the one hand and the reacting gas on the other hand. Products from this alkylation reactor pass olf through line 6| to separator |62 and the liquid is pumped back to the reactor through valve |63 and line |64by pump |65. Spent catalyst can be removed through valved line |66 and fresh catalyst introduced from time to time through valved line |61 as required. 'I'he products from separator |62 pass through heat exchanger |68 into fractionating column |69 which is heated at its base by coil |10 and cooled at the top by reflux coil I1|.

Overhead gas from fractionator |69, largely two and three carbon atom hydrocarbons, can be used in synthesisgas manufacture, passing to this step via lines |12 and 66. If this gas contains a deleterious amount of sulfur it can be routed through hydrogensulfide removal plant 8 via line |13 by opening valve |14 and closing valve |15.

An important 4feature of my invention is that of cycling to alkylation reactor |53 or isomerization reactor |5| heavy alkylate, particularly alkylate heavier than octanes.. resulting from the reaction of pentenes and hexenes from the Fischer synthesis with isobutane under the influence of an aluminum halide or aluminum halide complex. One or both of two procedures can be followed: first, this heavy alkylate can be recycled to the alkylation reactor from the base of fractionator |69 via'pump |16, valve |11, line |18 and line |42 for conversion to lower boiling, highly'branched hydrocarbons; and/or, second, especially -when sulfuric acid is used as the catalyst in reactor |53 rather than an aluminum halide or aluminum halide complex type catalyst, the heavier-thanoctane alkylate can be cycled to isomerization revactor |5| via pump |16, valve |19 and lines |80,

|8l and 62 and then recycled to completion in this so-called paraflln isomerization process which results in the conversion of this heavy alkylate to high knock rating gasoline range hydrocarbons.

alkylation reaction through lines |4| and I 42, I prefer to include all or a large part of these six and seven carbon atom hydrocarbons in the stream withdrawn from the upper trapout plate |84 through line |85 and passed through this line and lines 8| and 62 to the so-called parailin isomerization step of my process. 'I'he reason for v this preference is that the paraffin hydrocarbons contained in the light Fischer naphtha are almost wholly normal and therefore pass through the alkylation step unchanged. Since these normal paraflin hydrocarbons are not desirable motor fuel components, they can be greatly improved A by the parailln isomerization operation. On the other hand, ii' the feed to the alkylation process does not include any large quantity o1' normal C5 and Cs paraffin hydrocarbons, the C5 and Cs components of the alkylation of! gases can better beincluded along with the C1 and Cs products in the stream withdrawn from trapout plate |83 and passed to product storage sincein this instance these materials are rich in high antie knocklaranched chain components.

Thenext step in my combination process can' be referred to as paraflln isomerization although f the strict accuracy oi' thisterm may be open to question since the process produces molecular weight changes as well as structural rearrangements withvcertain particular feed stocks. In any event under appropriate conditions normal paramnic hydrocarbons heavier than propane are If desired, a portion, or even all, of this heavy alkylate can be included in the heptane-octane converted in large measure into branched chain paraffin hydrocarbons under the iniluence of metal halide catalysts such as aluminum chloride, aluminum bromide and their hydrocarbon complexes in the presence of hydrochloric acid or other hydrohalogen acid or material which will decompose to form lhyclrohalogen acid. Thus normal butano charged to a process of this type gives high yields of isobutane While heavier normal hydrocarbons charged to the process result in the production of the same or lower molecular weight branched chain paramn hydrocarbons. A still more interesting phenomenon results when a. heavy normal paramnic charging stock is sub- Jected to such catalysts in the presence of alight parafnic material such as butane or isobutane. These two charging stocks react together in the presence of so-called paramn isomerization catalysts under reaction conditions to be described and produce materials of molecular weight intermediate to the molecular weight of the two parts of the charge or, in other words, within the gasoline boiling point range. vMoreover, these products are highly branched and therefore constitute extremely valuable motor fuel components particularly where a very high grade aviation motor fuel is required.

I prefer that the charge to isomerization reactor include as its main components the Fischer synthesis products heavier than gasoline or a selected fraction of those products and butanes from the alkylation process or from the distillate recovery operation or both. Thus the Fischer products boiling above 400 F. removed from the base of Abubble tower |101 can be introduced into isomerization reactor |5| by means of line ||0, pump and line 62 together with the butane-containing fraction withdrawn from column |69 of the alkylation process through lines |85 and |8|. Similarly all or part of the butane from the distillate recovery operation can. if desired, be introduced into isomerization reactor |5| through line |54, valve |86, pump |81, and lines |88 and 62. Another source of butane for the paraflln isomerization process is from the aromatization process which will later be described. The butane from this source can be introduced into isomerization reactor |5| through lines |89 and 62.

While the main preferred charges to the socalled paran isomerization process are the heavier-than-gasoline Fischer process products and butanes, ve and six carbon atom hydrocarbons from the distillate recovery operation can likewise be charged to this reaction in order to isomerize them and improve their quality as motor fuel components and this can be accomplished by means of pump 59, line 60, valve 6| and line 62. Another feed which can be used to advantage in the isomerization reaction is the heavy alkylate from lines |80, |8| and 62, as previously described. The light naphtha from the Fischer process can likewise go to isomerizer 5| through valve |49, pump |50 and line 62 but, as previously mentioned, it is preferable toA send it to alkylation reactor |53. Similarly the heavy Fischer naphtha can be isomerized by introducing it into isomerizer |5| via pump |90, valve |9| and line 62 but I prefer to aromatize this heavy naphtha as will hereinafter appear. Hydrogen chloride or other promoter is likewise introduced into isomerizer |5| by means of valved line |92.

The optimum conditions existing in isomerization reactor 5| are influenced to some extent by the nature of the charge to the isomerization reaction and by the nature of the desired products as well as by the precise catalyst chosen. However, in general these conditions include a temperature of from about 100 F. to about 450 F., for instance 280 F. and a pressure sufficient to maintain at least a substantial portion of the charge in the liquid phase. While aluminum chloride is the preferred catalyst, aluminum bromide and other polyvalent metallic halides can be used and the aluminum chloride or other metal halide forms a complex in the course of the reaction, which complex in all probability acts as the actual catalyst.

It is desirable to introduce hydrogen into isomerization reactor |5| in order to improve catalyst life, and the partial pressure of the hydrogen in the reactor can suitably be from about 50 to about 1500 or more pounds per square inch, for instance about; 1000 pounds per square inch. The lower the temperature, the lower should be the hydrogen partial pressure. In connection with my process, the hydrogen fed to the isomermatization operation via valve |99, lines |94 and |95. valve |96, compressor |91 (it required). and lines |98 and 92. Alternatively it can come from the Fischer synthesis gas (particularly ii the aromatlzation step is omitted) via line |02, valve |02a, line |95, valve |99, compressor |91 and lines |98 and 62. This synthesis gas contains, of course, carbon monoxide as well as hydrogen but it can nevertheless be used effectively in the isomerization reactor. Alternatively, however, synthesis gas from line |02 can be passed through a carbon monoxide removal step shown diagrammatically as purifier 200 by closing valve |02a and opening valves 20| and 202. This carbon monoxide purication can, for example, be carried out by oxidation of the carbon monoxide to carbon dioxide with steam followed by the removal of the carbon dioxide by some such means as absorber 85 and related elements.

Various types of isomerization reactors can be used, for instance a tubular reactor or a vessel equipped with an agitator as in the case of alkylation reactor |53. However, as shown, the charge is fed to a tower |5| which can be a bubble tower, packed tower, or baille tower and is therein agitated with the aluminum chloride complex or other liquid catalyst entering through conduit 208. The products pass overhead through linev 204 to separator 205 in which any entrained catalyst together with some of the heavier product liquid settles out and returns to the isomerizer through valved conduit 203. The catalyst reaching the bottom of the isomerizer tower is removed by means of pump 206 and passed through valve 201 to regenerator 208 and thence back to isomerizer |5| through line 209, separator 205 and conduit 203. Regeneration can be accomplished by hydrogenation or by decomposition and rechlorination or by electrolytic or other means. However, this regeneration does not constitute a part of the present invention and is merely indicated diagrammatically. From time to time fresh catalyst must be added and this is accomplished by means of valved line 2|0. Similarly spent catalyst can be removed through valved line 2| The reaction liquid from isomerizer |5| passes into separator 205 and, as previously mentioned,

f' the l'quid catalyst separates out as a lower liquid ization reactor preferably comes Afrom the arophase and is returned under the control of the valve in con`duit.203. The hydrogen and other fixed gases likewise separate at the top of separator 205 and are recycled by means of compressor 2|2 and lines |98 and 62.

The liquid isomerization product passes through line 2|3 to a wash tower 2|4 where residual hydrogen chloride is removed by an alkaline and/or water wash, the wash liquid being introduced by valved line 2|! and withdrawn through valved line 2|6. The washed product then passes through line 2|1 to i'ractionating tower 2|8 where it is preferably divided into three fractions. 'I'his tower is heated at its base by reboiler coil 2|9.

A light fraction is removed from tower 2|8 via line 220 and passed through condenser 22| to separator 222 from which a portion of it is returned to the top of the tower as reflux by means of pump 223, valve 224 and line 225, and the remainder, consisting largely of four' carbon atom hydrocarbons and rich in isobutane, is preferably cycled through valve 226 and lines 221, |51 and |42 to alkylation reactor |53 and normally constitutes the main part of the iso-paramn fed t the alkylation reaction:

The intermediate cut from fractlonating tower alkylation reaction.

The process as thus far described converts into valuable motor fuel components practically all of hydrocarbons produced in the aromatization reaction, as will hereinafter appear.

Reverting to" fractionating columnn 2|8, the Y heaviest fraction isv preferably recycled to isomerizer |5| by means o f line 23|, pump and line 62. This heavy fraction tends to break down into lighter compounds in the presence of the butane which is likewise introduced into isomerizer l5| and the process converts the heavy hydrocarbons charged to itin large measure into valuable motor fuel components and isobutane for use in th the original well fluids with the exception of hydrocarbons boiling between the boiling point of the seven carbon atom vhydrocarbons andthe gasoline end point which is typically in the neighborhood of 400 F. These hydrocarbons can be utilized most effectively and efficiently by raising their antiknock value without any radical change in their vo1atility,vthereby producing a final motor fuel of balanced distillation range and of uniformlyhigh antiknock quality throughout its distillation range. drocarbons from the Fischer synthesis and also from the distillate recovery operation, if such operation is utilized, constitute an extraordinarily effective 4charge for a. vcatalytic aromatization process. l.

Processes oi.' this latter type are'sometimes referred to as catalytic reforming in the presence of hydrogen or as dehydroaromatization processes. They are also 'sometimes popularly described as hydroforming processes. These processes are conducted in the presence of added hydrogen but nevertheless produce hydrogen since they operate by transforming naphthenic, olenic and parafllnic hydrocarbons into aromatic hydrocarbons with resultant dehydrogenation. In the present instance the charge is almost exclusively parainic and thus very substantial amounts of hydrogen are produced. Nevertheless the addition of hydrogen is essential in order to direct the reaction and particularly in order to prolong catalyst life.

Catalytic aromatizaton as contemplated in connection with the present invention is carried on in the presence of metal oxide catalysts and more particularly catalysts made up of oxides of one or more metals selected from the left hand columns of groups IV, V and VI of the periodic table. These oxides, preferably vanadium oxide, chromium oxide or still better molybdenum oxide, are most effective when supported on alumina. Thus, for instance, one particularly suitable catalyst is made'up of about 8% molybdenum oxide precipitated on the surface of activated alumina` Reaction conditions are criticaland include a pressure of from about 50-t9 about 450 Ypounds.

per square inch, for instance about 250 pounds'A4 per square inch, a temperature offrom about 850 These heavy naphtha hy-J to about 1025u F., for instance about 975 F..

a space velocity of from about 0.04 to Vabout 10 volumes of charge measuredA as liquid per gross volume of catalyst per hour, for instance about 0.7 volume per volume per hour. The volume of catalyst referred to above is the volume which the catalyst would occupy at rest or in a pelleted or compacted condition. The amount of hydrogen usedl should range from about l to about 10 mols per mol of charge, for instance about 2 mols of hydrogen per mol of charge.

This catalytic aromatization reaction can be carried out with the catalyst ina fixed bedlor a moving -bed but, as shown, it is carried .out by the Yuse of 'an upiiow reactor in which the flow velocity is adjusted to obtain a phenomenon" sometimes krrown'as'hindered settling in which the catalyst is maintained in a state of agitation resembling in appearance a boiling liquid. Typically this flow velocity is from 0.2 to about 3.0 feet per second, for instance 1.0 foot per second. 'I'hese velocities are for finely powdered catalyst and the velocity used will vary with the particle size of the catalyst. The velocity hould be high f enough to avoid bridging of the catalyst, but -low enough to maintain a dense catalyst suspension phase in the reactor.

Turning more particularly to the flow diagram, the heavy naphtha from the Fischer process passes through pump |50, valve 232 and lines 233 Y and 64 into coil 234 of aromatization furnace 235. Heavy naphtha from the distillate recovery operation, if any, passes through pump 63 and joins the Fischer heavy naphtha in line 64 and passes with it through coilsv 234 and 236 in furnace 235 and thence through transfer line 231 into upow reactor 238. In this reactor the naphtha meets a stream of hydrogen which is preferably part of the hydrogen produced inthe process recycled by means of valve 239, compressor 248, line 24| and coils 242 and 243 of furnace 235. Alternatively, the hydrogen and charge can be heated together but better control can be obtained by heating them in separate coils.

AS shown. the hot hydrogen is sent through valved line 244 and picks up regenerated catalyst from standpipe 245 and then enters the base of the reactor. A portion of the hydrogen can likewise be passed through valved line 246 and serves This latter catalyst is recycled without regeneration since in general in .this process catalyst can effectively be passed through the reactor more than once before regeneration is required. Furthermore, in order to maintain the desired fluid conditions within reactor 238 a, high ratio of catalyst to oil fed to the reactor is required and this can effectively be maintained by recycling Va portion of the catalyst without regeneration, thereby reducing the load on the regeneration system.

' The upflowing catalyst for reaction gases passes through thrcat248 to deector 249 which causes the bulk of the catalyst to precipitate on top of bafile 250. Further means can be utilized to recover residual catalyst including cyclone separators located in the separation zone of the reactor and lCottrell precipitators but these have been omitted in the interests of simplicity. A portion of the catalyst passes through standpipe 25| to, conduit 252 where it is picked up in a stream ofA air introduced' through valved line 253 .and passed to a regenerator 255 which is preferably of the` upflow hindered settling type as described previously. The amount of air introing with consequent deleterious effect on the catalyst. Since the upilow type of reactor gives good heat distribution with practically no temperature gradients, the problems of local hot spots usually encountered in other oxidation and revivication systems is avoided.

As in the case of reactor 238, a portion of the catalyst reaching the top of regenerator 255 is cycled via standpipe 256 back to the regenerator while another portion passes through standpipe 245 and is picked up by the stream of hydrogen and introduced into reactor 238 as previously described. The catalyst passing through standpipe 256 can advantageously be cooled by cooler 251. The catalyst in standpipes 245, 241, 25| and 256 is maintained in aerated condition by introducing small amounts of gas through valved lines 258. As previously described, the ilue gas from regenerator 255 can suitably be passed through lines 99, 91, 98 and 84 and stripped. of its carbon dioxide in absorber 85, thereby augmenting the supply of carbon dioxide for use in connection with the manufacture of synthesis gas for the Fischer process.

Returning to reactor 238, the products after catalyst removal pass through line 259 and condenser 260 to separator 26| and the hydrogen remaining uncondensed along with some contaminants passes in part back to the aromatization v process and in part to the so-called paraffin isomerization process via valve |93, lines |94 and |95, valve |86, compressor |91 and lines |98 and 62. Also, as previously mentioned, a portion of this hydrogen can be used as part of the synthesis gas for the Fischer reaction, being sent to Fischer reactor as hereinabove described. EX- cess hydrogen, if any, can be vented through valved line 262.

The liquid from separator 26| passes through line 268 into fractionating tower 264 which serves to separate the aromatization products into the desired fractions. This aromatization product contains a very small amount of the C2 and C: hydrocarbons and the former can suitably be used in the manufacture of synthesis gas for the Fischer process being sent to the reformer coils 13 and 14 from the top of fractionator 264 through lines 265, |12 and 66. This gas can also be used as fuel for furnace 61 and/or 235 using valved line 16 in the first instance and valvedline 266 in the second.

The four carbon atom hydrocarbons are taken ofi' from tower 264 through trapout plate 261 and can suitably constittue part of the feed to the socalled parailin isomerization process passing to that process through pump 268 and lines 269, 210, |89 and 62. Any five carbon atom hydrocarbons present in the products from the aromatization reaction can, if desired, also be sent to this isomerization process and this is desirable since five carbon atom hydrocarbons cannot be aromatized but do form branched chain hydrocarbons of high antiknock value in the isomerization reaction. Care should be taken, however, that no aromatics are cycled to the isomerization process since they poison the catalyst.

Aromatization products heavier than C5 and within the motor fuel distillation range, say up to 400 F., can be utilized in the manufacture of automotive and aviation motor fuels in various ways. Thus a cut rich in hydrocarbons having six and seven carbon atoms per molecule is particularly desirable for use in aviation gasoline 0f extremely high knock rating. A fraction of this type can be withdrawn from trapout plate 21| 75 The fraction from the aromatization operation' which contains the xylenes and heavier components within the gasoline range can suitably be withdrawn from trapout plate 28| through line 282 and introduced into storage tank 263 for subsequent blending in the manufacture of high grade gasoline engine motor fuels since these heavier aromatics are somewhat inferior to the highly branched parains of like boiling point when used in high temperature engines such as are increasingly typical of aviation practice.

The aromatization reaction produces a small quantity of hydrocarbons heavier than gasoline and these are often referred to as polymer." Also when hydrocarbons heavier than gasoline from the distillate recovery process are charged to the aromatization step, they are aromatized and form a part of the heavy product. This heavier-than-gasoline aromatized fraction con stitutes an extremely effective absorption medium for use in connection with distillate recovery by high pressure absorption and all or a portion of this fraction can thus suitably be withdrawn from tower 264 by means of trapout plate 284 and directed through line 285, valve 286, pump 261, lines 288 and |1, heat exchanger I8 and cooler |9 into absorber I0. The remainder can be passed through valved line 289 to polymer storage tank 290 and can subsequently be used in .various Ways suitable for highly aromatic ma.-

terials of this boiling range.

Any small amount of catalyst passing overhead from reactor 238 finds its way to the base of fractionating tower 264 and can be recycled along with any extremely heavy hydrocarbons back into the aromatization reactor by means o1' pump 26|, lines 292, coils 234 and 236 and transfer line 281.

While I have described my invention in connection with certain preferred embodiments thereof, it is to be understood that these are by way of illustration and not by way of limitation and I do not mean to be restricted thereto but only to the scope of the appended claims.

I claim:

1. A method for manufacturing high quality motor fuel comprising passing well fluids produced by a well of the distillate type into an absorption zone at high pressure, contacting said well fluids in said absorption zone with a heavy absorber oil comprising aromatic hydrocarbons produced by an aromatization operation, recovering from the rich absorber oil a fraction rich in butanes and a heavy naphtha fraction, compressing a portion of the unabsorbed gases from said absorption zone and injecting same into a subsurface hydrocarbon reservoir, passing a further portion of the unabsorbed gases from said absorption zone to a hydrogen sulfide removal zone and removing hydrogen sulfide from said gases in said last mentioned zone, reacting said purified gases with steam in the presence oi' a catalyst to produce carbon monoxide and hydrogen therefrom, reacting said carbon monoxide and hydrogen in the presence of a catalyst of the Fischer synthesis type to produce a synthetic crude oil therefrom, fractionating said synthetic accesos crude oil to produce fractions including a predominantly paramnic heavy naphtha fraction, passing at least a substantial part o! said heavy naphtha fraction from said synthetic crude oil and at least a substantial part of said heavy naphtha fraction recovered from said rich absorber oil to a catalytic aromatization zone and contacting it in said zone with an aromatization catalyst of the metal oxide type in the presence of hydrogen to produce aromatic hydrocarbonsl and hydrogen, fractionating said aromatic hydrocarbons into at least one gasoline range fraction and at least one heavier-than-gasoline fraction, and cycling at least a substantial part of said last mentioned fraction to said contacting step as said heavy absorber oil. s

2. A method formanufacturing high quality motor fuel comprising passing Well fluids produced by a well of the distillate type into an absorption zone at high pressure, contacting said well iiuids in said absorption zone with a heavy absorber oil comprising aromatic hydrocarbons produced by an aromatization operation, recovering from the rich absorber oil a fraction rich in butanes and a heavy naphtha fraction, compressing a portion of the unabsorbed gases from said absorption zone and passing said compressed gases to a subsurface hydrocarbon reservoir, reacting at least a substantial part of said gases with steam in the presence of a catalyst to produce carbon monoxide and hydrogen therefrom,

- reacting said carbon monoxide and hydrogen in the presence of a catalyst of the Fischer synthesis type to produce a synthetic crude oil therefrom, fractionating said synthetic crude oil to produce a light fraction rich in oleilnic hydrocarbons, a heavy naphtha fraction and a predominantly paraffinic fraction boiling predominantly above the gasoline boiling point range, passing at least a substantial part of said last mentioned fraction and at least a substantial part of said heavy naphtha fraction recovered from said rich absorber oil into a parailln isomerization zone and contacting it therein with butane to react said predominantly parainic fraction with said butane in the presence of a catalyst of the aluminum chloride type to produce isobutane and predominantly branched chain hydrocarbons intermediate in boiling point between said butane and said predominantly parailinic fractionmpassing at least a substantial part of said isobutane together with at least a substantial part of said light fraction from said synthetic crude'oil to an alkylation zone and contacting said isobutane and said last mentioned fraction with each other and with a catalyst of the aluminum chloride type to produce a highly branched chain alkylation product together with a normal butane fraction passing through said alkylation zone, passing at least a substantial part of said normal butane fraction and at least a substantial part of said fraction rich in butanes recovered from said rich absorber oil to said isomerization zone as at least a part of the butane contacted in said isomerization zone as aforementioned, passing at least a substantial part of said heavy naphtha fraction from said synthetic crude oil to a catalytic aromatization zone and contacting it in said zone With an aromatizationcatalyst of the metal oxide type in the presence of hydrogen to produce aromatic hydrocarbons and hydrogen, cycling at least a substantial portion of said produced hydrogen to said paraflln isomerization zone to increase the life of said catalyst of the aluminum chloride type present in said isomerization zone, fractionatlng said aromatic hydrocarbons into at least one gasoline range traction and at least one heavier-than-gasoline fraction, cycling at least a substantial part of said last mentioned iraction to the nrst mentioned contacting step as said heavy absorber oil, and blending selected gasoline range fractions of the products produced in said alkylation, isomerization and aromatization steps to produce a superior motor fuel of high antiknock value.

3. A method for manufacturing high quality motor fuel comprising passing well fluids produced by a well of the distillate type into an absorption zone at high pressure, contacting said well fluids in said absorption zone with a heavy absorber oil comprising aromatic hydrocarbons produced by an aromatization operation, recovering from the rich absorber oil a fraction rich in butanes and a heavy naphtha fraction, compressing a portion ci the unabsorbed gases from said absorption zone and passing said compressed gases to a subsurface hydrocarbon reservoir, passing a further portion of the unabsorbed gases from said absorption zone to a hydrogen sulfide removal zone and removing hydrogen sulde from said gases in said last mentioned zone, reacting said purified gases with steam in the presence of a catalyst to produce carbon monoxide and hydrogen therefrom, reacting said carbon monoxide and hydrogen in the presence of a catalyst of the Fischer synthesis type to produce a synthetic crude oil therefrom, fractionating said synthetic crude oil to produce a light fraction rich in olenic hydrocarbons, a heavy naphtha fraction and a predominantly parafllnic fraction boiling predominantly above the gasoline boiling point range, passing at least a substantial part of said last mentioned fraction and at least a substantial part of said heavy naphtha fraction recovered from said rich absorber oil into a parailin isomerization zone and contacting it therein with butane to react said predominantly parafiinic fraction with said butane in the presence of a catalyst of the aluminum chloride type to produce isobutane and predominantly branched chain hydrocarbons intermediate in boiling point between said butane and said predominantly paraiiinic fraction, passing at least a substantial part of said isobutane together with at least a substantial part of said light fraction from said synthetic crude loil to an alkylation zone and contacting said isobutane and said last mentioned fraction with each other and with a catalyst of the aluminum chloride type to produce a highly branched chain alkylation product together with a normal butane fraction passing through said alkylation zone, passing at least a substantial part of said normal butane fraction and at least a, substantial lpart of said fraction rich, in butanes recovered from said rich absorber oil to vsaid isomerization zone as at least a part of the butane contacted in said isomerization zone as aforementioned, passing at least a substantial part of said heavy naphtha fraction from said synthetic crude oil and at least a substantial part of said heavy naphtha fraction recovered from said rich absorber oil to a catalytic aromatization zone and contacting it in said zone with an aromatization catalyst of the metal oxide type in the presence of 'hydrogen to produce aromatic hydrocarbons and hydrogen, cycling at least a substantial portion of said produced hydrogen to said paraiiin isomerization fzone to increase the life of vsaid catalyst of the aluminum chloride type present in said isomerization zone, fractionating said aromatic hydrocarbons into at least one gasoline range fraction and at least one heavier-than-gasoline fraction, cycling at least a substantial part of said last mentioned fraction to the first mentioned contacting step as said heavy absorber oil, and blending selected gasoline range fractions of the products produced in said alkylation, isomerization and aromatization steps to produce a superior motor fuel of high antiknock value.

4. A method for manufacturing high quality motor fuel from a stream of light and heavy hydrocarbon vapors comprising the combination of \steps of contacting said stream with an absorber oil comprising aromatic hydrocarbons produced by a catalytic conversion, recovering from the rich absorber oil a heavy naphtha fraction, subjecting at least a substantial part of said heavy naphtha fraction to a, catalytic conversion in the presence of a catalyst of the metal oxide type to prsduce aromatic hydrocarbons, recovering from he conversion products at least one fraction boiling in the motor fuel range and at least one heavy fraction containing substantial amounts of aromatics, and supplying at least a part of said heavy fraction as the absorber oil in the contacting step.

5. A method for manufacturing high quality motor fuel from a mixture of light and heavy hydrocarbon vapors comprising the combination of steps of contacting said mixture with an absorber oil comprising aromatic hydrocarbons produced by subsequent catalytic conversion, recovering from the rich absorber oil a heavy naphtha fraction, recovering an unabsorbed predominantly methane fraction from said mixture and producing a mixture of hydrogen and oxides of carbon therefrom, preparing a synthetic crude oil by catalytic reduction of the oxides of carbon, `fractionating the said synthetic crude oil to produce fractions including a predominantly paralinic heavy naphtha fraction, subjecting at least a substantial part of said parailnic heavy naphtha fraction from said synthetic crude oil to a catalytic conversion in the presence of hydrogen and a Icatalyst of the metal oxide type to produce aromatic hydrocarbons, recovering from the conversion products at least one fraction boiling in the motor fuel range and at least one heavy fraction containing substantial amounts of aromatics, and supplying at least a part of said heavy fraction as the absorber oil to the contacting step. i

6. A method for manufacturing high quality motor fuel from a stream of light and heavy hydrocarbon vapors comprising the combination of steps of contacting said stream with an absorber oil comprising aromatic hydrocarbons pro- -duced by a catalytic conversion, recovering from the rich absorber oil a heavy naphtha fraction, producing a mixture of hydrogen and oxides of carbon from the unabsorbed gases and preparing a synthetic crude oil by catalytic reduction of the oxides of carbon, recovering from the said synthetic crude oil a fraction including a predominantly paraiiinic heavy naphtha fraction, subjecting at least a substantial part of said paraflinic heavy naphtha fraction from said synthetic crude oil and said rst mentioned heavy naphtha fraction to a catalytic conversion in the presence of hydrogen and a catalyst of the metal oxide type to produce aromatic hydrocarbons. recovering from the conversion products at least one fraction boiling in the motor fuel range and at least one heavy fraction containing substantial amounts of aromatics, and supplying at least a part of said heavy fraction as the absorber oil in the contacting step.

7. A method for manufacturing high quality motor fuel from a stream of light and heavy hydrocarbon vapors comprising the combination of steps of contacting said stream with an absorber oil comprising aromatic hydrocarbons produced by a catalytic conversion, yrecovering from the rich absorber oil a fraction rich in butanes and a heavy naphtha fraction, producing a mixture of hydrogen and oxides of carbon from the unabsorbed gases and preparing a synthetic crude oil by catalytic reduction of the oxides of carbon, fractionating the said synthetic crude oil to produce fractions including a, light olenic fraction and a predominantly parailnic heavy naphtha fraction, catalytically alkylating said butanes and said light olenic fraction to produce hydrocarbons boiling in the motor fuel range, subjecting at least a substantial part of said parafnic heavy naphtha fraction from said synthetic crude oil to a catalytic conversion in the presence of hydrogen and a catalyst of the metal oxide type to produce aromatic hydrocarbons, recovering from the conversion products at least one fraction boiling in the motor fuel range and at least one heavy fraction containing substantial amounts of aromatics, and supplying at least a part of said heavy fraction as the absorber oil inthe contacting step.

MAURICE H. ARVESON.

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