US3246951A

US3246951A - Recovery of hydrogen from refinery gas

Info

Publication number: US3246951A
Application number: US197987A
Authority: US
Inventors: Ramella Amilcare
Original assignee: Socony Mobil Oil Co Inc
Current assignee: ExxonMobil Oil Corp
Priority date: 1962-05-28
Filing date: 1962-05-28
Publication date: 1966-04-19
Anticipated expiration: 1983-04-19

Description

Aprll 19, 1966 A. RAMr-:LLA 3,246,951

RECOVERY OF HYDROGEN FROM REFINERY GAS Filed May 28, 1962 2 Sheets-Sheet 1 Amilcore Rcmello ATTO/BVD A. RAMELLA April 19, 1966 RECOVERY OF HYDROGEN FROM REFNERY GAS 2 Sheets-Sheet 2 Filed May 28, 1962 INVENTOR Amilcare Ramello ATM/wf] United States Patent 3,246,951 i RECOVERY F HYDRQGEN FROM REFINERY .GAS Amilcare Ramella, Woodbury, NJ., Yassigner to Socony Mobil Gil Company, Inc., a corporation of New York Filed May 28, 1962, Ser. No. 197,987 14 Claims. (Cl. 23-212) The present invention relates rto .the recovery `of hydrovgen of high purity, of at least 90% `by volume, from industrial gas streams comprising hydrogen and components not reactive with extracting .agent and, more particularly, to the recovery of hydrogen from refinery ygas streams containing at least 20 percent by volume .of hydrogen and comprising hydrogen and hydrocarbons with or without ammonia, hydrogen .sulfide and the like.

Reforming of hydrocarbons and particularly petroleum hydrocarbons involves dehydrogenation of naphthene, dehydrogenation of paraffins, dehydrocyclization of paraffins, isomerization of normal parains to iso-parans, and hydrocracking of paraliins. of reforming conditions the `greater the dilution -of hydrogen produced -in the dehydrogenation of naphthenes and the dehydrocycliza-tion of parat-:fins with hydrocarbon fragments such as methane, ethane, and propane.

With the increased demand for high octane gasoline, e.g., C+ reformate having an octane rating l(Research +3 -ml. TEL) Vof 100 the purity of :reformer hydrogen has decreased. 4Present demands for hydrogen on the other hand have increased.

Thus, for example, it is a matter yof economics whether MidContinent naphtha be hydrogenated :to reduce the sulfur content to provide reformer feed for a reformer unit employing particle-form platinum-group metal reforming catalyst. On the other hand, when thermal naphtha is to be reformed in the presence of platinum-group :metal reforming catalyst there is no choice. The nitrogen content of thermal naphtha is so high that the platinum-group metal reforming catalyst is inactivated in an impractically short time when the thermal naphtha is not hydrodesulfurized and hydrodenitrogenized prior to reforming. Consequently, the increased demand for upgrading previously accepted gasolines has krequired reforming of `thermal naphtha and consequently increased demand for hydroge to hydrodecontaminate the thermal naplitha.

It is now rather general practice to hydrodecontaminate domestic `fuel oil and similar hydrocarbon fractions `boiling above gasoline to upgrade `these fractions thus making a further demand for hydrogen.

The problem of meeting the more stringent specifications for jet fuels is being solved at least in part by hydrogenation of the naphthenes of the jet fuel.

The treatment of lubricating voil fractions in the presence of hydrogen has also been added to the list of refinery operations using hydrogen produced in the reforming unit. ln fact, the demand for hydrogen for pretreating reformer feed, hydrodesulfuring domestic fuel oil, hydrogenating jet fuel fractions, hydrotreating lubricatingfractions, and hydrocracking crudes, deasphalted crudes, topped crudes, and gas oil fractions exceeds the supply with the consequent result that the reformer make-gas is passed from one refinery operation to another provided the contaminants picked up in one operation are not deleterious to the succeeding operation.

Hydrogen can be produced by partial oxidation of hydrocarbons and by other means but not at .a cost which makes its use attractive in the aforementioned and other retinery hydroprocesses. On the `the hand, there are refinery gas streams containing hydrogen but `of such low concentrations that the use thereof is not practical.

For example, the off-gas of catalytic cracking of gas The greater the severity 3,246,951 Patented Apr. 19, 1966 oil such 'as the olf-,gas from the Thermofor Catalytic Cracking of gas oil 4has the following composition:

In other words, were'ICC olf-gas to 'be used in hydro.- processing 'for every 1000 s.c.f. of hydrogen circulated 2855 s.-c.f. ,of diluent gas must also be compressed and transported. Many of the hydroprocesses can be carried out at lower pressure. That is .to say, .-it is the partial pressure of hydrogen not the -total pressure which is controlling in many hydroprocesses Hence, the higher the purity of vthe l'hydrogen-containing gas the lower the .total pressure for a given hydrogen partial pressure. The lower r the pressure the lower the capital cost; the higher the purity .of the hydrogen-containing gas stream .the lower the compression costs; the higher .the purity of the hydrogen-containing gas the lower the aging rate andthe longer the interval between regenerations, i.e., the longer the onstream .time and the fewer the off-stream Ytimes for regeneration .of the catalyst in a given calendar interval. For example, a decrease .of about 20% in investment cost of a new hydrodesulfurization unit treating middle kdistillate .can `be realized by employing recovered hydrogen of .90% purity over `the Vuse of .reformer make-gas of `60% purity. For a hydrodesul'furization unit treating 7,000 barrels .of middle distillate per day the saving in investment cost .can amount to about .$250,000. On an .existing hydrodesulfurization unit employing lthe recovered hydrogen of purity can double the on-stream time of the unit between regenerations.

Th-us, it is manifest that a means of raising the hydrogen `content of refinery gas streams will make available hydrogen which cannot be presently used and the availability of 90 Vpercentnby volume hydrogen will present additional advantages. The present invention provides the solution tothe twin problems.

In general, the ,present invention provides for contacting any refinery .stream containing at least 20 percent .by volume of hydrogen with a hydrogenatable extracting agent in the presence of a hydrogenation catalyst under hydrogenating conditions of temperature, pressure, and liquid hourly space velocity to react the hydrogen in the gas stream .with the hydr-ogenatable extracting agent to prolduce fat extracting agent and off-gas of reduced concenration of hydrogen. The fat extracting agent is separated from thedenuded gas stream. The fat extracting 4agent is then contacted with dehydrogenating catalyst under dehydrogenating conditions of temperature, pressure, and liquid hourly space velocity to produce a gaseous fraction containing at least 90 percent by volume of hydrogen and :lean extracting agent. The high purity gas stream flows lto .other hydroprocesses while the lean extracting 'agent is returned to the hydrogenation stage.

"The method of recovering hydrogen of at least 90 percent purity .briefly and generally described hereinbefore is to be distinguished over the method ofpurifying reformer `recycle Agas described in U.S. Patent No. 2,328,828 issued in 1943.

The patentee describes a process in which naphtha having `an ASTM knock rating of about 25 `to 45 mixed with about 0.5 .to 4 molsof hydrogen per mol of said naphtha is contacted with dehydrogenating catalyst such as chromia or molybdena on alumina at pressures of about 50 to 450 p.s.i., 900 to 1025 F. and at liquid hourly space velocities of about 0.1 to 10 to produce a product which is fractionated into a bottoms product boiling higher than gasoline, a heavy' naphtha, a benzene fraction boiling in the range of about 160 to 185 F. and uncondensed overhead vapors. The overhead of the fractionation is stabilized at 150 to 500 p.s.i. to produce a sidestream -comprising C3 and C4 hydrocarbons, a light naphtha whichis mixed with the aforesaid heavy naphtha, and a stabilizer overhead -containing 25 to 75 percent by volume of hydrogen. The stabilizer overhead in part can be discharged from the system although it is preferred to contact the stabilizer overhead in admixture with the aforesaid benzene fraction with a hydrogenating catalyst at hydrogenating conditions of temperature of about 300 to 800 F. and pressures of about 50 to 5000 p.s.i. The benzene fraction is hydrogenated and separated from the treated stabilizer overhead. The hydrogenated benzene fraction is then contacted in admixture with the feed naphtha or separately with dehydrogenating catalyst at 900 to 1025 F., at 50 to 450 p.s;i., and at liquid hourly space velocities of about 0.1 to 10. The patentee also statesthat for naphthas having end boiling points below 450 F. diphenyl may be employed and for naphthas having boiling temperatures below 375 to 400 F. naphthalene can be used.

Since in the patented process gasoline is produced in admixture with the produced hydrogen and dehydrogenated benzene or diphenyl or naphthalene the condition of dehydrogenation must -be regulated to be the optimumv for the production of gasoline of the required octane rating regardless of the conditions which are optimum for recovery of the hydrogen and for minimum degradation of the naphthalene, diphenyl, xylene, toluene, or benzene. Degradation of the aromatic hydrocarbons by dealkylation during dehydrogenation dilutes the hydrogen produced and undoubtedly at least in part accounts for the failure to exploit this process industrially.

In direct contrast to the foregoing method of returning to the reformer a hydrogen donor or a mixture of dehydrogenated hydrogen donor and hydrogen the present method provides for subjecting the fat extracting agent to dehydrogenating conditions which are optimum for dehydrogenation without substantial degradation of the extracting agent.

In contrast to the off-gas from the catalytic cracking of gas oil which contains about 35 percent by volume of hydrogen (Table I) the make-gas of reforming unit treating Mid-Continent naphtha having a boiling range of about 160 to about 360 F. to produ-ce C5+ reformate having an octane rating (Research-F5 ml. TEL) of 100 has a hydrogen concentration of about 73 percent by volume as is shown in Table II. On the other hand, the make-gas from a reforming unit treating a Mid-Eastern naphtha having a boiling range of about 160 to about 360 F. to produce C3+ reformate having an octane rating (Research-l-S ml. TEL) of 100 has a hydrogen concentration of about 61 percent by volume as is shown in Table III.

TABLE II [Make-Gas from Mid-Continent Naphtha at 100 O.N. (Research-i-S m1. TEL)] TABLE III Component: Mol percent Hydrogen 61.0 C1 22.0 C2 12.0 C3 4.0 I-C.L 0.5 N4C3 0.4 I-C5 0.1

Total 100.0

Hydrogen sulfide, p.p.m. 20

In many reneries the off-gas of the naphtha pretreater is contacted with the raw naphtha to remove C4 and heavier hydrocarbons from the pretreater -gas while removing water and oxygen from the raw naphtha. Hence, the concentrations of C4 and heavier concentrations are low enough to be disregarded. In most refineries the naphtha pretreater gas from which the hydrogen is to be recovered has the lfollowing 4composition (Table IV) TABLE IV Component: Mol percent Hydrogen 65.0 C1 20.0 C3 11.0 C3 3.7 C4+ about 0.2 Hydrogen sulfide 0.1

However, it is conventional because of local antipollution laws to remove the hydrogen sulfide from this gas. The hydrogen sulfide content is generally reduced below p.p.m.

The off-gas from a hydrodecontaminating unit for the treatment of domestic fuel oil, kerosine, jet fuel and the like has the average composition set forth in Table V.

TABLE V Component: Mol percent Hydrogen 50.0 C1 31.0 C3 14.0 C3 4.5 C4 0.5 Hydrogen sulfide, p.p.m. 10

In general, the aforedescribed off-gases will be treated to reduce the sulfur content to at least 10,000 p.p.m. by volume when using a sulf-active hydrogenating catalyst and to at least 400 p.p.m. by volume when using a platinum-group metal hydrogenating catalyst. Suitable extracting agents are pure single or multi-ring aromatic hydrocarbons, pure single or multi-ring hydrocarbons having one or more saturated, i.e., alkyl side-chains, e.g., benzene, toluene, xylenes, trimethyl benzenes, ethyl benzene, diethyl benzene, ethyl-methyl benzene, naphthalene, methyl naphthalene, anthracene, methyl anthracene, phenanthrene, methyl phenanthrene, mixtures of the foregoing, aromatic petroleum fractions such as synthetic tower bottoms obtained in the catalytic cracking of gas oil, mixtures of aromatic hydrocarbons recovered from the extract obtained by extraction aromatic hydrocarbon-containing naptha, kerosine synthetic tower bottoms, and, in general, any aromatic hydrocarbon mixture with a selective solvent such as sulfur dioxide, alkylene glycols, e.g., diethylene glycol, triethylene glycol, Chlorex and others. In general, the hydrogen extracting agent comprises aromatic hydrocarbon substantially devoid of hydrogenatable material not readily dehydrogenatable at extracting agent regeneration conditions without the production of unsaturated hydrocarbons. Suitable hydrogenation catalysts are hydrogenation catalysts having a hydrogenation component selected from the group consisting of the metals having atomic numbers 44 to 96 inclusive and 76 to 78 inclusive on refractory oxide base comprising alumina, silica-alumina, silica-zirconia. The

'hydrogen-ation catalyst can be selected on the basis-of the "catalyst poisons present in the gas from which hydrogen 'is `to be recovered. Thus, when the hydrogen-containing gas contains hydrogen sulfide it is preferred to use a 'catalyst which 'is not seriousiy Vinactivated by sulfur, eg., -having la hydrogenation component such as nickeltungsten sulfide, tungsten disulfide, nickel-molybdenum sulfide. -On the vother hand, when the sulfur content of the #hydrogen-containing ,gas to be treated does not exceed 30 p.p.m. A(parts per million) by weight of the acceptor platinum-group metal hydrogenating vcatalysts such as a `catalyst comprisingabou-t 0.30 to about 1.00 percent by weight of platinum, palladium, osmium, iridium on alu- Vmina `base can be used. However, when desirable or .necessary or convenient, the gas containing the hydrogen :sulfide can be contacted with an aqueous solution of -al-kanolamine such -as diethanolamine or other treating agent for removal of the hydrogen sulfide to provide a hydrogen-containing gas containing not more than ,p.-p.=m. of sulfur by weight based on the treating agent and hydrogenating catalyst having as the hydrogenating component platinum-group meta-l used to recover hydrogen from `the so-treated hydrogen-containing gas. n

Hydrogenating conditions are as set forth in Table VI.

TABLE v1 Gperating Conditions Broad Preferred Temperature, It.,A 200 tomo 30o to 650. Pressure, p.s.i.g 150 to 2.500-... 500 to 1,500. 'Liquid Hourly Space Velocity (v.'/hr./ 0.25 t0 10 0.5 to 4.

vw l

1 v. .,-volurne ol treating agent in barrels: v.-volume of hydrogenating `catalyst (barrel).

able dehydrogenating vcatalysts are catalysts comprising platinum-group metal on refractory oxide support. Thus, for example, a dehydrogenating catalyst comprising about -0.35 to about 0.6 percent `of platinum or alumina support is suitable. However, the preferred dehydrogenating catalyst comprises about 0.35 to about 0.6 percent yby weight of platinum and not more than about 0.7 percent by weight of chlorine on alumina. Y

Dehydrogenating conditions are set forth in Table VII.

4TABLE VII Operating Conditions ABroad Preferred Temperature, F v i 800 to 1,050 850 to 950. Pressure, psig 100 to 1.000.... 200 to 500. Liquid Hourly Space Velocity v.fea/ 0.5 to 100 2 to 30.

hr. v. Hydrogsen Circulation (mol of purified 1 to 20.V 3 to 10.

hydrogen per mol of FEA).

fea, FEA-Fat Extracting Agent.

For the recovery of hydrogen from dilute hydrogencontaining gas such as 'the off-gas from catalytic cracking of gas oil, naphtha pretreater off-gas, oit-gas from middle distillate hydrodecontaminat-ion and similar refinery olfgases containing at least percent by volume of hydro- :gen the dilute `refinery gas is contacted with the extracting agent -i-n a lean extracting agent-to-hydrogen mol ratio in the range of 0.20 to 5.

Illustrative of the flow of liquids and gases in the present method of recovery hydrogen of `at least 90 percent purity from refinery gas containing at least 20 percent of hydrogen `are the flow sheets of FIGURES 1 and 2.

The flow sheet FIGURE 1 illustrates the use of two hydrogen extractors or reactors and one fat extracting agent regenerator or reactor. The hydrogen extractors or reactors are used alternately. That is tonsay, while one is on-stream the other is off-stream and the catalyst is regenerated usually by combustion of the coke deposited on the catalyst during the on-st-ream period of the cycle in any inert gas containing lfree oxygen. The deposition of coke in the fat extract-ing agent regenerator or reactor is minimal since .temperatureand liquid hourly space velocity are `correlated to minimize hydro-cracking and dealkylation of alkyl-substituted extracting agent. It is -to be noted that regeneration of the fat extracting agent FEA is an endothermic reaction which under adiabatic reaction conditions results in a temperature drop in the reactor. Since Adealkyl'ation '-of alkyl-substituted acceptors occurs at temperatures below that at which dehydrogenation of the 'hydrogenated extracting agent occurs it is desirable, and in fact preferred, to separate the regenerated or lean extracting agent from the dehydrogenating catalyst substantially as soon as the temperature of the reactan-t reaches that at which the dehydrogenation reaction ceases. This is mos-t readily accomplished by contacting the fat extracting agent with the dehydrogenating catalyst at a catalyst-to-oil ratio in the range of about 0.2 to 0.5 tons of catalyst per 1000 barrels of fat -extracting agent per day, i`.e., at a liquid hourly space velocity (v.fe /hr./v.c) in the range of about l0 to about 25 when using a dehydrogenating catalyst comprising about 0.35 to about 0.6 percent by weight of platinum and `not more than 0.7 percent of chlorine and/ or fluorine on alumina support. y

As illustrated in FIGURE 1 at start-up a lean extracting agent such as the hydrocarbons recovered from the extraction of kerosine with sulfur dioxide or the synthetic tower bottoms recovered from the product of the catalytic cracking of gas oil, or in general a hydrogenatabile aromatic hydrocarbon comprising at least 60 volume percent including volume percent, of aromatic hydrocarbon and substantially devoid of other hydrogenatable material is drawn from a source not shown through pipe 1 under contnol of val-ve 2. After start-up only make-up quantities of fresh extracting agent are drawn through valve 2. The extracting agent (lean) ilows through pipe 1 to the suction side of pump 3. Pump 3 discharges the `lean extracting agent (LEA) into pipe 4 at a pressure in excess of that in the on-stream extractor or

reactor

15 or 16 by the amount of pressure drop between pump 3 and the 'on-stream extractor or react-or. The lean extracting agent ows through pipe 4 to indirect heat exchanger 5 where the lean extracting agent is in heat transfer relation with lthe eiiluent of the on-s'tream extractor or reactor flowing from eiuent manifold 26. From heat exchanger 5 the lean extracting agent ows through pipe 6 to coil 7 in heater or furnace 8. In furnace 8 Vthe lean extracting agent is heated to a temperature to `at least 2.0 percent by volume of hydrogen contains not more than about 10,000 p.p.m. by volume of sulfur as hydrogen suliide.)

Contemporaneously with the flow of lean extracting agen-trthrough reactor 15 refinery off-gas containing at least 20 percent by volume of hydrogen and not more `than 10,000 p.p.m. by volume of sulfur as hydrogen suliide ows from a source not yshown through conduit lene.

' much as live.

17 to compressor 18. Compressor 18 discharges the refinery off-gas into conduit 19 at a pres-sure in excess of that in reactor 15 by the pressure drop b-etween cornpressor 18 and reactor 15. The compressed refinery offgas flows through conduit 19 to coil 20 in furnace or heater 21. In heater 21 the refinery off-gas is heated to a temperature such that when mixed with the heated lean extracting agent in inlet manifold 10 the mixture will have a hydrogenating temperature. The heated lean extracting agent and heated refinery off-gas flow respectively through pipe 9 and conduit 22 to inlet manifold 10.

The lean extracting agent-to-hydrogen mol ratio depends on the extracting agent employed. For example, if naphthalene is used, one mol can bind as many as five mols of hydrogen when completely hydrogenated in accordance with the following equation:

CioHs-i-'ZHz'e-CioHiz-t 3H2C1nHis Naphthalene Tetralin Decalin Thus, the lean extracting agent-to-hydrogen mol ratio is 0.20 (l/) for complete hydrogenation of the naphtha- For partial hydrogenation (to tetralin), the lean extracting agent-to-hydrogen mol ratio is 0.50 (l/ 2). In the case of benzene, one mol can react with a total of three mols of hydrogen when completely hydrogenated in accordance with the following equation:

The lean extracting agent-to-hydrogen mol ratio for benzene is 0.33 (1/3). H-owever, to ensure maximum recovery of hydrogen from the dilute (with respect to hydrogen) refinery olf-gas an excess of hydrogenatable aromatic hydrocarbon above that which theoretically can be hydrogenated by the hydrogen of the dilute refinery oE-gas is generally used. Accordingly, it is preferred to employ lean extracting agent-to-hydrogen mol ratios as This excess of lean extracting agent not only ensures maximum extraction of the hydrogen from the dilute refinery off-gas but also more ready control of the temperature of the endothermic reaction.

Returning vto a description of FIGURE 1; the heated lean extracting agent and dilute refinery off-gas are mixed in inlet manifold 10 in the proportion of at least 0.20 mol of hydrogenatable aromatic hydrocarbon per mol of hydrogen in the dilute refinery off-gas to provide an extraction mixture. The extraction mixture at hydrogenation temperature flows through inlet manifold 10 to manifold branch 11 (valve 14 closed; valve 12 open). The extraction mixture flows downwardly in recator 15 in contact with hydrogenation catalyst. The fat extracting agent, i.e., at least partially hydrogenated aromatic hydrocarbon, and denuded refinery off-gas, designated extractor effluent, flow from reactor 15 through manifold branch 23 (valve 24 open; valve 28 closed) to outlet manifold 26. The extractor effluent flows through outlet manifold 26 to indirect heat exchanger 5 to conduit 30 and cooler 31. In cooler 31 the extractor effluent is cooled to a temperature at which the fat extracting agen-t is liquid at the existing pressure. The cooled extractor efuent flows from cooledl 31 through conduit 32 to liquid-gas separator 33. In liquid-gas separator 33 the uncondensed denuded refinery off-gas separates from the condensed fat extracting agent. The uncondensed denuded refinery off-gas comprising C1 to C5 hydrocarbons flows from separator 33 through conduit 34 to hydrocarbon recovery, processing, the refinery fuel main, or the refinery flare.

The condensed fat extracting agent flows from separator 33 through pipe 35 to the suction side of pump 36. The fat extracting .agent is discharged into pipe 37 by pump 36 at a pressure greater than the pressure in reactor 25 by the pressure drop between pump 36 and reactor 25.

The fat extracting agent flows through pipe 37 to indirect heat exchanger 38. In indirect heat exchanger 38 the fat extracting agent is in heat transfer relation with the effluent of reactor 25 comprising hydrogen and lean extracting agent flowing from reactor 25 through conduit 43. From indirect heat exchanger 38 the fat extracting agent flows through conduit 39 to coil 40 in furnace or heater 41. At a point in conduit 39, intermediate to indirect heat exchanger 38 and to furnace 41, percent or better hydrogen flowing from compressor 51 through conduit 52 is mixed with the fat extracting agent in the mol ratio of aboutl 1 to about 20, preferably about 3 to 10 mols of hydrogen per mol of hydrogenated aromatic hydrocarbon -to provide a recovery charge.

The recovery charge is heated in furnace 41 to a dehydrogenating temperature to provide a vapor inlet temperature at reactor 25 in the range of aboutSOO" to about 1050" F. The heated recovery charge flows from heater 41 through conduit 42 to reactor 25. The recovery charge flows downwardly through reactor 25. The extracting agent regenerator eluent comprising hydrogen and lean extracting agent flows from reactor 25 through conduit 43, indirect heat exchange 38 and conduit 44 to cooler 45. In cooler 45 the extracting agent regenerator effluent is cooled to a temperature at which the extracting agent is condensed at the existing pressure. The uncondensed hydrogen and minor concentration of hydrocarbons together with the condensed lean extracting agent flow from cooler 45 through conduit 46 to liquid-gas separator 47.

In liquid-gas separator 47 the uncondensed recovered hydrogen and hydrocarbons boiling below the boiling point of the extracting agent separate from the condensed lean extracting agent and ow therefrom, as la gaseous stream containing at least 90 percent by volume of hydrogen balance, to make percent by volume, hydrocarbons boiling below the boiling point of the extracting agent, through conduit 48 to hydroprocesses. A portion of the at least 90 percent pure hydrogen is divertcd from conduit 48 under control of valve 50 to conduit 49 and the suction side of compressor 51. The quantity of 90 percent pure hydrogen so diverted is sufficient to provide the aforesaid 1 to 20, preferably .3 to 10 mols of hydrogen per mol of hydrogenated aromatic hydrocarbon in the recovery charge. Compressor 51 recompresses the 90 percent pure hydrogen to a pressure in excess of that in conduit 39. The recompressed 90 percent pure hydrogen flows from compressor 51 through conduit 52 to conduit 39.

The condensed lean extracting agent (LEA) separated in liquid-gas separator 47 from the uncon-densed at least 90 percent pure hydrogen, flows from separator 47 through pipe 53 and 1 to the suction side of pump 3 (valve 54 open; valve 2 closed except to admit make-up extracting agent to pipe 1). The lean extracting agent is then recycled to the hydrogenating-reactor or extractor 1S or 16 which is on-stream to extract hydrogen from further quantities of refinery off-gas containing at least 20 percent by volume of hydrogen. Those skilled in the art will recognize that for simplicity the piping required for regeneration of coked catalyst has been omitted since regentration of catalyst deactivated with coke is too well known to require description. Briefly, the catalyst in the off-stream extractor, presently assumed for descriptive purposes to be reactor 16, is regenerated by isolating the reactor or extractor from the extracting agent and refinery off-gas and from extractor effluent manifold. The extractor or reactor is pluggedV with inert gas such as flue gas, steam, nitrogen. The purged reactor is then pressured with flue gas and the flue gas heated and encirculated until the temperature in the catalyst bed is at a burning temperature. Free oxygen, e.g., air is then admixed which the circulating inert gas and the coke burned from the catalyst. When the burning front reaches the bottom of the catalyst bed and free oxygen is detected in the effluent gas the supply of air is cut-off and the reactor purged until the `oxygen content of the etliuent gases does not exceed 1 mol-percent. The reactor or extractor is then pressured with refinery off-gas to be extracted and the ilow of heated lead extracting agent and renery off-gas to be extracted diverted from the on-steam extractor -or reactor to the olf-steam extractor in which the catalyst has been regenerated'.

FIGURE 2 is a flow sheet likewise illustrating the use of three: reactors for extracting hydrogen from dilute reinery gas streams and recovering they extracted hydrogen. However, whereasv inthe embodiment illustrated in FIG- URE l one reactor is used only for the recovery of extracted hydrogen while the. other two reactors are used alternatively for extracting. hydrogen in the embodiment illustrated in FIGURE 2 all of the reactors are charged with a catalyst'y having not only hydrogenating capabilities but also having dehydrogenating capabilities.

Catalysts having hydrogenating and dehydrogenating capabilities suitable for use in the present method are for example, a mixture of oxides of cobalt and molybdenum on alumina support; platinum-group metal on refractory oxide support, such as` alumina, silica-alumina, and the like; and catalysts comprising a mixture of at least 18 percent by weight of at least one. oxide of chrominum, molybdenum, and vanadium, and at least one refractory oxide such as alumina, silica, zirconia.

i The embodiment illustrated in FIGURE 2 has parvticular advantages when recovering hydrogen from dilute refinery gas containing more than 30 p.p.in. of sulfur 'based upon the weight of the hydrogenatable aromatic hydrocarbons used as, for example, when recovering hydrogen from the hydrogen-containing off-gases from a hydrodesulfurizing unit treating middle distillate fuels such as kerosine, domestic fuel oil, jet fuels and the like. For example, the hydrogen-containing gas from a hydrodesulfurization unit hydfrotreating a middle distillate e.g., domesticV fuel oil has the following average cornposition:

t P.p.ni.-parts per million by volume on gas. While a dilute off-gas containing such a concentration of sulfur and ammonia must be pretreated to reduce the sulfur concentration to not more than 30 p.p.m. by weight based on the extracting agent (15 to 400 ppm. by volume :based on the gas for treating agent-to-hydrogen mol ratios in the range of 0.20 to 5) and the ammonia concentration to not more than 5 p.p.m. by weight based on the said extracting agent (5 Ito 150 p.p.in. by volume based on the gas for hydrogenatable aromatic-to-hydrogen mol ratios in the range of 0.20 to 5 when platinum-group metal catalyst having hydrogenating and dehydrogenating capabilities is employed for extracting hydrogen (from dilute off-gas) and recovering hydrogen (from lean extracting agent), in contract when using sulfur and nitro- -gen-insensitive catalyst having hydrogenating and dehydrogenating capabilities such as a mixture of oxides of chromium and aluminum, or of oxides of chromium, molybdenum, and aluminum, or of oxides of cobalt,

lmolybdenum, and aluminum, the concentrations of sulfur and ammonia in the dilute off-gas can be as high as 10,000 p.p.m. by volume and 1500 p.p.m. by volume respectively, i.e., 750 p.p.ni. and 50 p.p.in. by weight based on the lean extracting agent using a mol ratio in the range of 0.20 to 5 mols of said aromatic hydrocarbon per mol of hydrogen. Dilute off-gas such as the aforesaid dilute off-gas from a hydrodesulfurization unit trea-ting middle 10 distillate fuel oil as characterized in Table VIII is-illustrative of dilute off-gas treated in accordance with the flow-sheet of FIGURE 2.

Reactors

114, 160, and 141 are charged with sulfur and nitrogen-insensitive catalyst having both hydrogenating and dehydrogenating capabilities. Lean extracting agent, for example, synthetic crude. tower bottoms from a catalytic cracking unit having the average composition set forth in Tabley IX is drawn from a source not shown through pipe 101 under control of valve. 102 by pump 103.

TABLE` IX [Lean Extracting Agent-#TCC Synthetic Crude Tower- Bottoms] Boiling Range, F. 300 to 371 IBP, F. 300 5%, F. 314 10%, F. 3-30 F. 344 EBP, F. 37'1 Aromatics (principally polyalkylbenzenes) 60-61`.8 vol. percent.

The lean extracting agent is discharged by pump 103 into pipe 104 at a pressure greater than that in the reactor of

reactors

114, 160, and 141 on-stream by at least the pressure drop between pump 103 and the on-stream reactor. The lean extracting agent flows Ithrough pipe 104 to coil 105 in furnace or heater 106. In heater 106 the lean extracting agent is heated to a temperature such that when mixed with dilute oif-gas in the ratio of about 0.20 to about 5 mols of hydrogenatable aromatic per mol of hydrogen in the dilute oif-gas the extraction mixture soformed has a hydrogenating temperature at the vapor inlet of the on-stream extractor or hydrogenating reactor. (For simplicity it will be assumed that the following cycle having the indicated phases is employed.)

TABLE X Ri (114) R2 (160) Ra (141) First Phase Extracting Idle Regeneration of Hydrogen. Extracting Agent,

Second Phase Regeneration of Regeneration of Extraction of Catalyst. Extracting Hydrogen to Agent. about 10% coke on catalyst.

Third Phase Regeneration of 'Extraction to Regeneration of extracting about 10% catalyst. agent. coke on catalyst.

Fourth Phase Extraction to Regeneration of Regeneration of about 10% catalyst. extracting coke on cataagent.

` lyst.

Fifth Phase Regeneration of Regeneration of Extraction to catalyst. extracting about 10% coke agent. on catalyst.

Sixth Phase Regeneration of Extraction to Regeneration of extracting about 10%. catalyst. agent. coke on catalyst.

The heated lean extracting agent Hows from furnace 106 through conduit 107 to inlet manifold 108 and manifold branch 112 (valve 113 open;

valves

159 and 140 closed). At a point in conduit 107 intermediate to inlet manifold 10S and heater 106 the dilute oif-gas from which hydrogen is to be recovered is mixed with the lean extracting agent in a mol ratio in the range of 0.20 to 5 mols of hydrogenatable aromatic per mol of hydrogen. The dilute off-gas iiows from a source not shown through conduit 110 to the suction side of compressor 109. Compressor 109 compresses the off-gas to a pressure at least equal to that in conduit 107 and discharges the off-gas into conduit 111 through which the off-gas ows to conduit 107 and admixture with the lean extracting agent. (Those skilled in the art will appreciate that when the off-gas is available contiguous to conduit 107 at a pressure at least equal to that in conduit 107 recompression is unnecessary and compressor 109 can be bypassed.)

y, The extractor mixture comprising off-gas and lean extracting agent in proportions to provide 0.20 to mols `of hydrogenatable aromatic per mol of hydrogen in the olf-gas flows through conduit 107 to extractor inlet manifold 108. From extractor inlet manifold 108 in the first yphase of the cycle the extractor mixture ows through manifold branch 112 (valve 113 open;

valves

159 and 140 closed;

valves

166, and 157 closed) to extractor reactor 114. The extractor mixture flows downwardly in extractor 114 in contact with particle-form solid hydrogenating catalyst 'in this description sulfurand nitrogeninsensitive hydrogenating catalyst. T he extractor effluent comprising fat extracting agent and extracted off-gas flows from extractor 114 to manifold branch 11.5. (Valve 116 open;

valves

168, 162, 164, and 143 closed.) The extractor etiluent flows through manifold branch 115 to extractor elfluent manifold 117. The extractor effluent flows through extractor efuent manifold 117 to indirect heat exchanger 118 Where the extractor effluent is in heat transfer relation with lean extracting agent flowing from indirect heat exchanger 120 through conduit 154. From indirect heat exchanger 118 the extractor efliuent ows through conduit 119 to indirect heat exchanger 120 where the extractor efuent is in heat transfer relation with lean extracting agent flowing from liquid-gas separator 151 through pipe 153. From indirect heat exchanger 120 the extractor eiuent flows through conduit 121 to cooler 122. In cooler 122 the extractor euent is cooled to a temperature at which the fat extracting agent is condensed at the existing pressure. (Fat extractving agent as used hereinbefore, here, and hereinafter designates extracting agent in which the hydrogenatable arolmatic has been in contact with hydrogen-containing gas in the presence of hydrogenating catalyst under hydrogenating conditions of temperature, pressure, and liquid hourly space velocity regardless of the degree of completeness ofthe hydrogenation of said hydrogenatable aromatic hydrocarbon. Thus, for extracting agents containing hydrogenatable aromatic hydrocarbons such as benzene and alkylbenzenes it is immaterial whether all of the molecules each have accepted six atoms of hydrogen or only one atom of hydrogen or that a portion of the molecules of the aromatic each have accepted one atom of hydrogen and the balance none of a portion of the molecules of the aromatic each have accepted one to six atoms of hydrogen and the balance none.) The cooled extractor etiluent comprising extracted off-gas and fat extracting agent flows from cooler 122 through conduit 123 to gas-liquid separator 124. In gas-liquid separator 124 the extracted off-gas separates from condensed fat extracting agent. The extracted off-gas, designated waste gas, ows from separator 124 through conduit 125 to means for recovery of hydrocarbons, isomerization of hydrocarbons, hydrocarbon conversion as of methane to acetylene, the refinery Vfuel main, the refinery fiare or other means of treating a hydrocarbon stream of this composition.

The condensed fat extracting agent ows from separator 124 through pipe 126 to indirect heat exchanger 127 Where the fat extracting agent is in heat transfer relation with extracting agent regenerator effluent flowing from regenerator outlet manifold 146, indirect heat exchanger 129, conduitl 147 to indirect heat exchanger 127. From indirect heat exchanger 127 the fat extracting agent flows through pipe 128 to indirect heat exchanger 129 where the fat extracting agent is in heat transfer relation with the regenerator efuent flowing from fat extracting agent regerenator manifold 146. From indirect heat exchanger 129 the fat extracting agent flows through pipe 130 to ythe suction side of pump 131. Pump 131 discharges the fat extracting agent into pipe 132 through which the fat extracting agent flows to coil 133 in furnace or heater 134:

In furnace 134 the fat extracting agent is heated to a temperature such that, when mixed with hydrogen in regenerator inlet manifold 136 to provide a regeneration feed comprising about l to about 20 mols of hydrogen of at least percent purity per mol of hydrogenated aromatic hydrocarbon, a vapor inlet dehydrogenation temperature of about 800 to about 1050 F. at thel regenerator-reactor inlet (141) is provided. From heater 134 the heated fat extracting agent (FEA) flows through pipe 135 to regenerator inlet manifold 136. In regenerator inlet mani-fold 136 the fat extracting agent is mixed with hydrogen of 90 percent purity flowing either from a source not shown through

conduits

171 and 169 under control of valve 172 or after start-up from liquid-gas separator 151, and

conduits

152 and 169 under control of valve 170. (When recovered hydrogen or extraneous hydrogenis not available at a pressure at least equal to that in regenerator manifold 136 the extraneous hydrogen and the recovered hydrogen is compressed to the neces.- sary pressure for introduction into regenerator manifold 136.)

From regenerator inlet manifold 136 the regenerator charge mixture comprising fat extracting agent and hydrogen in the ratio set forth hereinbefore flows through manifold branch 137 (valve 138 open;

valves

157 and 166 closed; valve 140 closed) and manifold branch 139 to regenerator 141. The regenerator charge mixture flows downwardly in regenerator 141 in Contact with dehydrogenating catalyst described hereinbefore. The regenerator efuent comprising recycle hydrogen, i.e., hydrogen admixed in regenerator manifold 136, make-hydrogen, i.e., hydrogen removed lfromthe fat extracting agent and lean extracting agent flows from regenerator 141 through manifold branch 144 (Valve 145 open;

valves

143, 164, and 168 closed) to regenerator manifold 146. From regenerator manifold 146 the regenerator effluent llows to indirect heat exchanger 129, conduit 147, indirect heat exchanger 127 and conduit 148 to cooler 149.

In cooler 149 the regenerator effluent comprising hydrogen and lean extracting agent is cooled to a temperature at which the lean extracting agent is liquid. The condensed lean extracting agent and uncondensed hydrogen ow from cooler 149 through conduit 150 to liquid-gas separator 151.

In liquid-gas separator 151 hydrogen of at least 90 percent purity separates from lean extracting agent. The hydrogen of at least 90 percent purity flows from separator 151 through conduit 152 to use in other hydro processes. A portion to -provide about 1 to about 20 mols of hydrogen per mol of hydrogenatable aromatic hydrocarbon in regenerator manifold 136 as described hereinbefore is diverted through conduit 169 under control of valve 170 (valve 172 closed).

The condensed lean extracting agent flows from separator 151 through pipe 153 to indirect heat exchanger 120, pipe 154, indirect heat exchanger 118 and pipe 155 to pipe 101 for use as extracting agent for the extraction of hydrogen from further amounts of 4dilute renery off-gas.

As set forth in Table X supra in the first phase of a cycle reactor 114 is in the extraction operation, reactor 141 is in `the regeneration of the extracting agent, and reactor is idle. When the activi-ty of `the catalyst in extractor 114 is reduced to an industrially appreciable extent as indicated by a decrease -in the difference in temperature between the temperature of the extractor effluent and the vapor inlet of the extractor, the extractor is taken .olf-stream and the catalyst regenerated. Regenerator 141 is then preferably taken off regenerating of the extracting agent and used for extracting hydrogen from the dilute refinery off-gas. Reactor 160 which in the rst phase of the cycle is idle is then placed in the ifat extracting agent regenerating operation. Operation in the second phase of the cycle then proceeds until the catalyst in reactor 141 has been contaminated by the deposition of an amount of coke equal to about 10 percent by weight 13 of the catalyst therein as determined, for example, bythe decrease in the difference between the temperature of the extractor effluent and the vapor inlet temperature. (Hydrogenation being an exotherrnic reaction the temperature .of the extractor effluent is higher than the vapor inlet temperature. The more vigorous the hydrogenation, Le., the more acti-ve the hydrogenat-ion, the `greater the temperature rise.) Reactor 141 is then taken off-stream and the catalyst therein regenerated while reactor 1160 (extracting agent regeneration in the second phase) is put on-stream as an extractor .until about percent by weight of coke is deposited thereon and reactor 114 (catalyst now regenerated) is put .on-stream for regeneration of the extracting agent in the third phase of the cycle. fIn the fourth phase of the cycle the .catalyst in reactor 160 (extraction stage in lthe third phaser) is regenerated, reactor .1-41 (regeneration of .catalyst -in the third phase) is yused for regeneration -of the extracting agent, and reactor 114 previously used in the regeneration of extracting agent) is used Afor extracting hydrogen from dilute refinery ott-gas until the amount of coke deposited on the catalyst amounts to about 10 percent -by weight thereof.

In the .fifth phase of .the cycle the catalyst -in reactor 114 is regenerated, reactor 1411 is used as .a hydrogen extractor, and reactor 160 is used to regenerate the fat extracting agent. In the sixth phase of the iirst cycle reactor 11'4 is .used for the regeneration of the :fat extracting agent, reactor 160 is used as an extractor of hydrogen from dilute refinery off-gas and the catalyst in reactor 14.1 is regenerated. This ends the first cycle.

Those skilled in the art Iwill rec-ognize that there is nothing immutable in .the sequence in which the reactors appear in any cycle. Tha-t is to say, in the iirst phase of the `first cycle reactor 141 can be used .as an extractor, reactor 160 used for regeneration of the extracting agent and reactor L14 can be idle. The other possible combination can .also be used in the first phase. However, it is preferred after the .first phase to use the reactor in 'which the catalyst is fresh or regenerated for regeneration of the extracting agent and to use the catalyst which in the previous phase was used `for regeneration of the extracting .agent for the extraction of hydrogen from refinery gas streams dilute with respect to hydrogen.

I claim:

I1. A method of recovering hydrogen from refinery gas dilute with respect to hydrogen which comprises in a gas extractor contacting renery gas comprising at least about Z0 percent by volume of hydrogen and an extracting agent having as its sole hydrogenatable material at least one aromatic hydrocarbon with hydrogenating catalyst at hydrogenating conditions of temperature, pressure, and liquid hourly space velocity, obtaining extracltor effluent comprising fat extracting agent and extracted refinery gas, cooling said extractor eiuent 4to a -temperature at which said fat extracting agent is condensed at the existing pressure, separating said cooled extractor effluent into waste gas comprising refinery gas having reduced hydrogen content and fat extracting agent, in an extracting agent regeuerator contacting only said separated fat extracting .agent in admixture .with hydrogen of at least 90 percent purity with dehydrogenating catalyst at dehydrogenating conditions of temperature, pressure, and liquid hourly space velocity whilst maintaining a concentration of at least `90 percent .of hydrogen in the fraction of reaction vapors boiling below the boiling point of said lean extracting agent, obtaining regenerator effluent comprising hydrogen produced in said extracting agent regenerator, the .aforesaid admixed hydrogen, and lean extracting agent, cooling said regenerator effluent to a temperature at which said lean extracting agent is condensed at the existing pressure, separating said cooled regenerator euent into a gaseous fraction comprising at least 90 percent hydrogen by volume and condensed lean extracting agent, recycling a portion of said gaseous fraction to said .extracting agent regenerator andrecovering the 'balance of said gaseous fraction as `hydrogen of at least 9.0 percent purity, and recycling 'at least a portion of said condensed -lean extracting agent to the aforesaid Vgas extractor.

2. The method of claim d wherein the hydrogenating catalyst and the dehydrogenating catalyst yare of .substan- .tially the same composition.

3. The method set forth. i-n claim I1 lwherein the hydrogenating catalyst and the .dehydrogenating catalyst are of substantially different composition.

4. The method set forth in claim -1 wherein the hydro- .genatingcomponent of the hydrogenating catalyst is substantially the lsame as the dehydrogenating component of .the dehydrogenating catalyst.

5. The method. set forth `in .claim 1 wherein the hydrogenating component ofthe hydrogenating catalyst is different .to the dehydrogenating component of the dehydrogenating catalyst.

f6. The method set forth in claim 1 lwherein three reactors are charged with substantially the same quantity of particle-form solid catalytic material having hydrogenating and dehydrogenating capabilities, wherein a cyclic manner catalyst having a coke deposit of about 10 percent by weight based on the catalyst is regenerated, then used -for extracting hydrogen from refinery gas dilu-te with respect to hydrogen.

7. The method set forth in claim 6 wherein the particle-form solid catalytic material is sulfur and nitrogen insensitive.

y8. The method set forth .in claim 6 wherein the particle-form solid catalytic material is platinum-group metal and the refinery gas to be extracted contains not more .than 400 p.p.m. volume of hydrogen sulfide.

9. A method of recovering hydrogen from refinery gas dilute with respect Ato hydrogen which comprises in a gas extractor contacting renery gas comprising at least about 20 percent by volume of hydrogen and an extracting agent having as its sole hydrogenatable material at least one aromatic hydrocarbon with hydrogenating catalyst at hydrogenating conditions of temperature, pressure, and liquid hourly space velocity, obtaining extractor effluent comprising fat extracting agent and extracted renery gas, cooling said extractor effluent to a temperature at which said fat extracting agent is condensed at the existing pressure, separating said cooled extractor effluent into waste gas comprising .refinery gas having reduced -hydrogen content and fat extracting agent, in an extracting agent regenerator contacting only said separated fat extracting agent admixed lwith hydrogen of at least A percent purity with platinum-group metal dehydrogenating catalyst at dehydrogenating conditions of temperature, -pressure and liquid hourly space velocity 'whilst maintaining a concentration of at least 90 percent of hydrogen in the .fraction of reaction vapors boiling below the boiling point of said lean extracting agent, separating regenerator vapor from dehydrogenating catalyst after said fat extracting agent has contacted not more than 0.2 to 0.5 ton of said platinum-group metal dehydrogenating catalyst per 1000 barrels of extracting agent regenerated per day, said separated regenerator vapors comprising said admixed hydrogen, hydrogen produced in said regenerator, and lean extracting agent, cooling said separated regenerator vapors to a temperature at which said lean extracting agent is condensed at the existing pressure, separating said cooled regenerator vapors into a gaseous fraction containing at least 90 percent by volume of hydrogen and liquid lean extracting agent, recycling a portion of said gaseous fraction to the regenerator to supply the admixed hydrogen of at .least 90 percent purity, recovering the balance of said gaseous fraction as hydrogen of at least 90 percent puri-ty, and recycling at least a portion of said lean extracting agent to said extractor.

10; The method set forth in claim 9 wherein the hy- -drogenating catalyst is selected from the group consisting of mixture v,of oxides of cobalt and molybdenum,

knickel-tungsten, sul-tide, tungsten disulde, and. nickelmolybdenum sulde..

11. The method set forth in claim 9 wherein the dehydrogenating catalyst comprises about 0.35 to about 0.6 percent by weigh-t of platinum on alumina support.

i finery gas is treated to reduce the sulfur and nitrogen con-- tent thereof to not more than 400 p.p.m. by volume and 150 p.p.m. by volume respectively, wherein the hydrogenating catalyst comprises about 0.35 to about 0.6 percent by weight of platinum on alumina and wherein the l platinum group metal dehydrogenatng catalyst comprises about 0.35 to about 0.6 percent by Weight of platinum on alumina support.

14. The method set forth in claim 9 wherein the renery gas contains not more than 10,000 p.p.m. by volu-me of sulfur and not more than 1,500 p.p.m. by volume of nitrogen, wherein the hydrogenating-catalyst is nickeltungstein sulfide, and wherein the dehydrogenating catalyst comprises about 0.35 to about 0.6 percent by weight 0f platinum on alumina support.

References Cit-ed by the Examiner UNITED STATES PATENTS 2,328,828 9/1943 Marschner 208-516 2,502,958 4195O Johnson 208-56 X 2,913,401 11/1959 Weikart et al 23-212 X MAURICE A. BRINDISI, Primary Examiner.

Claims

1. A METHOD OF RECOVERING HYDROGEN FROM REFINERY GAS DILUTE WITH RESPECT TO HYDROGEN WHICH COMPRISES IN A GAS EXTRACTOR CONTACTING REFINERY GAS COMPRISING AT LEAST ABOUT 20 PERCENT BY VOLUME OF HYDROGEN AND AN EXTRACTING AGENT HAVING AS ITS SOLE HYDROGENATABLE MATERIAL AT LEAST ONE AROMATIC HYDROCARBON WITH HYDROGENATING CATALYST AT HYDROGENATING CONDITIONS OF TEMPERATURE, PRESSURE, AND LIQUID HOURLY SPACE VELOCITY, OBTAINING EXTRACTOR EFFLUENT COMPRISING FAT EXTRACTING AGENT AND EXTRACTED REFINERY GAS, COOLING SAID EXTRACTOR EFFLUENT TO A TEMPERATURE AT WHICH SAID FAT EXTRACTING AGENT IS CONDENSED AT THE EXISTING PRESSURE, SEPARATING SAID COOLED EXTRACTOR EFFLUENT INTO WASTE GAS COMPRISING REFINERY GAS HAVING REDUCED HYDROGEN CONTENT AND FAT EXTRACTING AGENT, IN A EXTRACTING AGENT REGENERATOR CONTACTING ONLY SAID SEPARATED FAT EXTRACTING AGENT IN ADMIXTURE WITH HYDROGEN OF AT LEAST 90 PERCENT PURITY WITH DEHYDROGENATING CATALYST AT DEHYDROGENATING CONDITIONS OF TEMPERATURE, PRESURE, AND LIQUID HOURLY SPACE VELOCITY WHILST MAINTAINING A CONCENTRATION OF AT LEAST 90 PERCENT OF HYDROGEN IN THE FRACTION OF REACTION VAPORS VOILING BELOW THE BOILING POINT OF SAID LEAN EXTRACTING AGENT, OBTAINING REGENERATOR EFFLUENT COMPRISING HYDROGEN PRODUCED IN SAID EXTRACTING AGENT REGENEREATOR, THE AFORESAID ADMIXED HYDROGEN, AND LEAN EXTRACTING AGENT, COOLING SAID REGENERATOR EFFLUENT TO A TEMPERATURE AT WHICH SAID LEAN EXTRACTING AGENT IS CONDENSED AT THE EXISTING PRESSURE, SEPARATING SAID COOLED REGENERATOR EFFLUENT INTO A GASEOUS FRACTION COMPRISING AT LEAST 90 PERCENT HYDROGEN BY VOLUME AND CONDENSED LEAN EXTRACTING AGENT, RECYCLING A PORTION OF SAID GASEOUS FRACTION TO SAID EXTRACTING AGENT REGENERATOR AND RECOVERING THE BALANCE OF SAID GASEOUS FRACTION AS HYDROGEN OF AT LEAST 90 PERCENT PURITY, AND RECYCLING AT LEAST A PORTION OF SAID CONCENSED LEAN EXTRACTING AGENT TO THE AFORESAID GAS EXTRACTOR.