fig1 illustrates feed entering the process through stream 1 and being combined with hydrogen in stream 28 to form stream 2 . hydrogen in stream 28 is prepared by compression of hydrogen in makeup compressor 85 . hydrogen enters compressor 85 through stream 27 . the invention includes an option to compress a stream 30 of recycle gas in the last stage of compressor 85 to meet the gas to oil ratio in reactor 10 , when required . stream 2 is heated , as depicted by exchanger 31 , prior to entering the first stage hydroprocessing unit , vessel 10 . vessel 10 is preferably operated as a hydrotreater . the feed flows downward through one or more beds of catalyst . streams 3 , 4 , and 5 depict interbed hydrogen quench . hydrotreated effluent exits vessel 10 through stream 32 and is reduced in pressure ( valve 33 ) to that required for hydrocracking in a clean environment . the effluent is heated in furnace 34 to approximately 825 ° f . in order to disengage the maximum material in very hot high pressure separator 20 . this separator functions as a simple flash drum , separating diesel and lighter fractions from heavier materials without the use of hydrogen stripping . hydrogen stripping is relatively ineffective at hydrocracking pressures . stream 11 , containing diesel and lighter materials , exits vessel 20 overhead . external feeds in the middle distillate boiling range , as well as fractionation recycle , are represented by stream 9 and are combined with stream 11 . stream 11 is heated in exchanger 35 and may be combined with hydrogen in stream 25 prior to entering a distillate upgrader , vessel 30 , in the case of co - current flow . flow in vessel 30 may be co - current or countercurrent . countercurrent flow may be preferred if aromatics saturation is desired . the amount of aromatics permitted in the ultra - low sulfur diesel being manufactured ( ulsd ) may affect whether co - current or counter - current flow is used . in the case of countercurrent flow , hydrogen is added below the catalyst beds and is directed upward . the catalyst in the bed or beds of vessel 30 is preferably hydrotreating catalyst , but hydrocracking catalyst may be used if fractionation recycle is being treated . the bottoms effluent of vessel 30 exits through stream 15 . material from stream 15 may be passed to stream 12 as feed to the hydrocracker , vessel 40 , when necessary . the dotted line depicts this . the upgraded diesel effluent in stream 15 is reduced in pressure ( valve 36 ), cooled ( exchanger 37 ), combined with the effluent stream ( stream 14 ) from vessel 40 ( in which second stage hydrocracking preferably occurs ) to become stream 16 . stream 16 is passed to the hot high pressure separator 70 , where it is separated into an overhead stream 18 and a bottoms stream 17 . bottoms stream 17 is sent to fractionation . overhead stream 18 is cooled prior to entering cold high pressure separator 50 by passage through exchangers 43 and 44 , as well as by water injection through stream 19 . sour water exits cold high pressure separator through stream 29 . stream 71 goes to fractionation . it may be reduced in pressure using valve 72 . overhead gaseous material in stream 21 enters amine absorber , vessel 60 at the bottom and flows upward , as lean amine moves downward , absorbing hydrogen sulfide . rich amine exits vessel 60 through stream 22 . stream 23 , comprising primarily hydrogen , exits overhead through stream 23 . stream 23 is compressed in compressor 75 , becoming stream 24 . stream 24 is divided into streams 25 and 26 . stream 26 is heated in exchanger 42 before combining with stream 12 to form stream 13 . the bottoms effluent of vessel 20 exits through stream 12 . valve 38 is a level control valve . stream 12 may be combined with material in stream 15 , along with hydrogen in stream 26 then is heated in exchanger 39 . streams 12 and 15 may be combined when naphtha or jet fuel is the preferred product . recycle stream 31 may be added to stream 15 when very high conversion levels are required . stream 13 exits exchanger 39 and enters vessel 40 . second stage hydrocracking preferably occurs in vessel 40 , which contains one or more beds of hydrocracking catalyst . effluent in stream 14 is cooled in exchanger 41 before being combined with stream 16 . a wide variety of hydrocarbon feeds may be used in the instant invention . typical feedstocks include any heavy or synthetic oil fraction or process stream having a boiling point above 392 ° f . ( 200 ° c .). such feedstocks include vacuum gas oils ( vgo ), heavy coker gas oil ( hcgo ), heavy atmospheric gas oil ( ago ), light coker gas oil ( lcgo ), visbreaker gas oil ( vbgo ), demetallized oils ( dmo ), vacuum residua , atmospheric residua , deasphalted oil , fischer - tropsch streams , light cycle oil and other fcc product streams . the process can be used over a broad range of applications as shown in the table 1 . the process of this invention is especially useful in the production of middle distillate fractions boiling in the range of about 250 ° f . to 700 ° f . ( 121 ° c . to 371 ° c .). a middle distillate fraction is defined as having an approximate boiling range from about 250 ° f . to 700 ° f . at least 75 vol . %, preferably 85 vol . % of the components of the middle distillate has a normal boiling point of greater than 250 ° f . at least about 75 vol . %, preferably 85 vol . % of the components of the middle distillate has a normal boiling point of less than 700 ° f . the term “ middle distillate ” includes the diesel , jet fuel and kerosene boiling range fractions . the kerosene or jet fuel boiling point range refers to the range between 280 ° f . and 525 ° f . ( 38 ° c . to 274 ° c .). the term “ diesel boiling range ” refers to hydrocarbons boiling in the range from 250 ° f . to 700 ° f . ( 121 ° c . to 371 ° c .). gas streams or naphtha may also be produced in the process of this invention . gas streams or naphtha normally boils in the range below 400 ° f . ( 204 ° c . ), or from c 5 to 400 ° f . ( 204 ° c .). boiling ranges of various product fractions recovered in any particular refinery will vary with such factors as the characteristics of the crude oil source , local refinery markets and product prices . a hydroprocessing condition is a general term which refers primarily in this application to hydrocracking or hydrotreating . hydrotreating conditions include a reaction temperature between 400 ° f . to 900 ° f . ( 204 ° c . to 482 ° c . ), preferably 650 ° f . to 850 ° f . ( 343 ° c . to 464 ° c . ); a pressure between 500 to 5000 psig ( pounds per square inch gauge ) ( 3 . 5 to 34 . 6 mpa ), preferably 1000 to 3000 psig ( 7 . 0 to 20 . 8 mpa ): a feed rate ( lhsv ) of 0 . 5 to 20 hr - 1 ( v / v ); and overall hydrogen consumption 300 to 2000 scf per barrel of liquid hydrocarbon feed ( 63 . 4 to 356 m 3 / m 3 feed . the second stage hydrotreating reactor is operating at a lower pressure than the first stage reactor , the vgo hydrotreater or moderate severity hydrocracker . typical hydrocracking conditions ( which may be found in stage 1 or stage 2 ) include a reaction temperature of from 400 ° f . to 950 ° f . ( 204 ° c . to 510 ° c .) preferably 650 ° f . to 850 ° f . ( 343 ° c . to 454 ° c .). reaction pressure ranges from 500 to 5000 psig ( 3 . 5 to 4 . 5 mpa ), preferably 1500 to 3500 psig ( 10 . 4 to 24 . 2 mpa ). lhsv ranges from 0 . 1 to 15 hr - 1 ( v / v ), preferably 0 . 25 to 2 . 5 hr hydrogen consumption ranges from 500 to 2500 scf per barrel of liquid hydrocarbon feed ( 89 . 1 to 445 m 3 h 2 / m 3 feed ). a hydroprocessing zone may contain only one catalyst , or several catalysts in combination . the hydrocracking catalyst generally comprises a cracking component , a hydrogenation component and a binder . such catalysts are well known in the art . the cracking component may include an amorphous silica / alumina phase and / or a zeolite , such as a y - type or usy zeolite . catalysts having high cracking activity often employ rex , rey and usy zeolites . the binder is generally silica or alumina . the hydrogenation component will be a group vi , group vii , or group viii metal or oxides or sulfides thereof , preferably one or more of molybdenum , tungsten . cobalt , or nickel , or the sulfides or oxides thereof . if present in the catalyst , these hydrogenation components generally make up from about 5 % to about 40 % by weight of the catalyst . alternatively , platinum group metals , especially platinum end / or palladium , may be present as the hydrogenation component , either alone or in combination with the base metal hydrogenation components molybdenum , tungsten , cobalt , or nickel . if present , the platinum group metals will generally make up from about 0 . 1 % to about 2 % by weight of the catalyst . if aromatic saturation is particularly desired , a preferred catalyst has a crystalline molecular sieve material component and a group viii noble metal component . the crystalline molecular sieve material component is a large pore faujasite structure having an alpha acidity of less than 1 , preferably less than 0 . 3 . zeolite usy is the preferred crystalline molecular sieve material component . hydrotreating catalyst , if used , will typically be a composite of a group vi metal or compound thereof , and a group viii metal or compound thereof supported on a porous refractory base such as alumina . examples of hydrotreating catalysts are alumina supported cobalt - molybdenum , nickel sulfide , nickel - tungsten , cobalt - tungsten and nickel - molybdenum . typically , such hydrotreating catalysts are presulfided . fig2 demonstrates that conversion in the instant invention is greater at lower temperatures , as opposed to the conventional hydrocracking configuration . conversion improves at higher gas to oil ratios . fig3 demonstrates that yield to conversion ratios are comparable in both the conventional configuration as well as the configuration of the instant invention .