US3239449A - Selective conversion of unstable liquids - Google Patents

Description

SELECTIVE CONVERSION oF UNSTABLE LIQUIDs Filed Nov. 19, 1962 United States Patent O 3,239,0349 SELECTIVE CNVERSIN F UNSTABLE LQUIDS Richard G. Graven, North Castle, and Vernon 0. Bowies,

Bedford Township, Westchester County, NPY., assignors to Socony Mobil Oil Company, Inc., a corporation of New York Filed Nov. 19, 1962, Ser. No. 233,690 21 Claims. (Cl. 208-143) The present invention relates to a process for the selective conversion into more stable substances of unstable liquids which tend to deposit solids upon heating. It is particularly concerned with improved control of vaporization of intermediate liquid products between reaction stages in selective conversion processes having both liquid or mixed phase reactions and vapor phase reactions. In a preferred embodiment it involves the selective hydrogenation in two or more stages of a liquid hydrocarbon mixture containing aromatic hydrocarbons, olefins, diolelins and sulfur compounds, and especially a mixture of narrow boiling range.

Selective hydrogenation serves many purposes and the instant invention is particularly concerned with a process in which an unstable hydrocarbon mixture of the type mentioned above is hydrogenated in at least two stages of increasing severity to prepare a stable product from which valuable aromatic hydrocarbons can readily `be separated by solvent extraction with a solvent such as diethylene glycol. In such extraction, it is relatively easy to separate benzene and other aromatic compounds `from paraiiins or naphthenes; but this is not true of separating benzene from unsaturated aliphatic components and especially from organic sulfur compounds in the mixture.

To prepare a suitable feed for the solvent extraction, it is necessary to convert the organic sulfur compounds to a readily separable material such as hydrogen sulfide gas, to saturate the unstable gum-forming dioletins and also to saturate the mono-oleiins without converting aromatic hydrocarbons into naphthenes by excessive hydrogenation. Although it is easy to specify the reactions with hydrogen for obtaining these results, performing them in commercial practice has -been an entirely different matter. There is an increasing demand for the production of aromatic hydrocarbons from petroleum so that the supplies of these compounds are not restricted to the current production level of the steel and coking industries. Despite this demand there was still no fully satisfactory commercial method for the hydrogenation of such mixtures of aromatic and unsaturated aliphatic hydrocarbons prior to the present invention.

It is not feasible to saturate and desulfurize such feed stocks in a single operation because the relatively high temperatures suitable for hydrodesulfurization also promote the formation of coke and oletins polymers or gums and may induce the hydrogenation of aromatics to naphthenes. Prior to the present invention, conducting the hydrogenation reactions in two stages to avoid or minimize the aforesaid deficiencies has not been entirely satisfactory by reason of the accumulation of polymeric deposits that reduce the activity of hydrogenation catalysts, thereby necessitating frequent regeneration. Such deposits also plug up piping and other equipment. Not only thermal polymerization but also catalytic polymerization 4must be minimized as some hydrogenation and desulfurization catalysts also catalyze the polymerization of diolens. While various techniques are known for at least partially reducing ypolymer formation in hydrocarbous at elevated temperatures, nevertheless polymer formation has remained a critical problem in commercial rrr ICC

plants for the selective hydrogenation of charging stocks oi the type described.

An object to the invention is to provide an improved process for the selective conversion of unstable liquids into relatively stable liquids having little or no tendency to polymerize or otherwise deposit solids upon standing.

Another object of the invention is to provide an improved process that includes simple, direct and positive control of the vaporization of liquid at a heat labile stage during a selective conversion process.

A further object of the invention is to provide an improved selective conversion process for the direct regulation of the vaporization of a liquid of narrow boiling range.

An object of the present invention is to provide an improved process for the selective hydrogenation of an unstable mixture of organic compounds.

Another object of the invention is to provide a process for hydrogenating a mixture boiling below aboutA 500 F. of aromatic and unsaturated aliphatic hydrocarbons without the formation of substantial quantities of napht'hene's or polymers.

A further object of the invention is to provide an improved process for the selective, nondestructive hydrogenation of unstable, thermally-cracked, petroleum products with a boiling range below about 500 F., which hydrogenation is performed in stages of increasing severity without excessive deactivation of the contact catalysts.

Still another object of the invention is to provide an improved process for 4the selective, nondestructive hydrogenation of a mixture of aromatic and oleiinic hydrocarbons boiling 'below about 5007 F., and preferably below about 275, in which contact catalysts are kept on stream for substantially longer periods.

A still further object of the invention is to provide an improved process for the selective hydrogenation of an unstable hydrocarbon mixture in which not more than a very small amount of polymeric material is formed and it is removed from the intermediate product stream prior to the final hydrogenation reaction.

Other objects and advantages of the invention will be apparent to those skilled in the art -upon consideration of the following detailed disclosure in which all temperatures are expressed in terms of degrees Fahrenheit, all proportions in terms of weight and all boiling point or ranges of temperatures are measured at atmospheric pressure according to the ASTM procedure unless otherwise stated.

The present invention relates to the selective conversion of unstable liquids with a pronounced tendency to deposit solids upon heating by a process which includes partially converting unstable compounds in a liquid feed into more stable substances within a confined initial reaction zone under conversion conditions in which a considerable proportion of said feed is maintained in the liquid phase, vaporizing a substantial portion of the eiiluent liquid from the initial conversion reaction by controlled heating, and passing the gaseous phase derived from said initial eiiiuent through a confined conversion zone while substantially completing the conversion of unstable components of said feed; and in which the improvement comprises separating in an enlarged separation zone a liquid iiux in an amount equal to at least about 0.5% (at least 5% being preferred) of said liquid feed from said gaseous phase, withdrawing said liquid ux at a substantially constant rate from a pool thereof maintained in said separation zone and regulating said controlled heating operation in direct response to the rate of collecting said liquid flux in said separation zone as determined from the level of said pool.

Narrower aspects of the invention include, inter alia the use of contact catalysts in the various reaction and conversion zones, increased severity of conversion conditions in a subsequent conversion stage in comparison with the initial reaction, retaining specified proportions of the initial eliluent in the liquid state after the controlled heating and separation, heating the efliuent during passage through a restricted transfer conduit (a term employed herein in its broad sense .to include heaters, etc.), heating the initial effluent by the injection of a hotter stable gaseous reactant either upstream or downstream from the separation-zone or both, which heating by injection may be Vregulated as described rhereinbefore, and the sources, amounts and circulation of liquid linx. Feeds of narrow boiling range (eg. those with a breadth ofboiling range which does not exceed about 150 or does not exceed about 80 on the rFahrenheit scale) are particularly contemplated since the vaporization control of-Y such materials by conventional methods lacks precision;-

whereas the present method is especially suitabler for'the vaporization of liquids of very narrow boiling ranges or even those having boiling point rather than a range because regulation of the controlled heating is independentY of the boiling and vaporization characteristics of the ini-v within certain specified ranges to providea hydrogenation effluent in which at least about 35%, and preferably at least about 60%, of the diolens have been at least partially saturated and in which an amount equal to at least about 20%, and preferably at least about 60%, of the liquid feed remains in the liquid phase. The, Bromine Number of the normally liquid fraction of said eiiluent is also desirably reduced at least 25% below that of the liquid feed.` Said conditions which are regulated therel include maintaining a hydrogen partialpressure within the range of Vabout 200e800 (about 300-600 being pre-l ferred) pounds 'per'square inch absolute pressure( herein-A after abbreviated p.s.i.a.), an hourly spacevelocity'with-IV in the range of about 0.2-15.0, and preferably about 0.5-8.0, based on the Volume of liquid feed, the hydrogen temperature within the broad range of about 75-300 F.,- and preferably :not over about 250. Controlled Vaporization of liquid in said hydrogenation eiiuent and separation of theV gaseous, phase thereof in an enlarged separationV zone is eifected in vthe presence of a liquid liuX, for example an amount of ilux equal to at least t about 5% (preferablyover 10%) of said liquid feed.

Y gaseous material derived from saidfyaporization step, is

passed through a subsequent conversion zone in contact with a porous solid conversion catalyst ofat least mod-'.-

erate hydrogenation activity and high` desulfurizration activity at a substantially higher average temperature than in said initialzonewhile lregulating conversion conditions within certain specified rangesto produce agconversion eiliuent with a normally liquid fraction having a Bromine Number'less than about 4, .preferably below about 2, and an organic sulfur content `belowabout 20 p.p.m., and -Y preferably below 15 p.p.m. The speciiiedpranges of conversion conditions includemaintaining a hydrogen par-vv tial pressure with-in the range of about `20G-800 (about 1 30G-600 being preferred) p.s.i.a., an hourly space velocity within the range of about 0.2-6.0 (about,0.54.0 being; preferred) based on the volume of original liquid feed,

the total hydrogencharge (unreacted hydrogen plus any newly introduced hydrogen) within the' range of about G-10,000 (preferablyfabout 2,000-5,000) s.c.f.b. ofV said liquid feedand an inlettemperature withinv the` z' wide range of about S50-700, and preferablyk about e 'In' performing the preferred modificationof/ the instant process, a feed stock with a pronounced tendency toward' jundesired polymerization is subjected to a selective hydrogenation pro-cess in which it is first hydrogenated mildly inthe liquid -or mixed phase; the resulting efliuent'is vaporized under controlled conditions and iinally treated in a the gaseous phase with hydrogen under more severe conditions. The initial hydrogenation isiconducted at a temf perature sufficiently low to avoid or minimize boththermal and catalytic polymerizationgwhile hydrogenating a substantial portion and usually almost all of the -diole- `tins, including all of the more reactive ones. During the' r initial hydrogenation little, if any, desulfurization isV acf.

complished and a substantialproportion of the mono-olelins usually remain unsaturated here as, under some con-V without depositing polymeric solids on the equipment, if

the vaporization is accomplished by heating in a carefully controlled manner up until the liquid and gaseous, phases are separated from one another completely and rapidly in an enlarged separation zone. It is to be noted that the Y charge within the range of about SOO-6000, and preferv ably about 1200-3000, standard cubic feet per barrel f ther special-precautions to a temperature suitable for de- (hereinafter abbreviated s.c.f.b.) of liquid and a feed Such vaporization isproduced by a controlled heating.

operation that is regulated in direct response to the rate of collecting said flux in said separation zone and preferably in direct response to the rate of collecting the liquid fraction of fresh initial hydrogenation eiiiuent. The flux may be any liquid substance or mixture which is rniscible v with and unreactive with said liquid feed including immediate reaction products, recycled partially or fully hydrogenated products of the instant process or vextra neous liquids, preferably having a substantial content of hydrocarbon vapor phase is sharplyV and completely separated for `the tir-st time from the circulating flux liquidand 4the, unvaporized fraction of the Veffluent' by removal of the vapor phase and lnot'by evaporating therliquidkf.

phase to dryness orl a close approach to dryness.

This gaseous phaseis further heated withoutanyI fursulfurization and also for olen saturation and then itis subjected to la catalytic' conversion with hydrogenat a distinctly higher temperature than in the initial hydro-` genat'ion. In this conversion,-thesaturationrof all-remaining mono-olefins and diolens is substantially corn- Vpleted andorganic sulfur compounds are converted into.

hydrogen sulfide without any appreciablevpolymerfor.

mation occurring in either the preliminary heating or in, f the vapor phase reaction even though the catalyst is euse.

tomarily of a type of high polymerization,potentialsY When the catalyst in the initial reactor is in relatively, .Y

fresh condition, most of the hydrogenation of monoole fins (often more than as well fas diolens, occurs there.k S little? hydrogenationr of the hydrocarbonsVoc-V tinued use, more of lthe hydrogenationl loadlris shifted to Y Y V; the second or linal reactorY andsubstantial'increases be#A tween theV inlet :and outlet temperatures'. of this reactor); are then apparent., Y Y Y Eventually as theY activity As the starting material, any mixture of aromatic and unsaturated aliphatic hydrocarbons with organic sulfur compounds may be employed in this embodiment process if the final boiling point of the liquid does not exceed about 500. A narrow boiling range material, for example, one having a boiling range between about 140 and 275, is desirable, and preferably a -charging stock boiling in the range about 160 to 220.

As `a source of aromatic hydrocarbons, the liquid feed -desirably contains a total -of between about 20 and 90% aromatics, especially benzene and toluene. Typically, it also has substantial contents of diolefins and olefins as evidenced by Diene Numbers of about to 22 which measure the proportion of lconjugated diolefins as determined by the ma'leic anhydride condensation method and Bromine Numbers of about to 30 which represent the total content of unsaturated aliphatic hydrocarbons. Feeds with Diene and Bromine Numbers as high fas about 40 and about 75 respectively may also be processed according to this phase of the present invention. The -organic sulfur content is typically about to 300 p.p.m. :and may be as high las about 700 p.p.m.

In other utilizations of the instant process, the charging stock need not be rich in aromatic hydrocarbons. For instance, in producing a stable gasoline blending stock from a pyrolysis liquid, a feed containing 6 to 20% aromatics compounds is typical.

Feed stocks of the nature described are unstable as they tend to form polymeric gurus readily. It has been found desirable to keep the period of storing them as brief as possible in order to minimize the introduction of gum into the present process. In addition, it is recommended that the liquid feed stock be free of dissolved oxygen and be ,stored in the substantial absence of oxygen or air, for example, under a blanket of an inert gas such as nitrogen. This prolongs the activity of the catalysts usable in this process. Such feed stocks :are generally obtainable by severely thermally cracking a petroleum fraction suitable for the manufacture of gasoline or light `olefins, as exemplied by ethylene. It is preferred here to depentanize the cracked product. A partially suitable feed is one within an end point not exceeding 220 and a maximum gum content of less than 15 milligrams per 100 milliliters.

The total consumption of hydrogen in this modification of the invention Varies of course -with the particular feed stock employed; but in general, it is in the range of about 150-800 s.c.f.b. of liquid feed stock. A typical value is 300 s.c.f.b. with a charging stock of Diene and Bromine Numbers of l5 and 24 respectively, and the consumption is usually found to be less than 500 s.c.f.b. Substantial ,excesses of hydrogen have been specified -hereinbefore to avoid a drop in the hydrogenation rates as a result of an inadequate supply of hydrogen. Although pure hydrogen may be used, it is customarily supplied as a mixture of hydrogen `and gaseous hydrocarbons in the off gases of units for reforming naphthas or hydrodesul-furizing gas oils, etc. The gas charge preferably has 1a hydrogen content of at least 60% by volume but gaseous mixtures with as little as 40% hydrogen may be used.

The partial pressure of hydrogen in the t-wo or more reactors -is important in avoiding undesired side reactions, such as the formation of gum or coke on the catalysts. It should be maintained within the range of about 200- `800 p.s.i.a., in which the 300-600 p.s.i.a. range is -preferred. The total pressure in the reactors is not critical, but it -s'hould not be so high as to interfere 'significantly with the vaporization of the `feed and reaction produ-cts described herein. A major proportion of the product gases with much unconsumed hydrogen is recycled to the process after `any excessive quantities of hydrogen sulfide have been scrubbed out and this usually constitutes ra major proportion of the total quantity of gases charged to the reactors.

The charging stream of combined recycle and makeup gases containing hydrogen is divided into several streams.

A substantial quantity of hydrogen must be introduced into the first reaction zone along with the liquid feed, and unreacted `hydrogen is present in the effluent of that reaction which is subjected to further hydrogenation reactions. While in theory all of the hydrogen-rich gas required for the series of selective hydrogenations can be charged to the initial reactor, this is not particularly desirable in practice. Especially since the circulating gas m-ay be employed after heating to a high temperature in a furnace as a heat source to aid in vaporizing the effluent from the initial reactor and also to regulate the temperature of the vapor phase charge of hydrocarbons and hydrogen to a subsequent reactor. Alone this hydrogen-rich gas stream can be heated without decomposition or other difficulty to a temperature several hundred degrees higher than it is possible to heat the liquid or mixed phase effluent of ythe first reactor without the coincident deposition of polymeric gum or coke. Such deposit-ion from the mixed phases can occur at temperatures of 300 and even lower. In serving as a heat source, a substantial part of the total circulating gas, say about 30 to 85%, is heated to a temperature in the range of about 500 to 950, and preferably in the range of about 60G-850, while the unheated balance of the gas is charged to the initial reactor. One stream desirably containing more than half of this heated gas is used to supply the final temperature increment to the initial mixed phase effluent just prior to entering the enlarged separating and vaporizing chamber, and the remainder may be introduced into the wholly gaseous stream leaving the top of said chamber on its way to the second stage conversion reactor as the final heat increment to adjust the charge to the desired inlet temperature.

In the preferred modification of the novel process, a catalyst of high hydrogenation activity is required for the initial reaction zone as it must hydrogenate at a relatively low temperature the more reactive conjugated diolens and usually at least some of the other olefins, but its polymerization activity must be relatively low in order to avoid the formation of gums which will deactivate the catalyst. While suitable hydrogenation catalysts also incidentally possess relatively high desulfur-ization activity initially, this property drops ofi rapidly in a period of a few days to a week, because such catalysts are readily poisoned with respect to desulfurizing ability at desulfurizing temperatures by feeds containing much organic sulfur.

These qualities of the catalysts may be defined in terms of arbitrary activity indexes which are described herein. Unless otherwise stated, all such indexes are measured using fresh new catalyst. The activity indexes enable one to clearly differentiate between the two or more catalysts employed at various stages in the instant process.

For delineating hydrogenating `activity in a preferred embodiment, two different indexes are available. The hydrogenation activity index is defined herein as the percentage or proportion of isoprene which is converted to pentenes and pentanes when a blend of 8-l0% isoprene and 50-500 p.p.m. of thiophene sulfur in benzene is passed over the catalyst with 1500-3000 s.c.f.b. of hydrogen gas at F., 300 pounds per square inch gage (hereinafter designated p.s.i.g.) as the total pressure and a liquid hourly space velocity of 5. Thus a conversion of half of the isoprene present, or 4.5% out of a total of 9.0% isoprene present, signifies that the activity index is 50. For the initial catalyst, a hydrogenation activity index of at least about 40 is recommended.

In determining the benzene conversion index as another and usually supplemental measure of hydrogenation activity, a sulfur-free mixture of 17% benzene and 83% cyclohexane is passed through the catalyst under test at 400 F. and 400 p.s.i.g. with 1500-3000 s.c.f.b. hydrogen circulation and a liquid hourly volumetric space velocity of 2. The test mixture must be sulfur-free inasmuch as organic sulfur in a content as small as 50 p.p.m., and even less in the form of hydrogen sulfide, totally inhibits the hydrogenation of benzene with such catalysts under the specified reaction conditions. A suitable catalyst for the initial reactor has a benzene conversion index of at least about 50, meaning that half of the benzene present or 8.5% is converted into cyclohexane, but an index of Vabout 100 is typical with the preferred catalysts.

Another means for designating suitable catalysts for the first reactor is the polymerization activity index. This is another arbitrary index which equals the `percentage of isoprene that is polymerized when 25 cc. of a mixture of 8-l0% isoprene in benzene is heated with 5 cc. of t-he monomer due to polymerization is calculated 4by difference between the isoprene content of the reaction product -and that of the test blend charged. A satisfactory catalyst for the first reactor has a polymerization activity index less than about 35, as polymerization there is undesirable.

An arbitrary desulfurization activity index may also be used in this modification of the present invention, princi pally for determining suitable catalysts for the subsequent desulfurization operation. This index is the percent reduction in sulfur content obtained when a blend of pure compounds consisting of 10% hexene and 10% isoprene in 80 volume percent of benzene with a total thiophene sulfur content of 500 p.p.m. Vis passed over the catalyst in question at 500 F. and 450 p.s.i.g. together with between 1500 and 4000 s.c.f.b. of hydrogen at a liquid hourly volumetric space velocity of 2. For sui-table desulfurization the final stage catalyst desirably has a desulfurization activity index of at least about 80, both fresh and t after one week of operating with the test feed s-tock.` It has been observed that good catalysts for the initial hydrogenation zone also have desulfurization indexes in the;

80-100 range initially but these catalysts are quickly poisoned or deactivated by a sulfur content equivalent to that in the test feed so that after one week of operation the activity index -is in the 0 to 50 range.

Catalysts of substantial acid activity are not desirable for this particular process since they produce unwanted cracking reactions, so silica-alumina catalyst supports are usually avoided. However, While it is preferable that 'the catalyst lsupport is lsubstantially free of halogens, a relatively low halogen content up t0 about 0.5% may be tolerated. Furthermore, a Acatalyst is favored which is substantially devoid of alkylation activity and thus does not promotethe alkylation of aromatics with olefins.

A variety of catalysts of differing chemical constitution may be employed in the initial hydrogenation step as long `as they have the necessary Iactivity described herein. Platinium in amounts ranging from about 0.105 to 2.0%, preferably about 0.\2,1.0%, supported on various aluminas, and especially gamma and chi alumina, is suitable as 4are the other noble' materials in group VIII of the Periodic Table of Elements, such as rhodium land palladium. The concentration of palladium in such catalysts may be about 0.05-10% and about 0.2-2.0% is preferred for the purpose. NickelV either unsupported or on known supporting materials in concentrations ranging down to about 10% nickel in the composite catalyst Ialso provides satisfactory results, as does copper chromite. For instance, good hydrogenating characteristics for the yfirst reactor are obtained with 55% nickel supported on kieselguhr. genation activity at low temperature, palladium or platinum on gamma alumina are recommended, palladium being preferred for the initial reaction, since its greater By reason of their high hydro.

activity catalyzes the desiredhydrogenation.reactions at a temperature about 100 lower than in the case of platinum. The palladium composite is desirably Vpromoted in sorne instances with a quantity of chromia in lthe same range as the palladium. Among the many suitable specilic catalysts are 5% palladium on `activated'carbon and 0.6% platinum oneta alumina of less than 0.01% chlorine content. The manufacture of such catalysts is well known in the art and accordingly is not describedhere.

The catalyst or catalysts employed inthe second or a final reactor operate under quite different reaction con. ditions from Ithose in the initial reactor. The charge is enti-rely inithe kgaseous phase yand substantially higher.

temperatures are required in the second reactor to desul.

furize the intermediate product; and =to complete the.

hydrogenation of the less reactive unsaturated hydrocarbons, namely the: mono-oleiins and any remaining dioleiins. This catalyst may also be Vdefined in terms of an arbitrary desulfurization lactivity.indexaotivity asset forth hereinbefore in the -1100 range, which isfretained for a period of at least one week and usually. muc-h longer;

It should have at leastmodera-te hydrogenation activity. Y The increased temperature of -the fnalreaction greatlyV increases the actual hydrogenation activity of these catalysts. Also it hasbeen foundv thatfsome and perhaps The preferred catalyst Ifor theyfnal, stage is a sulfded.

composite of cob-alt molybdate on the surface of gamma alumina.

In illustration, the ypresulfidingfo-r finalstep in thepreparation of a preferred type of`desulfuriz-ation catalyst may desirably be performed in situ inthe reactor.

Afresh contact catalyst containing cobalt molybdate on` the surface of a suitablek support such as gamma alumina or a catalyst regenerated tothe Yoxide state byjcombustion with air diluted by steam isgsubjected first-to prereduction for six hours at*v 700 p.-s.i.g.` and 700ol with a hydrogen-rich recycle gas substantially free of hydrogen sulfide. Following ,the .prereduction step athe catalyst is then contacted with a circulating stream .of mixed hydrogen sulfide and hydrogen under conditions such that the minimum partial pressures are 8 p.s.i.a.y for hydrogen zsuliideand 'p.s.i.a. in the casebfhydrogen and the temperature is in the range of 500 to 700. ment is concluded at a temperature; of about 700 after being continued until the sulfur content of the content of the composite -catalyst rises to the range. 6.5-7.5%

whereupon the catalyst is ready to be placed on-strearn'.,V

Later during desulfurizing operations, the sulfur in the catalyst drops from that range to an equilibrium content;

of about4.6%

Returning now tothe first reactor, suitable ranges of reaction conditions for this preferred processhave been described earlier and the actual reaction conditions are selected and regulatedvwitliin those ranges in amanner' known to `those ,skilled inlthe art to produce an initial hydrogenation eiuent in which at lea-strabout 35 and preferablyr'at least .60%, of the original diolefins have been converted ,into mono-olefinsor p-arafn's andin whichV an yamouutequal to at least about`.20%, and prefeffect of one operating Ivariable upon another are well understood by those skilled in the art and need not be explained in detail here. For instance, if the Vdegree ofy The treatv hydrogenation tends to drop Ibelow the minimum specified or perhaps below the preferred value, this condition can be corrected by increasing the feed temperature or decreasing ythe space velocity or both. Also if the proportion of initial reactor effluent in the liquid phase drops below the minimum specified, the feed temperature may be decreased, the space velocity increased to reduce the total exothermic lheat generated and provide a greater quantity of reactants to absorb the heat liberated, or t-he pressure increased or any combination of these measures may be employed in reducing the degree of vaporization in the initial reactor. Using the same circulating gas, an increase in total pressures of course results in a corresponding increase in hydrogen partial pressure.

With a fresh catalyst, either nevi or regenerated, it is obviously most economical to maintain the feed or charging temperature at the lowest temperature at which the gaseous and liquid components of the charge are readily available thus avoiding any heating or cooling expense. The feed temperature should of course be within the stated range and, in the preferred operation, the feed temperature is maintained at a substantially constant value within the narrow range of 75-l90 F., and desirably in the lower part of that range, While the catalyst is fresh. Generally, this temperature is subsequently increased either gradually or by steps but not beyond about 300 F. in order to maintain a diolefin saturation above the specified minimum, and preferably substantially constant degree of saturation, as the hydrogenation activity of the catalyst decreases with continued use. Even with a fresh catalyst, the first hydrogen treatment customarily does not fully saturate all of the olefnic or unsaturated aliphatic compounds for the Bromine Number reduction usually is in the range of about 25-95%.

Regeneration of the initial hydrogenation zone catalyst is required when the Diene Number reduction is less than the prescribed minimum of 35%, or the degree of vaporization exceeds 80%, or both, even after the feed temperature has been adjusted upward to the stated maximum. These are better criteria than prescribing a maximum outlet temperature for the initial reactor inasmuch as the degree of vaporization of-the efiiuent and the degree of saturation of its more reactive original components, are more significant than the outlet temperature in the instant embodiment of the novel process. In addition, it appears that the maximum permissible outlet temperature can vary considerably for different feed stocks over the range of about 275 to 400. For instance, a reactor outlet temperature of 325 is considered excessive for certain low boiling feed stocks but will give satisfactory results with other feeds boiling at higher temperature ranging up to end points near 500'.

An attempt to define acceptable outlet temperatures is not feasible in consideration of the variations in permissible pressures in the reactor. For example, an outlet temperature of 325 would be suitable for a relatively high reaction pressure in retaining the necessary proportion of liquid phase efiiuent while a lower temperature would be necessary if the minimum pressure were employed while all other conditions were held constant. Thus as a result of the close interrelation of the various operating conditions, it is more significant to define the initial hydrogenation in terms of the regulation of certain reaction conditions within restricted ranges to provide an intermediate product in which a certain proportion is retained in the liquid phase and a certain amount of the more reactive feed components are at least partially saturated.

An entirely different situation prevails at the outlet of the desulfurization reaction zone as it is unlikely that any exotherm created by the stated reaction conditions can reach a temperature sufficiently high to deactivate the catalyst. To permit the use of ordinary construction materials, the maximum outlet temperature should not exceed about 850 Although catalysts in the form of palladium or platinum supported on alumina retain their activity for extremely long periods, as for instance, 3 months or more in the Case of palladium catalysts, regeneration of the catalyst is eventually necessary and this may be readily accomplished by heating the reactor to a temperature of about 700-900 for a palladium-alumina bed while passing a gas containing l or 2% oxygen therethrough. A diluent is usually introduced with the air to avoid excessive regeneration temperatures which can reduce catalyst activity considerably. Nitrogen or flue gas may be used generally for that purpose and the more convenient medium of steam may be utilized as the diluent with a palladium catalyst.

The desulfurization or final stage catalyst is conventionally regenerated in similar fashion at even longer intervals of about 6 months or more. In the case of a sulfided composite of cobalt and molybdenum on alumina, this converts the cobalt and molybdenum compounds to oxides and a presulfiding treatment such as the one described hereinbefore is employed to restore the catalyst to its original form.

It has also been found that purging the initial catalyst with hydrogen at 200-500 p.s.i.a. and 750-850" for 16-4 hours sometimes serves to regenerate certain catalysts, such as palladium, almost as effectively as conventional regeneration by combustion with air diluted to an oxygen content of 1 or 2 percent. Accordingly, it is contemplated that, in the absence of severe deactivation of the catalyst, this catalyst may be regenerated several times by such treatment with hydrogen-rich gas before it is necessary to regenerate the contact agent by the combustion technique.

Only a limited amount of hydrogen sulfide may be tolerated without substantial deactivation by certain of the catalyst suitable for the first reaction stage of this species of the invention. Although this loss of activity may be readily restored either by regeneration of the catalyst in the usual fashion or the hot hydrogen treatment described earlier, frequent regenerations reduce the over-all efficiency of the process. Accordingly, in the case of a platinum catalyst supported on alumina, it is desirable that the concentration of hydrogen sulfide in the gaseous phase should not exceed 0.05 p.s.i.a. and preferably should be less than 0.03 p.s.i.a. The effect on a palladium catalyst is similar. Organic sulfur generally has a lesser effect on the catalyst and it is a relatively simple matter to control the hydrogen sulfide which is introduced in the hydrogen-containing gas by simply passing either or both of the make-up and recycle gases through an alkaline scrubber, or other unit for removing hydrogen sulfide such as a diethylamine absorber.

Under severe conversion conditions, for example a high desulfurization temperature in combination with a low space velocity of perhaps less than 1, a sulfided composite of cobalt and molybdenum on alumina may catalyze the hydrogenation of a part of the aromatic hydrocarbons, as exemplified by the conversion of benzene to cyclohexane. This is usually undesirable and may be easily avoided by inhibiting the reaction by maintaining a concentration of sulfur compounds in the charge equivalent in inhibiting effect to at least about 50 p.p.m. of thiophene sulfur (e.g., about 20 p.p.m. of hydrogen sulfide). Where the charge contains less of such compounds it is a simple matter to supply additional hydrogen sulfide in the hydrogen-rich gas which is introduced upstream of the final reactor. Selection of a make-up gas of suitable hydrogen sulfde content or by-passing the recycle gas around the caustic soda scrubber are some of the methods useful in attaining any additional inhibiting effect.

Despite the unstable nature of the hydrocarbon feed stock, especially when subjected to substantially complete vaporization, very little if any gum is formed in the first reactor of this particular process. The relatively low reaction temperature is not conducive to thermal polymerization. A catalyst having little or no polymerization activity is employed. A substantial proportion of the reaction mixture is maintained in the liquid phase to avoid approaching the point of dryness in the reactor. In addition, the usually substantial aromatic content of this liquid'makes it a good solvent for polymeric gums, so the though this technique has been suggested in the prior art.

Such procedure deposits polymer either in the heater or in the catalyst mass or both, and stoppages of this nature call for much cleaning and/or regeneration that reduce the overall operating efiiciency. Instead the instant process is concerned with vaporization of the initial eiuent in the presence of a flux liquid. This may be accomplished by various methods, one of which involves a combination of stages in which the initial hydrogenation effluent is gradually heated undei good temperature control in the presence of a flux, preferably circulating in substantial quantity through the transfer line between the` initial reactor and a vaporizing and separating chamber of enlarged cross section. In that chamber vaporization of the initial original feed and products thereof is completed to the desired extent of about 90 to 99% and seldom more than 99.5%. The small but significant balance of unvaporized effluent is withdrawn from the process as a liquid leaving the enlarged chamber and carrying a small amount of polymer formedrduring the vaporization operation and possibly also in the initial hydrogenation step or perhaps present in the original charge stock. Once this separation of the gaseous and liquid phases has been accomplished, there is no longer a tendency toward any significant polymerization in the gaseous phase containing the major proportion of the hydrocarbons even when it is heated up to temperatures of 350 to 700 which would have produced an unacceptable degree of polymerization in the mixed phase material from the initial reactor. Y

The gradual heating of the initial effluent to effect controlled vaporization during passage of the eluent through the restricted transfer conduit (including heater passages, etc.) leading from the initial reactor to the vaporizing and separating chamber may be accomplished by several means. One, comprises an optional but preferred technique in which a circulating liquid flux at a substantially higher temperature, than the reffluent typically of the order of 75-200 higher, is injected into the initial hydrogenation effluent near the outlet of the first reactor. It will be appreciated that the exotherm of the initial reaction has already increased the temperature of -this effluent substantially above the temperature of the feed to that reactor. The temperature of the mixture of flux and reaction effluent is preferably increased further during passage through an indirect heater which is desirably heated with steam or another easily controllable medium for even heating. A relatively low'temperature difference between the heating and the heated media is highly desirable to provide the gentle heating that minimizes polymerization in such equipment. Indirect heat exchange is recommended for the major heat input into the stream passing through the transfer conduit. In one embodiment of the invention, an amount of the initial reaction efluent equal to betweenabout and 50% 0f the original liquid feed may be retained in the liquid phase in this controlled heating operation whenifurther Vaporization is' subsequently produced by the injection into the heated effluent of a stable gaseous reactant at a higher temperature as described hereinafter. Finally, and preferably closely adjacent to the inlet of the separator,

Accomplishing 12 an additional stream of the hydrogen-rich gas used in this process may be injected into the mixture at a temperature several hundred degrees higher than-the temperature of the mixture.Vv Thisdirect contact heating with jet of hot gases is an optional but highly 'desirable feature which minimizes polymer' deposition on equipment surfaces. With each of these increments of heat,l more of the lirst reactor elluent is converted in the transfer conduit from the liquid phase into the gaseous -state under conditions in which the presence at allrtimes of a substantial liquid phase assistsin preventing or at least in `minimizing the deposition of polymeric material or heated.

surfaces. The enlarged crosssectionl of the chamber provides good conditions for separating the two-phases bycatch any traces of entrained liquid in the rising vapors.

The supply of steam to the indirect heater may be manually controlled vto maintain a predetermined temperature in the vaporizing chamber as steady as possible, but far better :results are usually obtainable in regulating the steam supply in response to the liquid level in `the. separating chamber.

In brief vsuch a regulating system involves. controlling the input of steam manually, but preferably automatically, in direct response tothe signals of a conventional liquid level indicator or: controller attached to theV vaporizing and separating chamber. The removaly of liquid streams from that chamber Ias well as any input ofV external flux is desirably maintained at constant flow rates under the` regulation ofv automatic flow controllers; therefore, a rise in the liquid level in the separating chamber represents a decrease in the vaporization of the `initial hydrogenation.

effluent and a fall in that level-means=that the eliiuent is being vaporized in a greater degree. To maintain a steady degree of vvaporization more steam or less steam respectively is supplied to the indirectheater. The heating steam may be adjusted Iby means of a valve'in :the steam supply line or one in the line used for draining condensed heating steam from the heater.

Conventionally, control 'of vaporization of a generally similar nature is regulated in response tothe temperature of the vapor or perhaps the liquid temperature., Such control is subject to the usual deviations encountered in efforts to obtain `precise elevated temperature measurements thatarise from radiation or evaporization-of liquid on a temperature sensing element, etc. not particularly satisfactory for liquids of narrow boil-ing range, such as the preferred feedstof the present invention,l

inasmuch as a small temperaturey differential of only a few degrees at a substantially elevated temperature generally is related to a largeV diiferential in the proportion Vof liquid vapor-ized.: Thus control of f heating of the liquid in direct response tothe actualproportion of unreacted feed Either manual or automatic eontrolof'the heating of 1 the initial eflluent in direct -responseto the liquid level in the flash chamber `may also; be extended-to controlling the quantity of heat supplied by the stream .of hot hydrogen-v rich gas injected into the transfer line near the inlet of the vaporizer pot'.

line or on the temperature at the jcharge outlet of the furnace ldescribed hereinafter for heating that gas. Also,

it is possible to control both the heat input to the indirect heater through which t-he initial Veffluent passes yand the heat furnished to the effluent by the: hot hydrogen-rich stream in response to the .liquid level controller on the vaporizing andV separating chamber,

Moreover, it is This regulation may be `exercised eitherl on the quantity of said gas being admitted to the transfer Y However, it is; usually preferred from a standpoint of practical opera,

13 tions t-o apply such regulation only to the steam input to the indirect heater.

The liux liquid comprising the liquid fraction of the effluent from the initial reactor and any inert liquid miscible therewith that is introduced into the transfer line may perform several functions before being separated from the gaseous portion of that eluent in the separation chamber. It minimizes or inhibits gum format-ion at this critical stage of the preferred process wherein a stream of mixed gaseous and liquid lhydrocarbons containing gum-forming precursors is carried to a relatively high degree of vaporization by heating, for the flux prevents the effluent from approaching dryness too closely, for example, not closer than about based on the original liquid feed rate. Secondly, the circulating flux serves as an economical and relatively gentle direct heating medium for vaporizing a portion of the initial effluent. Finally, t-he flux liquid prevents, or at least minimizes, t-he deposition of any gums or polymeric solids on the pipes and other apparatus by reason of its washing -action on the surfaces thereof and its solvent characteristics which enable it t-o retain in solution any polymeric material Whether formed at t-his stage or earlier.

Although any hydrocarbon liquid of suitable boiling and stability characteristics may be employed as the flux in this particular embodiment of the new process, it is preferred that the content of aromatic compounds should amount to lat least 15% to improve its capability for dissolving gummy material. A flux liquid from an external source may be used, and it is suggested that its volatility should be suficiently low that a major proportion and preferably substantially all of the iiux remains in the liquid state under the conditions in the vaporizing chamber while its resistance to coking and polymerization should desirably be at least as good as that of the initial efuent. Its boil- -ing range is preferably located between about the boiling point of benzene and about 950. However, an economical and readily available ux liquid may be obtained very simply by merely reducing the -heating in the transfer line between the first reactor Iand vaporizer pot at the start of a run in order to retain a larger than usual proportion of the initial reactor etiluent in the liquid phase until a suicient body of flux liquid has been built up in the system. This of course amounts to :accumulating the least volatile fraction of the feed stock as the liquid liux.

The rate of recirculating the flux liquid in the preferred :process may `amount to at least 5%, and preferably at least 10%, of the rate of introducing the liquid feed stock into the frst reactor, and lesser amounts may be Irecirculated where an appreciable proportion of the initial effluent is retained in the liquid-phase throughout the vaporizing step. As used herein, all flux (liquid efllue-nt plus any added liquid) quantities or rates relate to the proportions at the moment when the maximum degree of vaporization of the initial eiiluent is attained; and, of course, the proportion 'of material in the liquid phase reaches its minimum--namely, the instant of separation of the gaseous and liquid phases-rather than at the confluence of a circulating tlux stream with the initial hydrogenation effluent. Much higher flux circulating rates can be employed ranging up to 40%, and even to 200% or more, for the only lreal limitations are physical ones relating to the capacities of the equipment and economic ones relating to pumping costs and the cost of larger equipment. When the total proportion of liquid in the transfer line and heater is ample by a substantial margin to avoid dryness and bathe the walls of the equipment, furthe-r increases in the flux circulation rate do not achieve a corresponding or even a significant reduction in the amount of polymer formed in the system or even in the polymer concentration in the ux liquid; hence, circulation rates provide no important advantages.

The concentration of polymer in the circulating flux is dependent on the small but significant proportion of spent ilux withdrawn from the vaporizing step and from the instant process either intermittently or preferably continuously. This spent liquid is derived from an unvaporized fraction of the eiuent of the initial reactor or from a supply of external flux or from both sources, and over any substantial periods the rate of withdrawal must equal the supply from these sources. Under the preferred steady state conditions, reducing the degree Iof vaporization of the initial eflluent and correspondingly increasing the spent flux withdrawal results in a decrease in the polymer concentration in the circulating ux and vice-versa. As indicated earlier, this removal of spent flux liquid amounts to at least about 0.5% and desirably about 1 to 10% based on the liquid feed rate. While the amount may be larger, it is generally uneconomical to withdraw much more in the liquid phase for purification or further processing. In actual practice of the instant embodiment a llow controller on the spent flux line from the vaporizing chamber may be adjusted manually as needed to keep the gum content of the circulating liquid low enough to avoid the deposition of polymerio material in the equipment; for example, by keeping the gum content below about 200 milligrams per milliliters.

Where a flux liquid from an external source is supplied to the system at a constant and usually relatively low rate, it is possible to vaporize a correspondingly greater proportion, in fact the whole, of the liquid eiiiuent fraction yof the first reactor. However, it is preferable to retain the least volatile 0.5% or 1% of said effluent in the liquid state in order to keep the temperature as low as possible during the vaporizing operation. For example, with all percentages based on `the liquid feed rate, one may continually charge 5% of a hydrocarbon -oil having an atmospheric boiling range of 600700 and a major proportion of aromatic hydrocarbons to the separating chamber as circulating flux, and recycle 25% liquid from this pot to the transfer line immediately downstream of the first reactor; then vaporization of the efflunt-flux mixture in the transfer line may be controlled by appropriate heating to retain 1% of the initial efuent in the liquid phase in that chamber and 6% spent llux may be withdrawn continually from the bottom thereof in maintaining steady operations.

The size and shape of the separating and vaporizing chamber are not critical. In avoiding or minimizing appreciable entrainment of liquid droplets in the vaporous phase that is leaving, it is desirable to keep the velocity of the gaseous phase relatively low, perhaps 2 feet/ second or less. This can be achieved by providing a reasonably large cross-sectional area perpendicular to the direction of gas flow in the upper part of the vessel. On the other hand, where the heat for vaporization is regulated in response to liquid level in the chamber, it is desirable to have a relatively small cross-sectional area in the neighborhood of that level in order that a significant change in level will occur whenever a significant change in the degree of vaporization of initial effluent occurs. Such factors pose no great problems, as there is no necessity `for maintaining a constant crosssectional area throughout the length tof the vessel. As one illustration, the vessel may be in the general form of a double cylinder having a lower section of considerably smaller diameter than the upper section.

After separation of the flux liquid from gaseous material derived from the efiluent of the initial reactor, this gaseous phase is heated if necessary to bring its temperature up to the desired inlet temperature lof the second reactor and its proportion of hydrogen is boosted, if necessary, to the desired level for that reactor by the introduction of a hydrogen-rich gas. These steps may be combined, if so desired, by introducing the extra hydrogen-containing gas at a substantially greater temperature, say about 100 to 400 more, than that of the gaseous phase leaving the separating chamber. This is one of the suitablemethods of making the final temperature adjustment in the charge to the nal reactor. It is preferably accomplished by regulating the volume of fuel gas burning in a furnace for heating circulating gas and consequently the outlet temperature of that circulating gas stream either manually or automatically in response to `signals from a temperature sensing device located in the conduit leading to the inlet of the second reactor.

For a better understanding of the nature and objects of this invention, reference should be had Ito the detailed description and examples hereinafter taken in conjunction with the accompanying drawing which is a simplied flow sheet or schematic representation of the process of the present invention. It will ybe appreciated that many details well known to petroleum engineers have been omitted from the drawing or simplifiedfor simplicity and greater clarity including Ipumps,- valves, alternate and parallel piping and equipment andV control equipment, especially instruments for indicating, recording or regulating temperature, pressure, level, flow, etc.

EXAMPLE 1 Turning now to the drawing, a freshly-distilled stream of thermally cracked and depentanized gasoline (170- 220 B. R.) of the composition set forth in part in column 1 of Table I hereinafter enters the feed conduit Z' at ambient temperature and a pressure of 740 p.s.i.g.

for catalytically hydrodesulfurizing gas oil is admitted iny pipe 4 at a pressure of 750 p.s.i.g. The Vquantity and` composition of this gas are specified in column 2 of Table I. This make-up gas joins the recycle gas stream, whichy lis described later, in conduit 6. The resulting mixture has a temperature of 125 F. and its composition and rate of flow are designated in column 3 of the table.

Half of the mixed gas stream in conduit 6 is taken oif in the valved line 8 for purposes that willbe apparent later. The other half of the gaseous material continues to travel along pipe 6 until it joins the hydrolysis liquid hydrocarbons in conduit 2, and this gas-liquid mixture of the composition and ilow rate given in column 4 of Table I passes through the heater 10 where its temperature is adjusted to 115 (herein designated as the feed temperature) by heating, if necessary, on its Way to reactor 12. This charge temperature produces good results with the catalyst described hereinbefore which has been partially deactivated by a liquid feed containing undesired polymeric materials.

Column 4 of Table I setsforth the total charge to the rst or initial reactor 12 which contains a xed or stationary catalytic bed of chromia-promoted palladium on a gamma alumina support in the form of 3/16 diameter i cylinders 5716 long. Based on the total weight, there is a surface deposit on the alumina of 0.50 percent of palladiurn metal and 0.51 percent of chromium in the form of oxides.

The reaction conditions in the first reactor 12` are:

In the rst stage reaction the primary reactionisone of the nondestructive hydrogenation ofk diolefnsespecially conjugated diolens, accompanied by considerably less t saturation of the less reactive mono-olens. The temperat tures are below the levelrequi'red for desulfurization and no significanthydrogenation vof aromatics or polymerization takesplace there.

Any trace of gum formed in the catalyst bed dissolves in the descending liquid and the reaction eluent is drawn off at the bottom of the reactor viaV conduit 14 .in'which it is transported to heater 15.' vA minor portion of the liquid feed stock or reaction products thereof vaporizes in reactor 12 'as a result ofthe heat evolved in the exothermic hydrogenation lreaction.V

A circulating iiux liquid at 350 is injected from the conduitl16 into the products in pipe 14 partly to increase the temperature ,of the initial reactoreluentabout 45 thus promoting its vaporization but chiey yto reduce any tendency toward the deposition I.of any gummy 'solidsfin the transfer` line 14. This flux liquid isJdraWn off near.

the bottom of the vaporizer pot 18in pipe 16 and recirculated by pump 22.at the rate of 9,220,1bs./l1r.`0r720 b./d. This liquid is composed of thehigherY boiling hy,-

drocarbons of the initial reactor eflluent' which'. are -retained in the liquid phase and a small quantity of dissolvedv polymeric material. The latter is a by-product of the present process and is readily soluble in the benzene and other aromatic hydrocarbons constituting most of the liquid flux.

Two other modes of heating the rst reaction euent` are also employed during its passage to `thevaporizervpotv 18.l Saturated steam at 220-p.s.i.g.. is admittedV to the heater 15 ,unde'ra control technique described Ihereinafter to indirectly heat the first reaction products to a tempora-l -ture of 337. In addition, a heated hydrogen-rich gaseous mixture is injected into ,the products in conduit 20 upstream Vbut close to the `chamber 18. rich stream is part of that drawn off in line8 fromthe total circulating gas (recycle and make-up gases) in conduit 6. The gas in pipe 8 flows through the heat exchange-r 24, where its tempe-rature is raised to 380, and finally into gas-fired heater 26. Firing of v this heater is controlled in a unique manner which is described'later; and it provides an effluent lleavingfinv conduit 28 -at a temperature of 645 which is divided by means of the three-way valve 30 with 20% of the rtotal circulating gas t being introduced into pipey 32 and -the lremaining' 30% passing through conduit 34 to join the rst reaction effluent in line 20. This fu-rther heating of the product stream in line 20 of course results inmore vaporization and j vaporization is completed to the desired extent'in'vapo-V, f yrizer 18. The latter is a vesselof enlarged cross section with an internal diameter of 4.5 feet and a height of 12.5 feet which provides'favorable conditions for the substantially complete separation of the Vgaseous phase fromV the.A

liquid phase in a mixture thereof at a temperature of1360 and pressure of 695 p.s.i.g.

Based on the vrate ofy feeding pyrolysisv gasoline, 4%

of Vsaid liquid feed and reaction products thereof vaporizes inreactor 12,:much more is evaporated during passage 1 through line 14 and heater 15,1further substantial vaporization is produced bythe hot gas Vinjected from vpipe 34 i` and only 8.5% is collected in .the liquid :phase in the.`

vaporizer pot '18 inV addition to they circulating flux.

The gaseous phase going'overhead passes throughv the demister blanket or pad 36 of coarse steel wool designedV to catch any entrained droplets ofliquid. No substantial deposition of polymers or gums occurs in the line 14 and., 20 or heater 15, but the .liquid in the bottom of pot 18 i contains an amount of dissolved'polymer; (ASTM gum content=76 mg./ 100 ml.) which issmallV but sutlcienty to foul and thereby deactivate a contactr catalyst within a fairly, short'time at desulfurizing temperatures. A portion of the flux is continually beingremoved ata constant rate of 200 b./d. as spent ux through the bottom line 38 under the regulation of the flow controller 40Y This hydrogenoperating the automatic valve 42. The rate of Withdrawing spent flux from the system is manually reset from time to time to the minimum rate that will hold the gum content thereof below about 100 milligrams per 100 18 The tiring of the furnace 26 for heating hydrogen-rich circulating gas is controlled by the automatic valve 64 operating in the fuel gas supply line 66 in response to two temperature controllers. Temperature controller 68 mls. The spent liquid ux is transferred to arerun tower 5 Senses the temperature in the outlet line 28 from the (not shown). heater and maintains a temperature 645 at this point,

While an extraneous flux, may be alternatively supbut this device is reset to other temperatures as may be plied to the system at a constant rate through the line 50 required in response to the temperature controller 70 connected to vaporizer 18, a suitable flux is obtained which is connected to conduit 56 and maintains a temfrOm the effluent f the iirst reactor by temporarily op- 10 perature of 515 in the charge entering the second reactor. erating heater 15 in the manner described hereinbefore Tb@ SeCOHd reaCiOr 58 COIltainS a bed 0f a COmpOSlte to accumulate sufficient liquid in vaporizer pot 18 for Of Cobalt and molybdenum suldes on a gamma alumina recycling as a flux; and thereafter normal operating con- 0f L9/10 inch Particle Sile Prepared by bydfogesuliide dtiOnS are employed in the vaporizing systeml The treatment in the manner described hereinbefore with a overhead or vapor phase passes through heat exchanger l Sulfur Content 0f 46% ai Operating equilibrium and a 52 0n its Way from vaporizing chamber 18 Via Conduit Weight IalllO Of A12O3ZMOICO Of 84.717.922] I'eSpClVely. 54 t0 join the hydrogenqich gas from pipe 32 in line 56 In the initial reaction eluent the remaining less reas the Charge for the second Stage reactor 58 In this active diolelins are saturated in the second reactor along passage, the heat exchanger 52 raises the temperature of With all of the mono-olefin in said eliiuent in a nondethe overhead e-iuent to 485 and admixmfe with the 20 structive manner with no substantial saturation of arohydrogen-rich gas at about 645 further raises the temmatic compounds. The reaction conditions in the secperature of the total charge to 515 at the reactor inlet. 0nd Stage reileiof are 3S fOiiOWS As indicated previously, two unique temperature con- Inlet mpemture 515' trol techniques are employed for heating and thereby Outlet temperature 535 vaporizing liquid effluent from the rst reactor to pi'e- 25 Total pressure 685 p Si'g pare a vapor phase charge for the second reactor. First, H2 partial pressure 335 psm: the rate of il-ow of heating steam .through conduit 59 to Total H2 charged 3350 S c f b heater is controlled by automatic valve 60 in response Space Velocity calculated to an external liquid level controller 62 which .is conon liquid feed 1 7 v /hL/vnected 1n conventional manner to sense the liquid level 30 catalysts activity indexes: in separator pot 18. Since the rates of circulation of Desulfurization 93400, iiux liquid and removal of the spent flux are customarily Polymerization 43, held constant, a rise in the level of liquid in pot 18 indicates that the liquid feed stock and its liquid products lrom the I nlet and Outlet tempfratures gwen I t 1S aP' are being vaporized at a lower rate. This is corrected palm. that Significant hydrqgenatmn ractlons Wlth Sub` automatically by the level controller 62 generating a Stalmd exotherms are mkmg .place m both reactqrs' This is borne out by a comparison of the unsaturation function or signal in response to which valve autoindexes of the reactor charges and etlluents of column 4 mil/cally opens to admlt more Steam Into heater 15 nd with column 5 in Table I and also column 6 with 7. The thus Vaporlze more of the rst reactor @gluem passing 40 latter two indicate that aminor hydrogenation of diolefins hrough the ilef 15' Corfvesel 2 a fau m hquld level is completed iri the iinal reactor along with the principal 111 the VaPOUZmg chamber mdlcate that a greater Pro' hydrogenation that saturates substantially all mono-olens Portion iS being VaPOIiZed, and this 1S Corrected by a and a substantially complete hydrodesulfurization of orsignal from the level controller 62 to the automatic valve game sulfur Compounds. Again there is no appreciable 6) which reduces the steam input to heater 15, and there- 45 polymerization or deposition of coke and no noticeable fore results in a lower rate of vap-orization in the liquid conversion of aromatic hydrocarbons to naphthenes ocpassing therethrough. curs.

Table I i 2 a 4 5 6 7 8 9 i0 Stream Fresh Total First First Final Final .1. Stabi- Stabi- Gasol HrRich H2-Rich Reactor Reactor Reactor Reactor Sep Oil lizer lized Feed Gas Gas Charge Effluent Charge Effluent as Ot Gas Liquid Flowlbsyhom 28,220 2, 520 14,020 35,230 35,230 39,080 39, 680 1,860 920 25,400 Flow, CHIII.. 42 r13 45 240 19 er Hour:

Mol 961. 0 480. 5 C1 501. 2 250. 6 C2 68. 3 34. 2 Il S in .111. C? p p 10.1 5.1

. l. 3 0. 7 C5 Saturates. 2.3 1.2 C5 Unsatuiates 3. 5 C@ Saturates 3.1 31a C@ Unsaturates 30. 0 Benzene 13. 6 277. 2 Toluene 0. 1 3. 6 Other 01+ 15. 3

Total, Mo1s.[Hr 3 1, 560. 9 1, 134. 2

1 Measured under actual process conditions. 2Based on weight of original liquid feed.

The gaseous product stream leaves the bottom of reactor 58 via conduit 74 and is cooled by passing through heat exchangers 52 and 24 respectively, as well as the cooler 76, on its way to the high pressure separator '78 where the vapor phase is separated from the newly condensed liquid at a temperature of 100 and pressure of 640 p.s.i.g. From this vessel the gaseous phase is taken overheadvin lines 80 and 82. About 15% of this gas is bled olf to the refinery fuel system through pipe 84and the pressure regulator 86 which maintains the desired pressure on the hydrogenation system. The rate of removal of this separator gas from the instant system is tabulated in column 8 of Table I. Most of the gaseous material, however, enters the line 90 wherein it meets f with any make-up gas diverted from supply conduit 4 via valved line 92 that also may require scrubbing to remove excessive hydrogen sulfide. These gases are introduced into the lower half of the combination washer 94 which is equipped with a lower caustic scrubber section 96 bey neath a water washing section 98.

Fresh aqueous sodium hydroxide solution is admitted f in conduit 100 and joins recirculating caustic soda soluf tion in the line 102 on its way to the perforated scrubber trays over which it cascades downwardly against the rising 1 gases. conduit 106 at the bottom and divided between an exit line 108 for spent solution and conduit Htl-leading to This alkaline liquid is drawn ol through the the water is collected in the trough 118 and is not allowed i.

to descend therebelow and dilute the caustic scrubbing solution. The gases rising countercurrently through the tower 94 at 640 p.s.i.g. lose most of their hydrogen sulfide content in being scrubbed irst by intimate contact with curtains of caustic soda solution, next they pass through the demisting pad 120 intothe washing section where they are washed with curtains of falling water to remove the last traces of H2S as well as any entrained particles of the caustic soda solution and then through the demisting pad 122.

The scrubbed and washed gases exit through the conduit 124 which connects with the valved by-pass line 126, that may be used to divert some or all of the separator olf-gas around tower 94. The by-pass conduitfis useful when the hydrogen sulfide content of the separator gas is low enough for a recycle gas.

These two pipes feed into the line 128 which leads to the knockout pot 130 in which any entrained liquidris separated. From here the hydrogen-rich gas passes through conduit 132 to compressor 134 where its pressure is boosted sufliciently to circulate it'through the recycle gas line 6 and associated conduits in the manner described earlier.

Returning n-ow to the scrubber 94, it is apparent that an extremely exible arrangement is shown for controlling the hydrogen sulfide content of the circulating gases passed into the two reactors with the feed. For example, the operator can divide the gaseous product from separator 7S between inlet line 90 of the caustic scrubber and the by-pass conduit 126 in any desired proportions. Similarly, the make-up gas entering in conduit 4 can be introduced directly into circulating gas line 6 or part or all of it can be taken off via conduit 92 for treatment in the caustic scrubber. Also, either or both of the rates of recirculation of caustic soda solutions in scrubber section 96 and the introduction of fresh caustic soda thereto can be controlled to set the rate of reaction and removal of hydrogen sulfide from the gas stream passing through the tower. Y

The liquid phase withdrawn from the bottom of high pressure separator 78 is treated in the stabilizing tower- 136 -at 180 p.s.i;g. after being carried in the conduit 138 through the pressure reducing valve 140 and heat exchanger 142 in which the temperature of the stream is raised to 240 F. Attached to the 30-tray stabilizer are valved inlet lines 144 to 146 =to introducethe charge selectivelyk and in any proportions onto the 18th and 12th trays respectively counting from the bottomof the column. A reboiler 148'is provided to maintain the bottoms at a temperature of aboutifSSfA F. and a stable, substantially saturated liquid product is withdrawn as the product of the process via pipe 150 into heat exchanger 142 at the rate given in column 10 of the table. This liquid, rich in aromatic hydrocarbons, is suitable for extraction processes, such as extraction with diethylene glycols, for the concentration of aromatics by reason of its Y negligible Ycontent of diolens, olefns and sulfur. It is essentially a mixture of'paraliinicrand ,aromatic hydrocarbons, and a sharp separation can kreadily be obtained between these constituents.

An-overhead fraction is conveyed via the conduit 152 and cooler 154 in which cold waterreduces its temperature from295 to 125 in transit to the reflux yaccumulator 156. Liquid reuxis returnedfrom the Vbottom of this accumulator to theV stabilizer 13.61through line 158v and pump 160 at a Vrate of 9840 lbs. per hour and-'aV gaseous by-product of the process is withdrawn through the valved conduit 162 atfrthe rate set forth in column 9 of Table I for use as fuel or other suitable purposes.

Starting up the process described herein in a commercial plant is relatively free of difhculties. Make-up gas obtained vfrom a catalytic reformer s charged. at; ambient temperature and the usual operatingl partial pressure .of hydrogen into the initial reactor 12 and also through the furnace 26 into the final reactor 58. This is vcontinued until the heat carried by theV gas from the furnace brings the second reactor close to its normal operating temperature. Meanwhile, recycle gas is substituted for most of the f resh supply ofl hydrogen-rich gas after a thorough purging of the system. Next, a typical reformate derived from naphtha and relatively free from runsaturated aliphatic compounds is introduced as a temporary feed along with the` circulating gas. An unusually high proportion of liquidaccumulates in the flash chamber 18 from kthe time theV reformate is; first chargedv until they chamber reaches normal operating temperature and none 1s withdrawn throughthe spent fluxline at'first.y After the circulating ux system is allowed to fill up `withfthe llquid phase collecting in the Vaporizer chamber, liquid is ydrained olf in the spent flux line at an abnormally high:

EXAMPLE 2 The process of Example 1 isrepeated using the same feed, equipment and reaction conditions except as otherwise specified herein. The `catalysts -are of identical composition with those-describedin Example .1 .except that the data tabulated hereinafter represents steady state conditionsafter operating for only live days with fresh new catalysts in both reactors.

Table Il v OPERATING CONDITIONS Flow rates-b/d.:

Pyrolysis liquid charge -3060 First reactor outlet 29,00 Flux oil circulation 710 Spent ux v 74 Stabilizer feed 3090` Circulating gas (72.5% H2)-m.s.c.f./h.:

From .caustic and water scrubber 742 To first-stage (Pd) reactorA 341v To furnace 26 336- To fuel line 84 21 To transfer line 34 218 To transfer line 32 118 Make-up gas (76% H2) 98.8 Pressures-p.s.i.g.:

Second reactor, AT 4 Gas furnace outlet 28 630 Outlet of heater 350 Vaporizer pot inlet 356 Circulating ux oil in line 16 350 Space velocities-LHSV:

First reactor 2.88

Second reactor 2.24

In comparison with Example 1, the temperatures listed immediately above demonstrate a considerably larger temperature rise in passing through the first reactor and a considerably smaller exotherm in the second reactor. This means that a distinctly greater degree of hydrogenation is taking place in the iirst reactor and less in the second than is the case in Example 1 where a catalyst of somewhat impaired activity is used in the rst reactor.

The following properties of several streams are established by analyses:

Table III First Spent Stabilizer Characteristics Charge Reactor Flux B ottorns Effluent (Product) Gum, nig/100 ml 1. 5 53 Specific Gravity 0.828 0. 830 0.837 O. 824 Unsaturation:

Bromine N o 18. 2 2. 9 Nil Dione No 9. 9 Nil Nil Sulfur, p.p.rn 39 l. 1

The results in Table III for catalysts with an age of ve days amount to reductions of 100% in Diene Number and 84% in Bromine Number for the product of the first reactor based on the values for the feed stock. Reductions of 85 and 55% respectively after 20 days of operation, and 80 land 50% reductions respectively after a total of 75 days are obtained in operating the initial reactor under the same reaction conditions. It is estimated that the rst stage reactor can be operated for a period of at least 6 months before regeneration of the catalyst is likely to be required to maintain a diolen reduction of 50% while keeping the charge temperature below 250 F.

The detailed examples given hereinabove are intended only to illustrate the invention. It will be apparent to those skilled in the art that many other modications and variations may be made in the embodiments set forth in the examples without departing from the invention. For instance, standby units arranged in parallel With -alternate piping may be provided for all equipment that requires periodic regeneration or cleaning. For simplicity and to provide comparative results, catalysts manufactured in the same manner are employed in both of the detailed examples; but the invention is not limited to such catalysts for a wide variety of other known hydrogenation and desulfurization catalysts may be used insteadY Accordingly, the present invention is not to be considered as limited in any respect other than the recitals of the appended claims.

Certain features of the selective hydrogenation process disclosed hereinabove are also described, and claimed in concurrently filed application Serial No. 238,693 of Richard G. Graven et al.

What is claimed is:

1. In a process for the selective conversion of unstable liquids with a pronounced tendency to deposit solids upon heating which includes partially converting unstable compounds in a liquid. feed into more stable substances Within a confined initial reaction zone under conversion conditions in which a substantial portion of said feed is maintained in the liquid phase, vaporizing a substantial portion of the effluent liquid from the initial conversion reaction by controlled heating and passing the gaseous phase derived from said. initial eflluent through a confined conversion zone while substantially completing the conversion of unstable components of said feed; the improvement which comprises separating in an enlarged separation zone a liquid flux in an amount equal to at least about 0.5% of said liquid feed from said gaseous phase, withdrawing said liquid` flux at a substantially constant rate from a pool thereof maintained in said separation zone and regulating said controlled heating operation in direct response to the rate of collecting said liquid ux in said separation zone as determined from the level of said. pool.

2. The process according to claim 1 in which said controlled heating is effected during passage of said initial eflluent through a restricted transfer conduit.

3. A process according to claim 1l in which said liquid flux is separated in an amount equal to at least about 5% of said liquid` feed.

4. A process according to claim 1 in which a substantial portion of said liquid flux is initial effluent liquid.

5. A process according to claim 1 in which a major portion of said liquid iiux is a material which has been subjected at least to said initial conversion reaction.

6. A process according to claim 1 in which a substantially constant portion of said initial effluent is retained in the liquid. phase and separated from the gaseous phase thereof in said separation zone.

7. A process according to claim 1 in which the breadth of the boiling range of said liquid feed does not exceed about 150 on the Fahrenheit scale.

8. A process according to claim 1 in which the breath of the boiling range of said liquid feed does not exceed about on the Fahrenheit scale.

9. A process according to claim 1 in which said initial eilluent liquid is separated in said enlarged separation zone in an amount equal to between about 1 and 10% of said liquid feed.

10. A process according to claim 1l in which an indirect heat exchanger is employed in said controlled heating operation.

11. A process according to claim 1 in which a stable reactant in gaseous form is introduced into said transfer conduit immediately upstream of said. separation zone at a substantially higher temperature than that of said heated initial effluent.

12. A process according to claim 1 in which an amount of said initial effluent equal to between about 20 and 50% of said liquid feed is retained in the liquid phase in said controlled. heating operation and further vaporization of said heated initial effluent is produced by the injection therein upstream of said separation Zone of a stable gaseous reactant at a temperature substantially higher than that of said efliuent.

13. A process according to claim 1 in which a stable gaseous reactant is injected into said gaseous phase at a location between said separation zone and said conversion Zone at an injection temperature substantially higher than that of said gaseous phase, which injection temperature is regulated in response to the inlet temperature of said conversion zone to maintain said inlet temperature constant.

14. A process according to claim 1 in which a hot stable gaseous reactant is injected into said heated initial eflluent and into said gaseous phase both upstream and downstream, respectively, from said separation zone.

15. A process for the selective nondestructive hydrogenation of a liquid hydrocarbon feed boiling below about 500 F. and containingk aromatic hydrocarbons, and oleiins, dioleiins and sulfur compounds which comprises passing said feed substantially in the liquid .phase and hydrogen through an initial hydrogenation zone-in contact with a porous solid hydrogenation catalyst of,

high hydrogenation activity' and low polymerization activity while controlling hydrogenating conditions in said zone including hydrogen partial pressure within therange of about 200-800 p.s.i.g., hourly space velocity Within the range of about 0.2-l5.0 based on the volume of liquid feed, the hydrogen charge within the range ofy about 50045000 s.c.f.b. of liquid feed and feed temperature within the broad range of about 75-300 F., said s` fluent by controlled heating and separation of the gaseous phase of said hydrogenation eiiuent in an enlarged sepy aration zone in the presence of a liquid flux in an amount equal to at least about 5% of said liquid feed, withdrawing said liquid llux from a pool thereof maintained in said separation zone at a substantially constant rate equal to at least about-0.5% of said liquid feed, regulating said controlled heating operation in direct response to the ratey of collecting said liquid uX in said separation zone as determined from the level of said pool, passing hydrogen together with gaseous material derived from said vaporization step through a subsequent conversion zone in contact Withfa porous solid conversion catalyst of at least moderate hydrogenation activity and high desulfurization activity at a substantially higher average temperature than in said initial zone While controlling conversion conditions in said conversion zone, including hydrogen partial'pressure Within the range of about 200-800 p.si.g.,` hourly space velocity within the range of about 0.2-6.0 based on the volume of said liquid feed, the total hydrogen charge Within the range of about SOO-10,000 s.c.f.b.

of said liquid feed and inlet temperature Within the wide` rangeof about 350-700v F., said conversion conditions being regulated to produce a conversion eluent from VWithin the range of about -275. F.

Y 24; said conversion zone with anormally liquid fraction having a Bromine Number less than about 4 and an organic sulfur Vcontentbelow .20-ppm.

16. A process according yto claim "15 1in which said liquidjfeed is richiin aromatic hydrocarbons;` and boils 17.' A process according to claim 15 in which said initialqcatalyst. has a hydrogenationVVV activity index of at least about 40 and, a polymerization activity index less than about 35 and said conversion catalyst hasV a desulfurization activity index of at least about =80.

18. A, process according to Vclaim` 15, in which said vaporizing step is effected in the preseucef of 'a quantity tfnf liiquid ilux equal to atlerastV about 10% of said liquid 19; A process according ito claim y15 Vin which saidV vaporizing step is effected in the presence of a liquid ux l recirculating at a rateequal-to at ,least about'l0% of that of said liquid feed .and saidfiux comprises the less ing said initial zone and said separation zone and a minor portion of said flux is removed at least` intermittently from said recirculating flux to minimize the 'accumula-` tion of polymeric material in said flux. Y

21. A process according to claim.20 in whichlfurther` vaporization of fsaid hydrogenation eluent is produced by the injection therein upstream of said separation zone of hydrogen at a temperature substantially higher than that of said effluent.

References Cited by the Examiner UNITED STATES PATENTSv 2,952,625, 9/1960 Kelley et a1., v 208-254 3,051,647 8/1962 White 20S- 255 3,075,917 l/l963 Kronig et al. 208-255 3,108,947 10/1963 Stijntjes 20S- 255` DELBERT E. GANTZ, Primary Examiner.

ALPHONSO D. SULLIVAN, IExaminer. v

Claims (1)

1. IN A PROCESS FOR THE SELECTEOVE CONVERSION OF UNSTABLE LIQUIDS WITH A PRONOUNCED TENDENCY TO DEPOSIT SOLIDS UPON HEATING WHICH INCLUDES PARTIALLY CONVERTING UNSTABLE COMPOUNDS IN A LIQUID FEED INTO MORE STABLE SUBSTANCES WITHIN A CONFINED INITIAL REACTION ZONE UNDER CONVERSION CONDITIONS IN WHICH A SUBSTANTIAL PORTION OF SAID FEED IS MAINTAINED IN THE LIQUID PHASE, VAPORIZING A SUBSTANTIAL PORTION OF THE EFFLUENT LIQUID FROM THE INITIAL CONVERSION REACTION BY CONTROLLED HEATING AND PASSING THE GASEOUS PHASE DERIVED FROM SAID INITIAL EFFLUENT THROUGH A CONFINED CONVERSION ZONE WHILE SUBSTANTIALLY COMPLETING THE CONVERSION ZONE WHILE SUBSTANTIALLY COMPLETING THE CONVERSION OF UNSTABLE COMPONENTS OF SAID FEED; THE IMPROVEMENT WHICH COMPRISES SEPARATING IN AN ENLARGED SEPARATION ZONE A LIQUID FLUX IN AN AMOUNT EQUAL TO AT LEAST ABOUT 0.5% OF SAID LIUQID FEED FROM SAID GASEOUS PHASE, WITHDRAWING SAID LIQUID FLUX AT A SUBSTANTIALLY CONSTANT RATE FROM A POOL THEREOF MAINTAINED IN SAID SEPARATION ZONE AND REGULATING SAID CONTROLLED HEATING OPERATION INDIRECT RESPONSE TO THE RATE OF COLLECTING SAID LIQUID FLUX IN SAID SEPARATIN ZONE AS DETERMINED FROM THE LEVEL OF SAID POOL.

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