Patent Publication Number: US-10781380-B2

Title: Process and apparatus for recovering hydrogen from hydroprocessed hot flash liquid

Description:
CROSS-REFERENCE TO RELATED APPLICATION 
     This application claims priority from Provisional Application No. 62/272,481 filed Dec. 29, 2015, the contents of which are hereby incorporated by reference. 
    
    
     FIELD 
     The field is the recovery of hydrogen and light hydrocarbons particularly from hydroprocessed streams. 
     BACKGROUND 
     Hydroprocessing can include processes which convert hydrocarbons in the presence of hydroprocessing catalyst and hydrogen to more valuable products. 
     Hydrocracking is a hydroprocessing process in which hydrocarbons crack in the presence of hydrogen and hydrocracking catalyst to lower molecular weight hydrocarbons. Depending on the desired output, a hydrocracking unit may contain one or more fixed beds of the same or different catalyst. Slurry hydrocracking is a slurried catalytic process used to crack residue feeds to gas oils and fuels. 
     Due to environmental concerns and newly enacted rules and regulations, saleable fuels must meet lower and lower limits on contaminates, such as sulfur and nitrogen. New regulations require essentially complete removal of sulfur from diesel. For example, the ultra low sulfur diesel (ULSD) requirement is typically less than about 10 wppm sulfur. 
     Hydrotreating is a hydroprocessing process used to remove heteroatoms such as sulfur and nitrogen from hydrocarbon streams to meet fuel specifications and to saturate olefinic or aromatic compounds. Hydrotreating can be performed at high or low pressures, but is typically operated at lower pressure than hydrocracking. 
     Hydroprocessing recovery units typically include an array of separators for cooling and depressurizing hydroprocessed effluent and separating gaseous streams from liquid streams and a stripping column for stripping hydroprocessed liquid with a stripping medium such as steam to remove unwanted hydrogen sulfide. The stripped stream then is typically heated and fractionated in a product fractionation column to recover products such as naphtha, kerosene and diesel. 
     In a refinery, hydrogen has a prime importance and recovery of hydrogen improves refinery profitability significantly. Pressure swing absorption (PSA) units are useful for purifying hydrogen by adsorbing larger molecules from the hydrogen stream at high pressure and then releasing the larger molecules at swing to lower pressure to provide a tail gas stream. 
     In many regions liquefied petroleum gas (LPG) is also important for petrochemical and fuel uses and additional recovery of LPG can also boost profit. Naphtha is also useful for fuel and petrochemical feed stock and its further recovery is desirable. 
     There is a continuing need, therefore, for improved methods of recovering hydrogen, LPG and naphtha from hydroprocessed effluent streams. 
     BRIEF SUMMARY 
     We have found that hydrogen loss from cold flash and hot flash liquid streams in a separation section can typically be 2 to 3 wt % of the hydrogen consumption in a hydroprocessing unit. Additionally, cold flash gaseous streams typically comprise 15-20 wt % of the LPG produced in the hydroprocessing unit. The process and apparatus are designed to recover these valuable components. 
     A medium pressure separator is used to recover a gaseous stream from a hydroprocessed liquid stream for hydrogen recovery in a hydrogen recovery unit. A medium pressure flash drum may further remove liquid components from the gaseous stream prior to hydrogen recovery in the hydrogen recovery unit. 
    
    
     
       BRIEF DESCRIPTION OF THE DRAWINGS 
         FIG. 1  is a simplified process flow diagram. 
         FIG. 2  is a further detail of a PSA unit of  FIG. 1 . 
         FIG. 3  is an additional embodiment of the process flow diagram of  FIG. 1 . 
     
    
    
     DEFINITIONS 
     The term “communication” means that material flow is operatively permitted between enumerated components. 
     The term “downstream communication” means that at least a portion of material flowing to the subject in downstream communication may operatively flow from the object with which it communicates. 
     The term “upstream communication” means that at least a portion of the material flowing from the subject in upstream communication may operatively flow to the object with which it communicates. 
     The term “direct communication” means that flow from the upstream component enters the downstream component without undergoing a compositional change due to physical fractionation or chemical conversion. 
     The term “bypass” means that the object is out of downstream communication with a bypassing subject at least to the extent of bypassing. 
     As used herein, the term “a component-rich stream” means that the rich stream coming out of a vessel has a greater concentration of the component than the feed to the vessel. 
     As used herein, the term “a component-lean stream” means that the lean stream coming out of a vessel has a smaller concentration of the component than the feed to the vessel. 
     The term “column” means a distillation column or columns for separating one or more components of different volatilities. Unless otherwise indicated, each column includes a condenser on an overhead of the column to condense and reflux a portion of an overhead stream back to the top of the column and a reboiler at a bottom of the column to vaporize and send a portion of a bottoms stream back to the bottom of the column. Absorber and scrubbing columns do not include a condenser on an overhead of the column to condense and reflux a portion of an overhead stream back to the top of the column and a reboiler at a bottom of the column to vaporize and send a portion of a bottoms stream back to the bottom of the column. Feeds to the columns may be preheated. The overhead pressure is the pressure of the overhead vapor at the vapor outlet of the column. The bottom temperature is the liquid bottom outlet temperature. Overhead lines and bottoms lines refer to the net lines from the column downstream of any reflux or reboil to the column unless otherwise indicated. Stripping columns omit a reboiler at a bottom of the column and instead provide heating requirements and separation impetus from a fluidized inert media such as steam. 
     As used herein, the term “True Boiling Point” (TBP) means a test method for determining the boiling point of a material which corresponds to ASTM D-2892 for the production of a liquefied gas, distillate fractions, and residuum of standardized quality on which analytical data can be obtained, and the determination of yields of the above fractions by both mass and volume from which a graph of temperature versus mass % distilled is produced using fifteen theoretical plates in a column with a 5:1 reflux ratio. 
     As used herein, the term “T5” or “T95” means the temperature at which 5 volume percent or 95 volume percent, as the case may be, respectively, of the sample boils using ASTM D-86. 
     As used herein, the term “diesel cut point” is between about 343° C. (650° F.) and about 399° C. (750° F.) using the TBP distillation method. 
     As used herein, the term “diesel boiling range” means hydrocarbons boiling in the range of between about 132° C. (270° F.) and the diesel cut point using the TBP distillation method. 
     As used herein, the term “diesel conversion” means conversion of feed to material that boils at or below the diesel cut point of the diesel boiling range. 
     As used herein, the term “separator” means a vessel which has an inlet and at least an overhead vapor outlet and a bottoms liquid outlet and may also have an aqueous stream outlet from a boot. A flash drum is a type of separator which may be in downstream communication with a separator which latter may be operated at higher pressure. 
     As used herein, the term “predominant” or “predominate” means greater than 50%, suitably greater than 75% and preferably greater than 90%. 
     DETAILED DESCRIPTION 
     In hydroprocessing units, hydrogen loss may be attributed to solution losses in hot separator liquid and cold separator liquid streams. Hydrogen from a cold flash drum vapor stream may be recovered in hydrogen recovery systems which may include a (PSA) unit. However hydrogen contained in hot flash and cold flash liquid streams can go into a stripper column and subsequently into the overhead off gas stream from the process unit. The stripper off gas stream is typically blended into the refinery fuel gas and burned. Hence, recovering hydrogen from the hot flash and cold flash liquid streams can reduce overall hydrogen consumption. 
     We have found that the bulk of the hydrogen is contained in the hot flash liquid stream and the bulk of the LPG is contained in the cold flash liquid stream. The cold flash drum gaseous stream typically contains 15-20 wt % of the hydroprocessed LPG based on reactor yields. LPG lost from the cold flash drum gaseous stream can be recovered as discovered. Moreover, we have found PSA tail gas contains LPG and naphtha hydrocarbons that can be recovered economically. 
     The apparatus and process  10  for hydroprocessing hydrocarbons comprise a hydroprocessing unit  12 , a separation section  30 , a hydrogen recovery unit  100 , a LPG recovery section  150  and a product recovery unit  14 . A hydrocarbonaceous stream in hydrocarbon line  16  and a hydrogen stream in hydrogen line  18  are fed to the hydroprocessing unit  12 . 
     A recycle hydrogen stream in recycle hydrogen line  20  may be supplemented by a make-up hydrogen stream from line  22  to provide the hydrogen stream in hydrogen line  18 . The hydrogen stream may join the hydrocarbonaceous stream in feed line  16  to provide a hydrocarbon feed stream in feed line  23 . The hydrocarbon feed stream in line  23  may be heated in a fired heater and fed to the hydroprocessing reactor  24 . The hydrocarbon feed stream is hydroprocessed in the hydroprocessing reactor  24 . 
     In one aspect, the process and apparatus described herein are particularly useful for hydroprocessing a hydrocarbon feed stream comprising a hydrocarbonaceous feedstock. Illustrative hydrocarbon feedstocks include hydrocarbonaceous streams having initial boiling points (IBP) above about 288° C. (550° F.), such as atmospheric gas oils, vacuum gas oil (VGO) having T5 and T95 between about 315° C. (600° F.) and about 600° C. (1100° F.), deasphalted oil, coker distillates, straight run distillates, pyrolysis-derived oils, high boiling synthetic oils, cycle oils, hydrocracked feeds, catalytic cracker distillates, atmospheric residue having an IBP at or above about 343° C. (650° F.) and vacuum residue having an IBP above about 510° C. (950° F.). 
     Hydroprocessing that occurs in the hydroprocessing unit  12  may be hydrocracking or hydrotreating. Hydrocracking refers to a process in which hydrocarbons crack in the presence of hydrogen to lower molecular weight hydrocarbons. Hydrocracking is the preferred process in the hydroprocessing unit  12 . Consequently, the term “hydroprocessing” will include the term “hydrocracking” herein. Hydrocracking also includes slurry hydrocracking in which resid feed is mixed with catalyst and hydrogen to make a slurry and cracked to lower boiling products. 
     Hydroprocessing that occurs in the hydroprocessing unit may also be hydrotreating. Hydrotreating is a process wherein hydrogen is contacted with hydrocarbon in the presence of suitable catalysts which are primarily active for the removal of heteroatoms, such as sulfur, nitrogen and metals from the hydrocarbon feedstock. In hydrotreating, hydrocarbons with double and triple bonds may be saturated. Aromatics may also be saturated. Some hydrotreating processes are specifically designed to saturate aromatics. The cloud point or pour point of the hydrotreated product may also be reduced by hydroisomerization. A hydrocracking reactor may be preceded by a hydrotreating reactor and a separator (not shown) to remove sulfur and nitrogen contaminants from the feed to the hydrocracking reactor. 
     The hydroprocessing reactor  24  may be a fixed bed reactor that comprises one or more vessels, single or multiple beds of catalyst in each vessel, and various combinations of hydrotreating catalyst and/or hydrocracking catalyst in one or more vessels. It is contemplated that the hydroprocessing reactor  24  be operated in a continuous liquid phase in which the volume of the liquid hydrocarbon feed is greater than the volume of the hydrogen gas. The hydroprocessing reactor  24  may also be operated in a conventional continuous gas phase, a moving bed or a fluidized bed hydroprocessing reactor. 
     If the hydroprocessing reactor  24  is operated as a hydrocracking reactor, it may provide total conversion of at least about 20 vol % and typically greater than about 60 vol % of the hydrocarbon feed to products boiling below the diesel cut point. A hydrocracking reactor may operate at partial conversion of more than about 30 vol % or full conversion of at least about 90 vol % of the feed based on total conversion. A hydrocracking reactor may be operated at mild hydrocracking conditions which will provide about 20 to about 60 vol %, preferably about 20 to about 50 vol %, total conversion of the hydrocarbon feed to product boiling below the diesel cut point. If the hydroprocessing reactor  24  is operated as a hydrotreating reactor, it may provide conversion per pass of about 10 to about 30 vol %. 
     If the hydroprocessing reactor  24  is a hydrocracking reactor, the first vessel or bed in the hydrocracking reactor  24  may include hydrotreating catalyst for the purpose of saturating, demetallizing, desulfurizing or denitrogenating the hydrocarbon feed before it is hydrocracked with hydrocracking catalyst in subsequent vessels or beds in the hydrocracking reactor  24 . If the hydrocracking reactor is a mild hydrocracking reactor, it may contain several beds of hydrotreating catalyst followed by a fewer beds of hydrocracking catalyst. If the hydroprocessing reactor  24  is a slurry hydrocracking reactor, it may operate in a continuous liquid phase in an upflow mode and will appear different than in  FIG. 1  which depicts a fixed bed reactor. If the hydroprocessing reactor  24  is a hydrotreating reactor it may comprise more than one vessel and multiple beds of hydrotreating catalyst. The hydrotreating reactor may also contain hydrotreating catalyst that is suited for saturating aromatics, hydrodewaxing and hydroisomerization.
     A hydrocracking catalyst may utilize amorphous silica-alumina bases or low-level zeolite bases combined with one or more Group VIII or Group VIB metal hydrogenating components if mild hydrocracking is desired to produce a balance of middle distillate and gasoline. In another aspect, when middle distillate is significantly preferred in the converted product over gasoline production, partial or full hydrocracking may be performed in the first hydrocracking reactor  24  with a catalyst which comprises, in general, any crystalline zeolite cracking base upon which is deposited a Group VIII metal hydrogenating component. Additional hydrogenating components may be selected from Group VIB for incorporation with the zeolite base.   

     The zeolite cracking bases are sometimes referred to in the art as molecular sieves and are usually composed of silica, alumina and one or more exchangeable cations such as sodium, magnesium, calcium, rare earth metals, etc. They are further characterized by crystal pores of relatively uniform diameter between about 4 and about 14 Angstroms (10 −10  meters). It is preferred to employ zeolites having a relatively high silica/alumina mole ratio between about 3 and about 12. Suitable zeolites found in nature include, for example, mordenite, stilbite, heulandite, ferrierite, dachiardite, chabazite, erionite and faujasite. Suitable synthetic zeolites include, for example, the B, X, Y and L crystal types, e.g., synthetic faujasite and mordenite. The preferred zeolites are those having crystal pore diameters between about 8 and 12 Angstroms (10 −10  meters), wherein the silica/alumina mole ratio is about 4 to 6. One example of a zeolite falling in the preferred group is synthetic Y molecular sieve. 
     The natural occurring zeolites are normally found in a sodium form, an alkaline earth metal form, or mixed forms. The synthetic zeolites are nearly always prepared first in the sodium form. In any case, for use as a cracking base it is preferred that most or all of the original zeolitic monovalent metals be ion-exchanged with a polyvalent metal and/or with an ammonium salt followed by heating to decompose the ammonium ions associated with the zeolite, leaving in their place hydrogen ions and/or exchange sites which have actually been decationized by further removal of water. Hydrogen or “decationized” Y zeolites of this nature are more particularly described in U.S. Pat. No. 3,130,006. 
     Mixed polyvalent metal-hydrogen zeolites may be prepared by ion-exchanging first with an ammonium salt, then partially back exchanging with a polyvalent metal salt and then calcining. In some cases, as in the case of synthetic mordenite, the hydrogen forms can be prepared by direct acid treatment of the alkali metal zeolites. In one aspect, the preferred cracking bases are those which are at least about 10 percent, and preferably at least about 20 percent, metal-cation-deficient, based on the initial ion-exchange capacity. In another aspect, a desirable and stable class of zeolites is one wherein at least about 20 percent of the ion exchange capacity is satisfied by hydrogen ions. 
     The active metals employed in the preferred hydrocracking catalysts of the present invention as hydrogenation components are those of Group VIII, i.e., iron, cobalt, nickel, ruthenium, rhodium, palladium, osmium, iridium and platinum. In addition to these metals, other promoters may also be employed in conjunction therewith, including the metals of Group VIB, e.g., molybdenum and tungsten. The amount of hydrogenating metal in the catalyst can vary within wide ranges. Broadly speaking, any amount between about 0.05 percent and about 30 percent by weight may be used. In the case of the noble metals, it is normally preferred to use about 0.05 to about 2 wt % noble metal. 
     The method for incorporating the hydrogenating metal is to contact the base material with an aqueous solution of a suitable compound of the desired metal wherein the metal is present in a cationic form. Following addition of the selected hydrogenating metal or metals, the resulting catalyst powder is then filtered, dried, pelleted with added lubricants, binders or the like if desired, and calcined in air at temperatures of, e.g., about 371° C. (700° F.) to about 648° C. (1200° F.) in order to activate the catalyst and decompose ammonium ions. Alternatively, the base component may first be pelleted, followed by the addition of the hydrogenating component and activation by calcining. 
     The foregoing catalysts may be employed in undiluted form, or the powdered catalyst may be mixed and copelleted with other relatively less active catalysts, diluents or binders such as alumina, silica gel, silica-alumina cogels, activated clays and the like in proportions ranging between about 5 and about 90 wt %. These diluents may be employed as such or they may contain a minor proportion of an added hydrogenating metal such as a Group VIB and/or Group VIII metal. Additional metal promoted hydrocracking catalysts may also be utilized in the process of the present invention which comprises, for example, aluminophosphate molecular sieves, crystalline chromosilicates and other crystalline silicates. Crystalline chromosilicates are more fully described in U.S. Pat. No. 4,363,718. 
     By one approach, the hydrocracking conditions may include a temperature from about 290° C. (550° F.) to about 468° C. (875° F.), preferably 343° C. (650° F.) to about 445° C. (833° F.), a pressure from about 4.8 MPa (gauge) (700 psig) to about 20.7 MPa (gauge) (3000 psig), a liquid hourly space velocity (LHSV) from about 0.4 to less than about 2.5 hr −1  and a hydrogen rate of about 421 (2,500 scf/bbl) to about 2,527 Nm 3 /m 3  oil (15,000 scf/bbl). If mild hydrocracking is desired, conditions may include a temperature from about 315° C. (600° F.) to about 441° C. (825° F.), a pressure from about 5.5 MPa (gauge) (800 psig) to about 13.8 MPa (gauge) (2000 psig) or more typically about 6.9 MPa (gauge) (1000 psig) to about 11.0 MPa (gauge) (1600 psig), a liquid hourly space velocity (LHSV) from about 0.5 to about 2 hr −1  and preferably about 0.7 to about 1.5 hr −1  and a hydrogen rate of about 421 Nm 3 /m 3  oil (2,500 scf/bbl) to about 1,685 Nm 3 /m 3  oil (10,000 scf/bbl). 
     Suitable hydrotreating catalysts for use in the present invention are any known conventional hydrotreating catalysts and include those which are comprised of at least one Group VIII metal, preferably iron, cobalt and nickel, more preferably cobalt and/or nickel and at least one Group VI metal, preferably molybdenum and tungsten, on a high surface area support material, preferably alumina. Other suitable hydrotreating catalysts include zeolitic catalysts, as well as noble metal catalysts where the noble metal is selected from palladium and platinum. It is within the scope of the present description that more than one type of hydrotreating catalyst be used in the same hydrotreating reactor  24 . The Group VIII metal is typically present in an amount ranging from about 2 to about 20 wt %, preferably from about 4 to about 12 wt %. The Group VI metal will typically be present in an amount ranging from about 1 to about 25 wt %, preferably from about 2 to about 25 wt %. 
     Preferred hydrotreating reaction conditions include a temperature from about 290° C. (550° F.) to about 455° C. (850° F.), suitably 316° C. (600° F.) to about 427° C. (800° F.) and preferably 343° C. (650° F.) to about 399° C. (750° F.), a pressure from about 2.1 MPa (gauge) (300 psig), preferably 4.1 MPa (gauge) (600 psig) to about 20.6 MPa (gauge) (3000 psig), suitably 12.4 MPa (gauge) (1800 psig), preferably 6.9 MPa (gauge) (1000 psig), a liquid hourly space velocity of the fresh hydrocarbonaceous feedstock from about 0.1 hr −1 , suitably 0.5 hr −1 , to about 4 hr −1 , preferably from about 1.5 to about 3.5 hr −1 , and a hydrogen rate of about 168 Nm 3 /m 3  (1,000 scf/bbl), to about 1,011 Nm 3 /m 3  oil (6,000 scf/bbl), preferably about 168 Nm 3 /m 3  oil (1,000 scf/bbl) to about 674 Nm 3 /m 3  oil (4,000 scf/bbl), with a hydrotreating catalyst or a combination of hydrotreating catalysts. 
     The hydroprocessing reactor  24  provides a hydroprocessing effluent stream that exits the hydroprocessing reactor  24  in a hydroprocessing effluent line  26 . The hydroprocessing effluent stream comprises material that will be separated in the separation section  30  comprising one or more separators into a liquid hydroprocessed stream and a gaseous hydroprocessed stream. The separation section  30  is in downstream communication with the hydroprocessing reactor  24 . 
     The hydroprocessing effluent stream in hydroprocessing effluent line  26  may in an aspect be heat exchanged with the hydrocarbon feed stream in line  16  to be cooled before entering a hot separator  32 . The hot separator separates the hydroprocessing effluent to provide a hydrocarbonaceous, gaseous hot gaseous stream in an overhead line  34  and a hydrocarbonaceous hot liquid stream in a bottoms line  36 . The hot separator  32  may be in downstream communication with the hydroprocessing reactor  24 . The hot separator  32  operates at about 177° C. (350° F.) to about 371° C. (700° F.) and preferably operates at about 232° C. (450° F.) to about 315° C. (600° F.). The hot separator  32  may be operated at a slightly lower pressure than the hydroprocessing reactor  24  accounting for pressure drop through intervening equipment. The hot separator may be operated at pressures between about 3.4 MPa (gauge) (493 psig) and about 20.4 MPa (gauge) (2959 psig). The liquid hydrocarbonaceous hot liquid stream  36  may have a temperature of the operating temperature of the hot separator  32 . 
     The hot gaseous stream in the overhead line  34  may be cooled before entering a cold separator  38 . As a consequence of the reactions taking place in the hydroprocessing reactor  24  wherein nitrogen, chlorine and sulfur are removed from the feed, ammonia and hydrogen sulfide are formed. At a characteristic sublimation temperature, ammonia and hydrogen sulfide will combine to form ammonium bisulfide and ammonia and chlorine will combine to form ammonium chloride. Each compound has a characteristic sublimation temperature that may allow the compound to coat equipment, particularly heat exchange equipment, impairing its performance. To prevent such deposition of ammonium bisulfide or ammonium chloride salts in the line  34  transporting the hot gaseous stream, a suitable amount of wash water may be introduced into line  34  upstream of a cooler at a point in line  34  where the temperature is above the characteristic sublimation temperature of either compound. 
     The cold gaseous stream may be separated in the cold separator  38  to provide a vaporous cold gaseous stream comprising a hydrogen-rich gas stream in an overhead line  40  and a liquid cold liquid stream in a cold bottoms line  42 . The cold separator  38  serves to separate hydrogen from hydrocarbon in the hydroprocessing effluent for recycle to the hydroprocessing reactor  24  in the cold overhead line  40 . The cold separator  38 , therefore, is in downstream communication with the overhead line  34  of the hot separator  32  and the hydroprocessing reactor  24 . The cold separator  38  may be operated at about 100° F. (38° C.) to about 150° F. (66° C.), suitably about 115° F. (46° C.) to about 145° F. (63° C.), and just below the pressure of the hydroprocessing reactor  24  and the hot separator  32  accounting for pressure drop through intervening equipment to keep hydrogen and light gases in the overhead and normally liquid hydrocarbons in the bottoms. The cold separator  38  may be operated at pressures between about 3 MPa (gauge) (435 psig) and about 20 MPa (gauge) (2,901 psig). The cold separator  38  may also have a boot for collecting an aqueous phase. The cold liquid stream may have a temperature of the operating temperature of the cold separator  38 . 
     The cold gaseous stream in the cold overhead line  40  is rich in hydrogen. Thus, hydrogen can be recovered from the cold gaseous stream. The cold gaseous stream in the cold overhead line  40  may be passed through a trayed or packed recycle scrubbing column  56  where it is scrubbed by means of a scrubbing extraction liquid such as an aqueous amine solution to remove acid gases including hydrogen sulfide and carbon dioxide by extracting them into the aqueous solution. Preferred lean amines include alkanolamines DEA, MEA, and MDEA. Other amines can be used in place of or in addition to the preferred amines. The lean amine contacts the cold gaseous stream and absorbs acid gas contaminants such as hydrogen sulfide and carbon dioxide. The resultant “sweetened” cold gaseous stream is taken out from an overhead outlet of the recycle scrubber column  56  in a recycle scrubber overhead line  58 , and a rich amine is taken out from the bottoms at a bottom outlet of the recycle scrubber column in a recycle scrubber bottoms line. The spent scrubbing liquid from the bottoms may be regenerated and recycled back to the recycle scrubbing column  56 . The scrubbed hydrogen-rich stream emerges from the scrubber via an overhead line  58  and may be compressed in a recycle compressor to provide a recycle hydrogen stream in line  20 . The recycle hydrogen stream in line  20  may be supplemented with make-up hydrogen stream in make-up line  22  to provide the hydrogen stream in hydrogen line  18 . A portion of the material in line  20  may be routed to the intermediate catalyst bed outlets in the hydroprocessing reactor  24  to control the inlet temperature of the subsequent catalyst bed (not shown). The recycle scrubbing column  56  may be operated with a gas inlet temperature between about 38° C. (100° F.) and about 66° C. (150° F.) and an overhead pressure of about 3 MPa (gauge) (435 psig) to about 20 MPa (gauge) (2900 psig). 
     The hydrocarbonaceous hot liquid stream in the hot bottoms line  36  may be fractionated as a hot hydroprocessing effluent stream in the product recovery unit  14 . In an aspect, the hot liquid stream in the bottoms line  36  may be let down in pressure and flashed in a hot flash drum  44  to provide a hot flash gaseous stream of light ends in an overhead line  46  and a hot flash liquid stream in a hot flash bottoms line  48 . The hot flash drum  44  may be any separator that splits the liquid hydroprocessing effluent into vapor and liquid fractions. The hot flash drum  44  may be in downstream communication with the hot bottoms line  36 . The hot flash drum  44  may be operated at the same temperature as the hot separator  32  but at a lower pressure of between about 1.4 MPa (gauge) (200 psig) and about 6.9 MPa (gauge) (1000 psig), suitably no more than about 3.4 MPa (gauge) (500 psig). The hot flash liquid stream in bottoms line  48  may be further fractionated in the product recovery unit  14 . The hot flash liquid stream in the hot flash bottoms line  48  may have a temperature of the operating temperature of the hot flash drum  44 . 
     In an aspect, the cold liquid stream in the cold bottoms line  42  may be fractionated as a cold hydroprocessing effluent stream in the product recovery unit  14 . In a further aspect, the cold liquid stream may be let down in pressure and flashed in a cold flash drum  50  to separate the cold liquid stream in the bottoms line  42 . The cold flash drum  50  may be any separator that splits hydroprocessing effluent into vapor and liquid fractions. The cold flash drum may be in downstream communication with a bottoms line  42  of the cold separator  38 . 
     In a further aspect, the hot flash gaseous stream in the hot flash overhead line  46  may be fractionated as a hot hydroprocessing effluent stream in the product recovery unit  14 . In a further aspect, the hot flash gaseous stream may be cooled and also separated in the cold flash drum  50 . The cold flash drum  50  may separate the cold liquid stream in line  42  and the hot flash gaseous stream in overhead line  46  to provide a cold flash gaseous stream in overhead line  52  and a cold flash liquid stream in a bottoms line  54 . In an aspect, light gases such as hydrogen sulfide may be stripped from the cold flash liquid stream. Accordingly, a cold stripping column  92  may be in downstream communication with the cold flash drum  50  and the cold flash bottoms line  54 . The cold flash drum  50  may be in downstream communication with the bottoms line  42  of the cold separator  38 , the overhead line  46  of the hot flash drum  44  and the hydroprocessing reactor  24 . The cold liquid stream in cold bottoms line  42  and the hot flash gaseous stream in the hot flash overhead line  46  may enter into the cold flash drum  50  either together or separately. In an aspect, the hot flash overhead line  46  joins the cold bottoms line  42  and feeds the hot flash gaseous stream and the cold liquid stream together to the cold flash drum  50  in line  47 . The cold flash drum  50  may be operated at the same temperature as the cold separator  38  but typically at a lower pressure of between about 1.4 MPa (gauge) (200 psig) and about 6.9 MPa (gauge) (1000 psig) and preferably no higher than 3.1 MPa (gauge) (450 psig). A flashed aqueous stream may be removed from a boot in the cold flash drum  50 . The cold flash liquid stream in the cold flash bottoms line  54  may have the same temperature as the operating temperature of the cold flash drum  50 . The cold flash gaseous stream in the cold flash overhead line  52  contains substantial hydrogen and may be transported to the hydrogen recovery unit  100  for recovering hydrogen from the cold flash gaseous stream. 
     We have found that substantial hydrogen may be dissolved in the hot flash liquid stream in the hot flash bottoms line  48 , more than in the cold flash liquid stream in the cold flash bottoms line  42 , which can be lost in the stripper off gas. Hence, we propose to recover hydrogen from the hot flash liquid stream in a hydrogen recovery unit  100 . The hot flash liquid stream in the hot flash bottoms line  48  is separated in a medium pressure separator  60  to provide a medium gaseous stream in a medium overhead line  62  and a medium liquid stream in medium bottoms line  64 . The medium pressure separator  60  may be in downstream communication with the hot flash bottoms line  48  and the hydrogen recovery unit  100  may be in downstream communication with the medium overhead line  62  of the medium pressure separator  60 . The medium pressure separator  60  may be operated at the same temperature as the hot flash drum  44  but typically at a lower pressure of between about 0.97 MPa (gauge) (140 psig) and about 1.59 MPa (gauge) (230 psig) and preferably no higher than 1.38 MPa (gauge) (200 psig). Hydrogen may be recovered from the medium gaseous stream in the hydrogen recovery unit  100 . Moreover, the medium liquid stream may be stripped in a stripper column  90 . The stripper column  90  may comprise a cold stripper  92  and a hot stripper  94 . The medium liquid stream in the medium bottoms line  64  may be stripped in the hot stripper  94 . 
     In a further embodiment, to further remove liquid hydrocarbons from the medium gaseous stream, the medium gaseous stream in the medium overhead line  62  may be flashed in a medium pressure flash drum  66  to provide a medium flash gaseous stream in medium flash overhead line  68  and a medium flash liquid stream in a medium flash bottoms line  70 . The medium flash drum may be in downstream communication with a medium pressure overhead line  62 . The hydrogen recovery unit  100  may be in downstream communication with an overhead line  68  of the medium pressure flash drum  66 . In an aspect, the medium gaseous stream may be cooled before separation in said medium pressure flash drum  66 . The medium pressure flash drum  66  may be operated at the same pressure as the medium pressure separator  60  but typically at a temperature of between about 40° C. and about 60° C. The medium pressure flash drum should be operated at a pressure of about 140 kPa (20 psi) to about 410 kPa (60 psi) above that of the stripper column  90 . The medium flash gaseous stream may be transported to the hydrogen recovery unit  100 . The medium flash liquid stream in the medium flash bottoms line  70  may be stripped in the stripper column  90 . In an aspect, the medium flash liquid stream may be heated and may be stripped in the cold stripper column  92 . In a further aspect, the medium flash liquid stream in the medium flash bottoms line  70  may be mixed with the cold flash liquid stream in the cold flash bottoms line  54  and a compressed, condensed stream in a knock out bottoms line  207 . The mixed stream in mixed line  74  may be heated together and stripped in the cold stripper column  92 . 
     In an aspect, the medium flash gaseous stream may be compressed in a separator compressor  72  to the hydrogen recovery pressure of about 1.4 MPa (gauge) (200 psig) to about 3.1 MPa (gauge) (450 psig) to provide a compressed medium gaseous stream in route to the hydrogen recovery unit  100 . The compressed, medium gaseous stream may be fed from the separator compressor  72  to a compressor knock out drum  71  to remove a compressed, condensed stream in a medium knockout bottoms line  73  for transport to the stripper column  90 . The knock out drum  71  may be in downstream communication with the separator compressor  72 . The medium knockout bottoms line  73  may transport a medium knockout liquid stream to be mixed with the medium flash liquid stream in a medium flash bottoms line  70  for transport to the stripper column  90 . The remaining compressed, medium gaseous stream may be transported from the knock out drum  71  to the hydrogen recovery unit  100  in a medium knock out overhead line  77 . The compressed, medium gaseous stream in the medium knockout overhead line  77  comprises substantial hydrogen such as about 70 to about 90 mole % which may transported to the hydrogen recovery unit  100  for hydrogen recovery. In an aspect, the cold flash gaseous stream in the cold flash overhead line  52  may be mixed with said medium vapor stream perhaps after compression to the cold flash drum pressure and transported to the hydrogen recovery unit  100 . 
     The compressed, medium gaseous stream in the medium knockout overhead line  77  may be passed through a trayed or packed separator scrubbing column  76  in the hydrogen recovery unit  100 . Alternatively or conjunctively, the cold flash gaseous stream in the cold flash overhead line  52  may be passed through a trayed or packed separator scrubbing column  76  in the hydrogen recovery unit  100 . In an aspect, the compressed, medium gaseous stream in medium knockout overhead line  77  may be combined with the cold flash gaseous stream and fed to the hydrogen recovery unit  100  together in a joint line  82  as a joint gaseous stream. In the separator scrubbing column  76 , the gaseous stream is scrubbed by means of a scrubbing extraction liquid such as an aqueous amine solution to remove acid gases including hydrogen sulfide and carbon dioxide by extracting them into the aqueous amine solution. The gaseous stream enters the separator scrubbing column  76  at an inlet near a bottom and flows upwardly, while a lean amine stream in a solvent line  75  enters the scrubber column at an inlet near a top and flows downwardly. Preferred lean amines include alkanolamines diethanolamine (DEA), monoethanolamine (MEA), and methyldiethanolamine (MDEA). However, other amines or solvents can be used. The lean amine stream contacts the scrubbed gaseous stream and absorbs acid gas contaminants such as hydrogen sulfide and carbon dioxide. The resultant “sweetened” gaseous stream is taken out from an overhead outlet of the separator scrubber column  76  in a separator scrubber overhead line  78 , and a rich amine is taken out from the bottoms at a bottom outlet of the separator scrubber column in scrubber bottoms line  79 . The rich amine stream may undergo regeneration to remove the hydrogen sulfide, for processing to generate elemental sulfur, and other gases and recycled back to the separator scrubbing column  76 . The separator scrubbing column  76  may be operated with a gas inlet temperature between about 30° C. (86° F.) and about 66° C. (150° F.) and an overhead pressure of about 1.4 MPa (gauge) (200 psig) to about 3.1 MPa (gauge) (450 psig). 
     A scrubbed gaseous stream emerges from the scrubber via a separator scrubber overhead line  78  and may be passed to a hydrogen recovery unit  100  which may comprise a membrane for concentrating a hydrogen stream. In an aspect, the hydrogen recovery unit  100  comprises a pressure swing adsorption (PSA) unit  102 . The PSA unit  102  comprises a plurality of adsorbent beds or vessels for hydrogen recovery. The PSA unit  102  may be in downstream communication with an overhead line of the separation section  30 . For example, the PSA unit may be in downstream communication with the hot overhead line  34 , the hot flash overhead line  46 , the cold flash overhead line  52 , the medium overhead line  62 , the medium flash overhead line  68  and the medium knockout overhead line  77 . 
     In the PSA unit  102 , impure gases are adsorbed from hydrogen in the scrubbed medium gaseous stream and/or the scrubbed cold flash gaseous stream. In an embodiment, the hydrogen in the gaseous stream can be purified in a pressure swing adsorption (PSA) unit  102  shown in  FIG. 2  to provide a hydrogen rich gaseous stream having a reduced concentration of hydrogen sulfide, ammonia, amines and hydrocarbons. The pressure swing adsorption process separates hydrogen from larger molecules in the scrubber overhead line  78 . The larger impurities are adsorbed on an adsorbent at a high adsorption pressure while allowing passage of the smaller hydrogen molecules. Pressure reduction is effected to a lower desorption pressure to desorb the adsorbed larger species. It is generally desirable to employ the PSA process in multiple bed systems such as those described in U.S. Pat. No. 3,430,418, in which at least four adsorption beds are employed. The PSA process is carried out in such systems on a cyclical basis, employing a processing sequence. 
     Referring to  FIG. 2 , the PSA unit  102  may have four beds  108 - 114  having inlet ends  108   a - 114   a  and outlet ends  108   b - 114   b , respectively. In the first step, the scrubbed gaseous stream in the overhead line  78  is fed to an inlet end  108   a  of a first adsorbent bed  108  at high adsorption pressure such as about 1 MPa (150 psia) to about 1.7 MPa (250 psia) to adsorb adsorbable species onto the adsorbent with passage of purified product hydrogen gas to a discharge end  108   b  of the bed  108  for 5 to 10 minutes. A purified hydrogen stream may pass from the PSA unit  102  through product line  104  with a greater hydrogen purity than in the overhead line  78 . Feed flow is terminated to the first bed  108  before the larger molecules break through to the discharge end  108   b  of the first bed. Second, the first bed  108  is cocurrently depressurized to an intermediate pressure such as about 0.7 MPa (100 psia) to about 1 MPa (150 psia) for 0.5 to 2 minutes by releasing void space gas from the discharge end  108   b  of the first bed to a discharge end  110   b  of a second bed  110  thereby repressurizing the second bed which has just been purged of desorbed larger molecules. Further cocurrent depressurization of the first bed  108  to a pressure of about 0.7 MPa (50 psia) to about 0.5 MPa (75 psia) can occur by releasing remaining void space gas to a discharge end  112   b  of a third bed  112  to purge the third bed of desorbed larger molecules for 5 to 10 minutes. In a third step, the inlet  108   a  to the first bed  108  is opened in a countercurrent depressurization or blow down step, in which gas departs the first bed through the inlet end  108   a  leaving the first bed  108  at sufficiently low pressure such as about 34.5 kPa (5 psia) to about 172 kPa (25 psia) to desorb adsorbed species from the adsorbent for about 0.5 to 2 minutes. Desorbed species are released through the inlet  108   a  and recovered in a tail gas line  106  with a greater concentration of adsorbable species than in the feed line  78 . In a fourth step, void space gas from a fourth bed  114  may be released through a discharge end  114   b  thereof and fed through the discharge end  108   b  of the first bed  108  to purge out the desorbed species. In a last step, void space gas from the second bed  110  is fed from its discharge end  110   b  into the discharge end  108   b  of the first bed  108  to repressurize the first bed. Product gas from the discharge end  112   b  of the third bed  112  is then fed into the discharge end  108   b  of the first bed  108  to achieve adsorption pressure in the first bed  108  of about 1 MPa (150 psia) to about 1.7 MPa (250 psia) for 5 to 10 minutes. Since the first bed  108  is now at adsorption pressure, the cycle in the first bed begins anew. The same process sequence is operated with the other beds  110 - 114 , with differences relating to the position of the bed  110 - 114  in the order. 
     A suitable adsorbent may be activated calcium zeolite A. Purified hydrogen with a hydrogen concentration greater and LPG concentration less than the gaseous stream in the scrubber overhead line  78 , the medium knockout overhead line  77 , the cold flash overhead line  52  and the joint line  82 , can be transported in line  104  for recycle or transport to the hydrogen header for use anywhere in a refinery. The tail gas line  106  contains hydrocarbons with a reduced concentration of hydrogen and an increased concentration of LPG hydrocarbons relative to the concentration in the scrubber overhead line  78 , the medium knockout overhead line  77 , the cold flash overhead line  52  and the joint line  82 . Because the tail gas stream in tail gas line  106  comprises substantial LPG hydrocarbons, it may be compressed in a tail gas compressor  142  and transported to an LPG recovery section  150 . 
     The product recovery section  14  may include a stripping column  90  and a fractionation column  130 . The stripping column  90  may be in downstream communication with a bottoms line in the separation section  30 . For example, the stripping column  90  may be in downstream communication with the hot bottoms line  36 , the hot flash bottoms line  48 , the cold bottoms line  42 , the cold flash bottoms line  54 , the medium bottoms line  64  and the medium flash bottom line  70 . In an aspect, the stripping column  90  may comprise a cold stripping column  92  and a hot stripping column  94 . The cold stripping column  92  may be in downstream communication with the hydroprocessing reactor  24 , the cold bottoms line  42  and, in an aspect, the cold flash bottoms line  54  for stripping a cold hydroprocessing effluent stream. The hot stripping column  94  may be in downstream communication with the hydroprocessing reactor  24  and the hot bottoms line  36  and, in an aspect, the hot flash bottoms line  48  for stripping a hot hydroprocessing effluent stream which is hotter than the cold hydroprocessing effluent stream. In an aspect, the cold hydroprocessing effluent stream may be the cold flash liquid stream in the cold flash bottoms line  54  which may be mixed with the medium flash liquid stream in medium flash bottoms line  70 . The hot hydroprocessing effluent stream may be the hot flash liquid stream in hot flash bottoms line  48 . In an aspect, the hot hydroprocessed effluent may be the medium liquid stream in medium bottoms line  64 . The hot hydroprocessing effluent stream is hotter than the cold hydroprocessing effluent stream, by at least 25° C. and preferably at least 50° C. 
     The cold hydroprocessing effluent stream which in an aspect may be in the mixed line  74  may be heated and fed to the cold stripping column  92  at an inlet which may be in the top half of the column. The cold hydroprocessing effluent stream which comprises at least a portion of the liquid hydroprocessing effluent may be stripped in the cold stripping column  92  with a cold stripping media which is an inert gas such as steam from a cold stripping media line  96  to provide a cold vapor stream of naphtha, hydrogen, hydrogen sulfide, steam and other gases in an overhead line  120 . At least a portion of the cold vapor stream may be condensed and separated in a receiver  122 . A stripper net overhead line  124  from the receiver  122  carries a stripper off gas stream for further treating. Unstabilized liquid naphtha from the bottoms of the receiver  122  may be split between a reflux portion refluxed to the top of the cold stripping column  92  and a stripper overhead liquid stream which may be transported in a stripper receiver bottoms line  126  to further naphtha, LPG and hydrogen recovery. A sour water stream (not shown) may be collected from a boot of the overhead receiver  122 . 
     The cold stripping column  92  may be operated with a bottoms temperature between about 149° C. (300° F.) and about 288° C. (550° F.), preferably about 260° C. (500° F.), and an overhead pressure of about 0.35 MPa (gauge) (50 psig), preferably about 0.70 MPa (gauge) (100 psig), to about 2.0 MPa (gauge) (290 psig). The temperature in the overhead receiver  122  ranges from about 38° C. (100° F.) to about 66° C. (150° F.) and the pressure is essentially the same as in the overhead of the cold stripping column  92 . 
     The cold stripped stream in a bottoms line  128  may comprise predominantly naphtha and kerosene boiling materials. Consequently, the cold stripped stream in cold stripped bottoms line  128  may be heated and fed to a fractionation column  130 . The fractionation column  130  is in downstream communication with the cold stripped bottoms line  128  of the cold stripping column  92 . In an aspect, the product fractionation column  130  may comprise more than one fractionation column. 
     The hot flash liquid stream which may be in the hot flash bottoms line  48  may be fed to the hot stripping column  94  near the top thereof. In an aspect, the hot flash liquid stream may be separated in the medium pressure separator  60  and the medium liquid stream in medium bottoms line  64  may be fed to the hot stripper column  94 . The hot flash liquid stream or the medium liquid stream may be stripped in the hot stripping column  94  with a hot stripping media which is an inert gas such as steam from a line  98  to provide a hot stripper gas stream of naphtha, hydrogen, hydrogen sulfide, steam and other gases in a hot stripper overhead line  118 . The overhead line  118  may be condensed and a portion refluxed to the hot stripping column  94 . However, in the embodiment of  FIG. 1 , the hot stripper gas stream in the overhead line  118  from the overhead of the hot stripping column  94  may be fed into the cold stripping column  92  directly in an aspect without first condensing or refluxing. The inlet for the cold hydroprocessing effluent stream may be at a higher elevation than the inlet for the overhead line  118 . The hot stripping column  94  may be operated with a bottoms temperature between about 160° C. (320° F.) and about 360° C. (680° F.) and an overhead pressure of about 0.35 MPa (gauge) (50 psig), preferably about 0.70 MPa (gauge) (100 psig), to about 2.0 MPa (gauge) (292 psig). 
     A hydroprocessed hot stripped stream is produced in a hot stripped bottoms line  146 . At least a portion of the hot stripped bottoms stream in the hot stripped bottoms line  146  may be heated and fed to the product fractionation column  130 . Consequently, the product fractionation column  130  may be in downstream communication with the hot stripped bottoms line  146  of the hot stripping column  94  and the hot stripping column  94 . The hot stripped stream in line  146  is at a hotter temperature than the cold stripped stream in line  128 . 
     The product fractionation column  130  may be in downstream communication with the cold stripping column  92  and the hot stripping column  94  and may comprise more than one fractionation column for separating stripped streams into product streams. The product fractionation column  130  may strip the cold stripped stream and the hot stripped stream with inert stripping media such as steam from line  132  to provide several product streams. The product streams from the product fractionation column  130  may include an overhead light naphtha stream in a net overhead line  144 , a heavy naphtha stream in line  136  from a side cut outlet, a kerosene stream carried in line  138  from a side cut outlet and a diesel stream from a side outlet  140 . An unconverted oil stream may be provided in a bottoms line  142  which may be recycled to the hydroprocessing reactor  24 . Heat may be removed from the fractionation column  130  by cooling at least a portion of the product streams and sending a portion of each cooled stream back to the fractionation column. These product streams may also be stripped to remove light materials to meet product purity requirements. The overhead naphtha stream in line  134  may be condensed and separated in a receiver with a portion of the liquid being refluxed back to the fractionation column  130 . The net light naphtha stream in line  144  may be further processed before blending in a gasoline pool. The product fractionation column  130  may be operated with a bottoms temperature between about 260° C. (500° F.), and about 385° C. (725° F.), preferably at no more than about 350° C. (650° F.), and at an overhead pressure between about 7 kPa (gauge) (1 psig) and about 69 kPa (gauge) (10 psig). A portion of the unconverted oil stream in the bottoms line  142  may be reboiled and returned to the product fractionation column  130  instead of adding an inert stream such as steam to heat to the fractionation column  130 . 
     The LPG recovery section  150  may comprise a stripper scrubber column  160 , a sponge absorber column  170 , an optional deethanizer column  180  and a debutanizer column  190 . The net stripper off gas stream in the net stripper overhead line  124  rich in LPG gas may be routed to a stripper scrubber column  160  before it is transported to the sponge absorber column  170 . The sponge absorber column  170  may be in downstream communication with the overhead line  120  of the stripper column  90 . The stripper overhead liquid stream transported in a stripper receiver bottoms line  126  may be routed to an inlet of the deethanizer column  180 . In the stripper scrubber column  160 , the stripper off gas stream enters the stripper scrubber column at an inlet near a bottom and flows upwardly, while a lean amine stream in solvent line  162  enters the stripper scrubber column at an inlet near a top and flows downwardly. In an aspect, a deethanizer gas stream in line  186  may be scrubbed along with the stripper off gas stream. Hence, the deethanizer gas stream in line  186  may be added to the net stripper off gas stream in the net stripper overhead line  124  before entering the stripper scrubber column together in line  129 . Preferred lean amines include alkanolamines DEA, MEA, and MDEA. Other amines can be used in place of or in addition to the preferred amines. The lean amine contacts the stripper off gas stream and absorbs acid gas contaminants such as hydrogen sulfide and carbon dioxide. The resultant “sweetened” stripper off gas stream is taken out from an overhead outlet of the stripper scrubber column in a stripper scrubber overhead line  164 , and a rich amine is taken out from the bottoms at a bottom outlet of the stripper scrubber column in scrubber bottoms line  166 . The rich amine may undergo regeneration to remove the hydrogen sulfide for processing to generate elemental sulfur. The stripper scrubbing column  160  may be operated with a gas inlet temperature between about 30° C. (86° F.) and about 66° C. (150° F.) and an overhead pressure of about 0.35 MPa (gauge) (50 psig) to about 1.7 MPa (gauge) (250 psig). 
     The sweetened stripper off gas stream comprising LPG hydrocarbons and some naphtha may be transported to the sponge absorber column  170  to recover LPG and naphtha hydrocarbons. Recovery of LPG hydrocarbons is improved in the sponge absorber  170  because hydrogen removed in the medium separator  60  and the medium flash drum  66  is not present in the sweetened stripper off gas stream in stripper scrubber overhead line  164 , which would otherwise reduce LPG absorption into the sponge oil. 
     The multi-tray sponge absorber column  170  may include a first inlet  164   i  at a tray location near a bottom of the sponge absorber column  170 . The sponge absorber  170  receives the sweetened stripper off gas from the stripper scrubber overhead line  164  at the first inlet  164   i . The sponge absorber column  170  may include a second inlet  106   i  at a tray location near a middle of the sponge absorber column  170  which is at an elevation that is above the first inlet  164   i . The sponge absorber column  170  receives the compressed tail gas in the tail gas line  106  from the PSA unit  102  at the second inlet  106   i . The sponge absorber column  170  may be in downstream communication with the PSA unit  102 . A lean sponge oil stream is fed into the sponge absorber column  170  through a lean sponge oil line  172  which may be a naphtha hydrocarbon stream such as from a debutanizer bottoms stream or a stream from a naphtha splitter column (not shown) which would be in downstream communication with the debutanizer bottoms line  194  carrying the debutanizer bottom stream. In the sponge absorber  170 , the lean sponge oil and the sweetened stripper off gas stream are counter currently contacted. The sponge oil absorbs, extracts, and separates hydrocarbons from the stripper off gas stream. Moreover, in the sponge absorber  170 , the lean sponge oil and the tail gas stream are counter currently contacted, and the sponge oil absorbs, extracts, and separates hydrocarbons from the tail gas stream. The hydrocarbons absorbed by the sponge oil include a substantial amount of methane and ethane and most of the LPG, C 3  and C 4 , hydrocarbons and the C 5 , and C 6+  light naphtha hydrocarbons in the tail gas stream and/or the stripper off gas stream. The sponge absorber  170  operates at a temperature of about 34° C. (93° F.) to 60° C. (140° F.) and a pressure essentially the same as the overhead receiver  122  stripper column  90  less frictional losses. A sponge absorption off gas stream is withdrawn from a top of the sponge absorber column  170  at an overhead outlet through a sponge absorber overhead line  174 . A rich absorbent stream rich in LPG and naphtha hydrocarbons is withdrawn in a sponge absorber bottoms line  176  from a bottom of the sponge absorber column  170  at a bottoms outlet. 
     A portion of the sponge absorption off gas stream in the sponge absorber overhead line  174  may be transported to the hydrogen recovery unit  100  for hydrogen recovery. A sponge absorption off gas recovery stream in off gas recovery line  200  may be split from sponge absorption off gas purge line  202  and transported to the hydrogen recovery unit  100 . In an aspect, the sponge absorption off gas stream may be compressed in a sponge absorber compressor  204  up to about 1.4 MPa (gauge) (200 psig) to about 3.1 MPa (gauge) (450 psig). The compressed, sweetened sponge absorption off gas stream may be fed from the sponge absorption off gas compressor  204  in a compressor line  201  to a sponge knock out drum  203  to remove a compressed, condensed stream in a knockout bottoms line  207  for transport to the stripper column  90 . The sponge knock out drum  203  may be in downstream communication with the sponge absorption off gas compressor  204 . The bottoms line  207  may transport the compressed, condensed stream to the mixed line  74 . The compressed, condensed stream in the knock out bottoms line  207 , the cold flash liquid stream in cold flash bottoms line  54  and the medium flash liquid stream in medium flash bottoms line  70 , which may include the medium knockout liquid stream from the medium knockout bottoms line  73 , may mix together in the mixed line  74  and be transported to the cold stripper column  92 . The remaining compressed, sweetened stripper off gas may be transported from the sponge knock out drum  202  to the hydrogen recovery unit  100  in a sponge hydrogen recovery feed line  208  which may be an overhead line from the sponge knock out drum  203 . The sponge absorber compressor  204  may be in downstream communication with the sponge absorber overhead line  174 . The sponge absorption off gas in purge line  202  may be treated and routed to the refinery fuel gas system. 
     In the hydrogen recovery unit  100 , the sponge absorption off gas stream in the sponge hydrogen recovery feed line  208  may be fed to the PSA unit  102  in which impurities are adsorbed from hydrogen in a portion of the sponge absorption off gas of the sponge absorption off gas recovery stream to provide the tail gas stream in tail gas line  106  and the purified hydrogen stream in product line  104  with a greater hydrogen purity than in the sponge absorber off gas stream in the sponge absorber overhead line  174  and the sponge hydrogen recovery feed line  208 . The PSA unit  102  is in downstream communication with the sponge absorber overhead line  174 . Additionally, the PSA unit  102  may be in downstream communication with the sponge absorber compressor  204 . 
     In an aspect, the PSA unit  102  adsorbs impurities from hydrogen in the sponge hydrogen recovery feed line  208  and the scrubbed medium gaseous stream and/or the scrubbed cold flash gaseous stream in the separator scrubber overhead line  78 . In an aspect, the scrubbed medium gaseous stream and/or the scrubbed cold flash gaseous stream in the separator scrubber overhead line  78  and the sponge absorber off gas recovery stream in the sponge hydrogen recovery feed line  208  may be combined in a common off gas line  206  and fed to the PSA unit together. 
     The LPG-rich absorbent stream in the sponge absorber bottoms line  176  may be fractionated to provide a hydrocarbon stream. The absorbent stream may be routed to an optional deethanizer column  180  to fractionate the absorbent stream and separate a C 2−  stream from a C 3+  hydrocarbon stream. Additionally, the net stripper overhead liquid stream in the stripper receiver bottoms line  126  also has substantial LPG hydrocarbons. Hence, the stripper overhead liquid stream in the stripper receiver bottoms line  126  may also be transported to the deethanizer column  180  to fractionate a C 2−  stream from a C 3+  hydrocarbon stream. In an aspect, the stripper overhead liquid stream in the stripper receiver bottoms line  126  may be combined with the rich absorbent stream and transported together to the deethanizer column  180  in a combined line  178 . A deethanizer overhead stream may be partially condensed and separated in an overhead receiver into a deethanizer liquid stream for reflux to the column and a deethanizer gas stream in the deethanizer net overhead line  186 . The deethanizer gas stream may be recycled to the stripper scrubber column  160  perhaps after mixing with the stripper off gas stream in the stripper net overhead line  124  and then routed to the stripper scrubber column  160  via line  129 . An ethane stream can be condensed and recovered from a deethanizer overhead receiver if desired by using cryogenic equipment. A deethanized bottoms stream may be withdrawn from a bottom of the deethanizer column  180 . A portion of the deethanized bottoms stream may be reboiled and sent back to the deethanizer column while a net deethanized bottoms stream comprising C 3+  hydrocarbons and concentrated in LPG and naphtha hydrocarbons is withdrawn in net deethanizer bottoms line  188 . The net deethanizer bottoms stream may be transported to the debutanizer column  190 . The deethanizer column  180  may be operated with a bottoms temperature between about 160° C. (320° F.) and about 200° C. (392° F.) and an overhead pressure of about 1 MPa (gauge) (150 psig) to about 2 MPa (gauge) (300 psig). 
     The debutanizer column  190  may fractionate the net deethanized bottoms stream in the net deethanized bottoms line  188  into an overhead stream comprising LPG product and a debutanized bottoms stream comprising naphtha. The debutanizer overhead stream may be condensed and a condensed stream from a debutanizer overhead receiver bottoms may be split between a reflux stream refluxed to the debutanizer column and a LPG product stream that may be recovered in LPG product line  192  or further processed such as in a caustic treatment process to remove sulfur compounds from the LPG product stream. If the deethanizer column  180  is omitted, an uncondensed net gas stream can be taken from a debutanizer overhead receiver overhead and routed to the stripper scrubber column  160 . A debutanized bottoms stream may comprise a full range naphtha product stream. A portion of the debutanized bottoms stream may be reboiled and returned to the debutanizer column  190 . Another portion of the debutanized bottoms stream may be recycled as the lean sponge oil hydrocarbon stream in sponge oil line  172  or further fractionated in a naphtha splitter (not shown). A further portion of the debutanized bottoms stream may be recovered as a naphtha product stream in the debutanizer bottoms line  194  which may be further processed to improve its value or taken as a fuel product as is. The debutanizer column  190  may be operated with a bottoms temperature between about 180° C. (356° F.) and about 220° C. (430° F.) and an overhead pressure of about 0.8 MPa (gauge) (120 psig) to about 1.7 MPa (gauge) (250 psig). 
       FIG. 3  shows an embodiment of an apparatus and process  10 ′ which utilized dual sponge absorbers  170 ′ and  210  for recovering even further LPG and naphtha hydrocarbons. Elements in  FIG. 3  with the same configuration as in  FIG. 1  will have the same reference numeral as in  FIG. 1 . Elements in  FIG. 3  which have a different configuration than the corresponding element in  FIG. 1  will have the same reference numeral but designated with a prime symbol (′). The configuration and operation of the embodiment of  FIG. 3  is essentially the same as in  FIG. 1  with the noted exceptions. In the embodiment of  FIG. 3 , the medium pressure separator  60  and the medium flash drum  66  can be omitted. In such an embodiment, the hot flash bottoms line  48  would deliver the hot flash liquid stream to the hot stripper column  94 , the cold flash bottoms line  54  would deliver the cold flash liquid stream to the cold stripper column  92  and the cold flash overhead line  52  would deliver the cold flash gaseous stream to the hydrogen recovery unit  100 . The stripper column  90  may be operated as previously explained as a single column or as a cold stripper column  92  and a hot stripper column  94 . The LPG recovery section  150 ′ may comprise a stripper scrubber column  160 , a first sponge absorber column  170 ′, a second sponge absorber column  210 , an optional deethanizer column  180  and a debutanizer column  190 . The stripper off gas stream in the net stripper overhead line  124  rich in LPG and naphtha may be routed to the stripper scrubber column  160  before it is transported to the first sponge absorber column  170 ′. The first sponge absorber column  170 ′ may be in downstream communication with the overhead line  120  of the stripper  90 . The stripper overhead liquid stream transported in a stripper receiver bottoms line  126  may be routed to an inlet of the deethanizer column  180 . 
     In the stripper scrubber column  160 , the stripper off gas stream enters the stripper scrubber column at an inlet near a bottom and flows upwardly, while a lean amine stream in solvent line  162  enters the scrubber at an inlet near a top and flows downwardly. In an aspect, a gas stream in line  186  may be scrubbed along with the stripper off gas stream. Hence, the gas stream in line  186  may be added to the stripper off gas stream in line  124  before entering the stripper scrubber together in line  129 . Preferred lean amines include alkanolamines DEA, MEA, and MDEA. Other amines can be used in place of or in addition to the preferred amines. The lean amine contacts the stripper off gas stream and absorbs acid gas contaminants such as hydrogen sulfide and carbon dioxide. The resultant “sweetened” stripper off gas stream is taken out from an overhead outlet of the stripper scrubber column  160  in a stripper scrubber overhead line  164 , and a rich amine is taken out from the bottoms at a bottom outlet of the stripper scrubber column in scrubber bottoms line  166 . The rich amine may undergo regeneration to remove the hydrogen sulfide for processing to generate elemental sulfur. The sweetened stripper off gas stream comprising LPG hydrocarbons and some naphtha may be transported to the first sponge absorber  170 ′ to recover LPG and naphtha hydrocarbons. 
     The multi-tray first sponge absorber column  170 ′ may include a first inlet  164   i  at a tray location near a bottom of the first sponge absorber column  170 ′. The first sponge absorber column  170 ′ receives the sweetened stripper off gas from the stripper scrubber overhead line  164  at the first inlet  164   i . The first sponge absorber column  170 ′ may include a second inlet  106   i  at a tray location near a middle of the first sponge absorber column  170 ′ which is at an elevation that is above the first inlet  164   i . The first sponge absorber column  170 ′ may receive the compressed tail gas in the tail gas line  106  from the PSA unit  102  at the second inlet  106   i . The first sponge absorber column  170 ′ may be in downstream communication with the PSA unit  102 . A lean sponge oil stream is fed into the first sponge absorber column  170 ′ through a lean sponge oil line  172  which may be a naphtha hydrocarbon stream such as from a debutanizer bottoms stream or a stream from a naphtha splitter column (not shown) which would be in downstream communication with the debutanizer bottoms line  194  carrying the debutanizer bottom stream. In the first sponge absorber column  170 ′, the lean sponge oil stream and the sweetened stripper off gas stream are counter currently contacted with each other. The sponge oil absorbs, extracts, and separates hydrocarbons from the stripper off gas stream. Moreover, in the sponge absorber  170 ′, the lean sponge oil and the tail gas stream may be counter currently contacted, and the sponge oil absorbs, extracts, and separates hydrocarbons from the tail gas stream. The hydrocarbons absorbed by the sponge oil include a substantial amount of methane and ethane and most of the LPG, C 3  and C 4 , hydrocarbons and the C 5 , and C 6+  light naphtha hydrocarbons in the tail gas stream and/or the stripper off gas stream. The first sponge absorber  170 ′ operates at a temperature of about 34° C. (93° F.) to 60° C. (140° F.) and a pressure essentially the same as the stripper column receiver  122  less frictional losses. A first sponge absorption stream is withdrawn from a top of the first sponge absorber column  170 ′ at an overhead outlet through a first sponge absorber overhead line  174 ′. A first absorbent stream rich in LPG hydrocarbons is withdrawn in a first sponge absorber bottoms line  176 ′ from a bottom of the sponge absorber column  170 ′ at a bottoms outlet. 
     The LPG-rich absorbent stream in the first sponge absorber bottoms line  176 ′ may be fractionated to provide a naphtha hydrocarbon stream and an LPG product stream. The absorbent stream is routed to an optional deethanizer column  180 , which is a fractionation column, to fractionate the absorbent stream and separate a C 2−  stream from a C 3+  hydrocarbon stream. In an aspect, the first sponge absorber bottoms line bypasses the second sponge absorber  210  and transports the first absorbent stream to an optional deethanizer column  180  which is in direct, downstream communication with the first sponge absorber bottoms line  176 ′. The deethanizer fractionation column  180  may be in downstream communication with the first sponge absorber column  170 ′. The stripper overhead liquid stream in the stripper receiver bottoms line  126  may also be transported to an optional deethanizer column  180  to fractionate a C 2−  stream from a C 3+  hydrocarbon stream. In an aspect, the stripper overhead liquid stream in the stripper receiver bottoms line  126  may be combined with the first absorbent stream and transported together to the deethanizer column  180  in a combined line  178  while bypassing the second sponge absorber column  210 . A deethanizer overhead stream may be partially condensed and separated in an overhead receiver into a deethanizer liquid stream for reflux to the column and the deethanizer gas stream in the deethanizer net overhead line  186 . The deethanizer gas stream may be recycled to the stripper scrubber column  160  perhaps after mixing with the stripper off gas stream in net overhead line  124  and then routed to the stripper scrubber via line  129 . If desired, an ethane stream can be condensed and recovered by using cryogenic equipment. A deethanized bottoms stream may be withdrawn from a bottom of the deethanizer column  180 . A portion of the deethanized bottoms stream may be reboiled and sent back to the deethanizer column while a net deethanized bottoms stream comprising C 3+  hydrocarbons and concentrated in LPG hydrocarbons is withdrawn in net deethanizer bottoms line  188 . The net deethanized bottoms stream may be transported to the debutanizer column  190 . 
     The debutanizer column  190  may fractionate the net deethanized bottoms stream in the net deethanized bottoms line  188  into an overhead stream comprising LPG product and a debutanized hydrocarbon stream comprising naphtha in the debutanizer bottoms line  194 . The debutanizer overhead stream may be totally condensed and a condensed stream from a debutanizer overhead receiver bottoms may be may be split between a reflux stream refluxed to the debutanizer column and a LPG product stream that may be recovered in LPG product line  192  or further processed such as in a caustic treatment process to remove sulfur compounds from the LPG product stream. If the deethanizer column  180  is omitted, an uncondensed net gas stream can be taken from a debutanizer overhead receiver overhead and routed to the stripper scrubber column  160 . The debutanized bottoms stream comprises a light naphtha product stream. A portion of the debutanized hydrocarbon stream may be reboiled and returned to the debutanizer column  190 . Another portion of the debutanized bottoms stream may be recycled as the lean sponge oil hydrocarbon stream in sponge oil line  172  to the first sponge absorber column  170 ′ or further fractionated in a naphtha splitter (not shown). The first absorber column  170 ′ may be in downstream communication with the deethanizer column  180  and the debutanizer column  190 , which are fractionation columns. A further portion of the debutanized bottoms stream may be recovered as a light naphtha product stream in the debutanizer bottoms line  194  which may be further processed to improve its value or taken as a fuel product as is. 
     We have found that the first sponge absorption stream in the first sponge absorber overhead line  174 ′ still has LPG and naphtha that can be recovered, so it may be transported to the second sponge absorber column  210  to absorb additional LPG and naphtha from the gaseous first sponge absorption stream. The multi-tray second sponge absorber column  210  may include an inlet at a tray location near a bottom of the second sponge absorber column  210  to receive first sponge absorption stream from the first sponge absorber column  170 ′. The second sponge absorber column  210  may be in downstream communication with the first sponge absorber overhead line  174 ′ and an overhead line  120  of the stripper column  90 . A second lean sponge oil stream is fed into the second sponge absorber column  210  through a second lean sponge oil line  212 . 
     In order to improve the recovery of naphtha, a portion of the cold stripped stream in the cold stripper bottoms line  128 ′ may be split into a cold stripper fractionation stream in cold fractionation feed line  214  and the second lean sponge oil stream in the second lean sponge oil line  212  and transported to the second sponge absorber column  210 . The second sponge absorber column  210  is in downstream communication with the cold stripper bottoms line  128 ′. In the second sponge absorber column  210 , the second lean sponge oil stream and the first sponge absorption stream are counter currently contacted with each other. The second sponge oil stream comprising a portion of the cold stripped stream absorbs, extracts, and separates hydrocarbons from the first sponge absorption stream. The hydrocarbons absorbed by the sponge oil include a substantial amount of methane and ethane and most of the LPG, C 3  and C 4 , hydrocarbons and light naphtha range hydrocarbons, C 5  and C 6+ , in the first sponge absorption stream absorbed from the tail gas stream and/or the stripper off gas stream. The second sponge absorber  210  operates at a temperature of about 34° C. (93° F.) to 60° C. (140° F.) and a pressure essentially the same as the stripper column receiver  122  less frictional losses. A second sponge absorption off gas stream is withdrawn from a top of the second sponge absorber column  210  at an overhead outlet through a second sponge absorber overhead line  216 . A second rich sponge oil stream rich in LPG and naphtha hydrocarbons is withdrawn in a second sponge absorber bottoms line  218  from a bottom of the sponge absorber column  210  at a bottoms outlet. 
     A portion of the second sponge absorption off gas stream in the second sponge absorber overhead line  216  may be transported to the hydrogen recovery unit  100  for hydrogen recovery. A sponge absorption off gas recovery stream in off gas recovery line  200  may be split from sponge absorption off gas purge line  202  and transported to the hydrogen recovery unit  100  as described with respect to  FIG. 1 . In an aspect, the sponge absorption off gas stream in the second sponge absorber overhead line  216  may be compressed in a sponge absorber compressor  204  up to about 1.4 MPa (gauge) (200 psig) to about 3.1 MPa (gauge) (450 psig) and transported to the hydrogen recovery unit  100  after knock out of condensed materials. The sponge absorber compressor may be in downstream communication with the sponge absorber overhead line  174 ′ and the second sponge absorber overhead line  216 . 
     The LPG and naphtha-rich second absorbent stream in the second sponge absorber bottoms line  218  may be transported to the stripper column  90  to have gases stripped from the second rich absorbent stream. In an aspect, the second rich absorbent stream may be stripped in the cold stripper  92 . In a further aspect, the second absorbent stream in the second absorbent bottoms line  218  may be combined with a medium flash liquid stream in a medium flash bottoms line  70  and transported to the cold stripper column  92 . Specifically, the medium flash liquid stream and the second rich absorbent stream may be mixed with the cold flash liquid stream in the cold flash bottoms line  54  and transported to the cold stripper column in the mixed line  74 . Alternatively, if no medium pressure separator  60  or medium pressure flash drum  66  are used, the second absorbent stream in sponge absorber bottoms line may be transported directly to the cold stripper column or mixed with the cold flash liquid stream in the cold separator bottoms line  54  and transported to the cold stripper in the mixed line  74 . Other aspects of  FIG. 3  are as described for  FIG. 1 . In the cold stripper column  92 , the hydrocarbon rich second rich absorbent stream and the medium flash liquid stream are stripped of gases together. 
     EXAMPLES 
     We calculated the improvement in hydrogen recovery obtained using a medium separator and medium flash drum to recover hydrogen from a hot flash liquid stream in a PSA unit. Hydrogen recovery improved by 8 wt % compared to allowing the hot flash liquid flow directly to a stripper. 
     SPECIFIC EMBODIMENTS 
     While the following is described in conjunction with specific embodiments, it will be understood that this description is intended to illustrate and not limit the scope of the preceding description and the appended claims. 
     A first embodiment of the invention is a hydroprocessing process comprising hydroprocessing a hydrocarbon feed stream in a hydroprocessing reactor to provide a hydroprocessing effluent stream; separating the hydroprocessing effluent stream in a hot separator to provide a hot separator gaseous stream and a hot separator liquid stream; separating the hot separator liquid stream in a hot flash drum to provide a hot flash gaseous stream and a hot flash liquid stream; and recovering hydrogen from the hot flash liquid stream. An embodiment of the invention is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph wherein recovering hydrogen from the hot flash liquid stream further includes separating the hot flash liquid stream to provide a medium gaseous stream and a medium liquid stream; and recovering hydrogen from the medium gaseous stream. An embodiment of the invention is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph wherein separating the hot flash liquid stream further comprises stripping the medium liquid stream in a stripper. An embodiment of the invention is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph further comprising flashing the medium gaseous stream to provide a medium flash gaseous stream and a medium flash liquid stream. An embodiment of the invention is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph further comprising cooling the medium gaseous stream. An embodiment of the invention is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph further comprising compressing the medium flash gaseous stream. An embodiment of the invention is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph further comprising removing acid gases from the medium gaseous stream. An embodiment of the invention is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph further comprising adsorbing impurities from hydrogen in the medium gaseous stream. An embodiment of the invention is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph further comprising stripping the medium liquid stream in a stripper. An embodiment of the invention is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph further comprising separating the hot separator gaseous stream in a cold separator to provide a cold separator gaseous stream and a cold separator liquid stream and recovering hydrogen from the cold separator gaseous stream. An embodiment of the invention is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph further comprising separating the cold separator liquid stream in a cold flash drum to provide a cold flash gaseous stream and a cold flash liquid stream. An embodiment of the invention is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph further comprising recovering hydrogen from the cold flash gaseous stream and stripping the cold flash liquid stream. An embodiment of the invention is one, any or all of prior embodiments in this paragraph up through the first embodiment in this paragraph wherein recovering hydrogen from the hot flash liquid stream further includes separating the hot flash liquid stream to provide a medium gaseous stream and a medium liquid stream, flashing the medium gaseous stream to provide a medium flash gaseous stream and a medium flash liquid stream, mixing the cold flash gaseous stream with the medium flash gaseous stream and recovering hydrogen from the medium flash gaseous stream and the cold flash gaseous stream and mixing the cold flash liquid stream with the medium flash liquid stream. 
     A second embodiment of the invention is a hydroprocessing process comprising hydroprocessing a hydrocarbon feed in a hydroprocessing reactor to provide a hydroprocessing effluent stream; separating the hydroprocessing effluent stream in a hot separator to provide a hot separator gaseous stream and a hot separator liquid stream; separating the hot separator liquid stream in a hot flash drum to provide a hot flash gaseous stream and a hot flash liquid stream; separating the hot flash liquid stream to provide a medium gaseous stream and a medium liquid stream; and recovering hydrogen from the medium gaseous stream. An embodiment of the invention is one, any or all of prior embodiments in this paragraph up through the second embodiment in this paragraph further comprising flashing the medium gaseous stream in a medium pressure flash drum to provide a medium flash gaseous stream and a medium flash liquid stream and recovering hydrogen from the medium flash gaseous stream. An embodiment of the invention is one, any or all of prior embodiments in this paragraph up through the second embodiment in this paragraph further comprising stripping the medium liquid stream in a stripper. An embodiment of the invention is one, any or all of prior embodiments in this paragraph up through the second embodiment in this paragraph further comprising stripping the medium flash liquid stream in a stripper. 
     A third embodiment of the invention is a hydroprocessing apparatus comprising a hydroprocessing reactor; a hot separator in communication with the hydroprocessing reactor; a hot flash drum in communication with a bottoms line of the hot separator; a hydrogen recovery unit in communication with a bottoms line of the hot flash drum. An embodiment of the invention is one, any or all of prior embodiments in this paragraph up through the third embodiment in this paragraph further comprising a medium pressure separator in communication with the bottoms line of the hot flash drum and the hydrogen recovery unit is in communication with an overhead line of the medium pressure separator. An embodiment of the invention is one, any or all of prior embodiments in this paragraph up through the third embodiment in this paragraph further comprising a medium flash drum in communication with an overhead line of the medium pressure separator and the hydrogen recovery unit is in communication with an overhead line of the medium flash drum. 
     Without further elaboration, it is believed that using the preceding description that one skilled in the art can utilize the present invention to its fullest extent and easily ascertain the essential characteristics of this invention, without departing from the spirit and scope thereof, to make various changes and modifications of the invention and to adapt it to various usages and conditions. The preceding preferred specific embodiments are, therefore, to be construed as merely illustrative, and not limiting the remainder of the disclosure in any way whatsoever, and that it is intended to cover various modifications and equivalent arrangements included within the scope of the appended claims. 
     In the foregoing, all temperatures are set forth in degrees Celsius and, all parts and percentages are by weight, unless otherwise indicated.