WO2001061097A1 - A wet processing textile machine - Google Patents

Abstract

A wet processing textile machine for wet treatment of yarns/fabrics in rope form comprises an autoclave (1) substantially of annular duct form placed horizontally and provided with a pair of inlet-cum-outlets (11, 12) at its top side for feeding the fabric to be treated and for taking out the treated fabric. A basket (2) substantially of annular duct form made of spaced apart rods or perforated sheet, open at top end and adopted to be rotatably placed inside the said autoclave (1). A processing fluid circulating means (13, 19, 29) provided for circulating the processing fluid, a pair of nozzles (22, 23) provided, one each, at the said inlet-cum-outlets (11, 12) and connected to the discharge side of the said fluid circulating means (13) through flow control means/valves (24, 25), a pair of sensors (26, 27) provided, one each, at the said inlet-cum-outlets (11, 12) for sensing the fabric movement and giving signals to the electronic control devices (36, 37) for regulating the fluid through the nozzles (22, 23), one by one (only one at a time), and controlling the direction of the rotation of the said basket (2) in a synchronized manner, after each rotation, thereby intermittently reversing the direction of fabric movement.

Description

A WET PROCESSING TEXTILE MACHINE

This invention relates to a process and a system for the production of middle distillate products comprising hydrocarbons having carbon atoms in the range of C₈ to C; in high \ ield. fiom heavier petroleum fractions through multistage catalytic ciackmg of varying seventy levels with solid acidic catalyst without using external hydrogen

Background

Conventionally, middle distillate range products e.g. Heavy Naphtha. Kerosene. Jet fuel Diesel oil and Light Cycle Oil (LCO) are produced m petroleum refinenes by atmosphenc/vacuum distillation of petroleum crude and also by the secondary processmg of vacuum gas oil and residues or mixtures thereof Most commonly practiced commercial secondary processes are Fluid Catalytic Cracking (FCC) and Hydrocrackmg. Hydrocrackmg employs porous acidic catalysts similai to those used m catalytic cracking but associated with hydrogenation components such as metals of Groups VI and VII of the Penodic Table to produce good quality of middle distillate products m the boiling range of Cg - C:4 hydrocarbons An excess of hydrogen is supplied to the hydrocrackmg ieactor under very high pressure (150-200 atm.) and at a relatively lower temperature (375-425°G m fixed bed reactors with two phase flow Due to severe hydrogenation. all hydrocarbon products from Hydrocracker are highly saturated with low sulfur and aromaticity The yield of middle distillate hydrocarbons ( 126-39 I boiling range) in hydrocrackmg is typically very high up to 65 - 80 wt° o of feed FCC process, on the other hand, is employed for essentially producmg high octane Gasoline and LPG In countnes, where demand of middle distillate pioduct is highei. Heavy Ciacked Naphtha (HCN C₈ - C₁ hydrocarbons) and Light Cycle Oil (LCO d , - C_:4 hydrocarbons) pioduction are maximized by manipulating operating vanables so as to vary the reaction and regenerator severity levels S Patent Nos 3.894,93 1 and 3.894 933 address such operations TypicalK diesel yield in FCC is maximized by maintaining a lower ieaction and iegeneiation seventy (I e , lower regenerator and reactor top temperatuie) and recy clmg of unconverted residual products Catalyst with lower zeohte/matnx ratio and MAT (Micro Activity Test) acnvity of 60-70 is normally prefened By proper selection of FCC vanables and innovations involving catalyst type and recycle of Heavy Cycle Oil and residual Slurry oil. distillate yield can be mcreased by considerable amount at the expense of Gasolme yield As the FCC unit operation is shifted from gasolme mode to middle distillate maximization mode, the LCO cetane number mcreases and thus could be more useful for blending to diesel pool

Howevei. while running at low seventy operations, for maximizing diesel yield, the unconverted bottom yield also mcreases to a significant extent and sometimes may even exceed 20 wt% of fresh feed as against 5-6 wt% for usual gasolme mode opeiation The other drawback of low seventy operation is the high amount of recycle oil bemg used m the nser bottom with fresh feed for furthei ciacking Firstly this reduces the throughput of riser reactor and secondly, with single πsei and pioduct fractionatoi . the recycle is nonselective This lesults into lecychng of un-crackable, aromatic components into the nser and thereby increases Coke and Gas without appreciably increasing the conversion level Consequently Diesel yield from FCC with the conventional cracking catalyst could be maximized upto 40-45 wt% m spite of running at low reaction seventy (495°C riser temperature) and fanly higher recycle ratio (30% of fresh feed)

Besides the operation of conventional FCC in middle distillate maximizanon mode, there are seveial other processes aimmg for improvement in middle distillate yield L S Pat No 5.098.554 discloses a process of fluid catalytic cracking with multiple feed injection pomts where fresh feed is charged to uppei injection points and unconverted slurry oil is recycled to a location below the fresh feed nozzles Essentially, the process conditions are similar to that of gasoline mode FCC operation ( e g . 527°C nser top temperature) which favois gasolme production By adopting split feed injection, middle distillate yield is marginally increased at the expense of Gasolme yield

L S Pat No 4,481, 104 descnbes about an ultra-stable Y-zeohte of high framewoik silica to alumina ratio having low acidity, large pores, use of which in catalytic cracking of gas oil. enhances distillate yield with production of low Coke and Dry gas It may be noted that yield of 420 - 650°F fraction is maximize about 28 wt% of feed and as 650°F- conversion mcreases beyond 67 wt%. the yield of 420-650°F fraction further reduces Therefore, as discussed earlier yield of the distillate is relatively more only at the higher yield of unconverted fraction

Yet another process in U S Pat No 4 606,810 discloses a scheme of two nser crackmg for improving total gasolme plus distillate yield Here, the feed is first ciacked in the first πsei with spent catalyst from the second nsei and the unconverted part is further cracked m a second nser in presence of regenerated catalyst The basic operation is of high seventy producmg maximum amount of Gasolme and the vield of LFO is around 15 - 20 wt% of feed It mav also be noted that while increase in Gasolme yield is in the range of 7.5 - 8 0 wt%. increase m LFO yield is merely in the range of 1 5 - 3 0 wt% on fresh feed basis

Two stage processing of hydrocarbon feedstock has been employed by different lesearcheis in the field of catalytic cracking Several processes have been developed m which first stage processing removes metals and Conradson cai'bon iesidue ( CCR) impurities from feed using a low activity cheap contact matenal with abundant surface aiea The demetalhzed feed is then processed in a more conventional second stage reactor under high seventy to maximize the conveision and gasoline production. U S Pat. No 4.436,613 descnbes such a process of two stage catalytic cracking using two different types of catalyst. In the first stage, the CCR matenals and metals are separated from the rest of the feedstock along with mild cracking over a relatively lower active catalyst. The residual un-cracked product of the first stage is then contacted with a high active catalyst under higher reaction seventy for gasolme maximization. It may be noted that in this process, two dedicated shippers and regenerators are used to avoid the mixing of two different types of catalysts

Dual πsei high seventy catalytic cracking process descnbed in U S Pat No 3.928.1 2 utilizes a mixture of large pore REY zeolite catalyst and a shape selective zeolite catalyst where gas oil is cracked in the first nser in the presence of the aforesaid catalyst mixture. The Heavy Naphtha product from the first nser and/or virgin straight run Naphtha are cracked m the second riser m the presence of catalyst mixture to produce high octane Gasolme together with C₃ and C₄ olefins S Pat No 4.830.728 discloses a process for upgrading straight run Naphtha, catalytically cracked Naphtha and mixtures thereof m a multiple fluid catalytic cracking operation utilizing mixture of amorphous crackmg catalyst and/or large pore Y-zeohte based catalyst and shape selective ZSM-5 to produce high octane gasolme U S Pat No 5,401.387 descnbes a process of multistage catalytic cracking where the first stage cracks a first feed over a shape selective zeolite to produce lightei products nch in iso-compounds which may be used for making ethers A. second feed which ma\ include 700°F- liquid from fust stage is cracked m the second stage Anofhei process as descnbed in U S Pat No 5,824,208. discloses a scheme m which hydrocarbon is initially contacted with cracking catalyst forming a first cracked product which after recovering of the product having boiling point of moie than 430°F. is subjected to cracking in a second nser The basic objective of this invention is to maximize the yield of light olefrns and minimize the formation of aromatic compounds by avoidmg undesnable hydrogen transfer reactions

So far, most of the pnor art methods have concentrated on multiple nser catalytic crackmg for maximization of gasolme yield and its octane numbers, mcreased yield of lso-olefm for production of ethers, mcreased yield of light olefrns. etc From the pnor art information and also from our expenence of operating low seventy FCC units, it is quite clear that maximizing middle distillate yield in FCC (without usmg external hydrogen) is not achieved beyond a level of 40-45 wt% of fresh feed Further, persons involved m fluid crackmg would be aware that middle distillate bemg an intermediate product in the complex catalytic ciacking reactions, its maximization is very difficult because hen the severity is increased, it is re-cracked to lighter hydrocarbons

Objects

Accordingly, the mam object of the present invention aims to propose a novel catalytic cracking piocess foi producmg middle distillate products in very high yield (about 50-65 wt%) Anothei object is to provide a multiple nser system that enables the productton of middle distillate products mcludmg Heavy Naphtha and Light Cycle Oil in high yield

Yet anothei object of the invention is to provide a multiple nser system to pioduce higher yield of Heavy Naphtha and Light Cycle Oil as compared to the prior art processes employing catalytic cracking of petroleum feedstock wnhout any use of external supply of hydrogen

Λ furthei objective of the piocess is to mimmize the yield of unwanted dry gas and coke and also the \ leld of unconverted bottom products, at the same time, improving the cetane quality of the middle distillate product

Summary

According to the present invention, there is provided a novel process for catalytic crackmg of vanous petroleum based heavy feed stocks m the presence of solid zeolite catalyst and high pore size acidic components for selective bottom crackmg and mixtures thereof, in a multiple nser type system wherem continuously circulating fluidized bed reactors aie operated at different sevennes to produce middle distillate products m high yield, in the range of 50-65 wt^0/o of fresh feed

The invention also provides an improved system foi catalytic cracking of heavy feed stock to obtain middle distillate products in high yield, employing the process herein descnbed

Detailed Description

The mvention i elates to a multi stage selective catalytic crackmg process for producmg high yield of middle distillate products having carbon atoms in the range of about C₈ to C₂₄, from heavy hydrocarbon feedstock, in the absence of added hydrogen, said process compnsmg the steps of

l) contacting preheated feed with a mixed catalyst in a first nser reactor under catalytic crackmg conditions mcludmg catalyst to oil ratio of 2 to 8. WHSV of 150-350 hr ^"'. contact penod of about 1 to 8 seconds and temperature m the range of about 400°C to 500°C to obtain first cracked hydrocarbon products. n) separating the first cracked hydrocarbon products from the first riser leactor into a first fraction compnsmg hydrocarbons with boiling points less than or equal to 370°C and a second fraction compnsmg unconverted hydrocarbons with boiling points greater than. or equal to 370°C, m) crackmg the unconverted second fraction from the first nser reactor compnsmg hydrocarbons having boiling pomts greater than or equal to 370°C, m the presence of regenerated catalyst, m a second nser reactor operating under catalytic crackmg conditions mcludmg WHSV of 75-275 hr^"1, catalyst to oil ratio of 4-12 and riser top temperature of 425 - 525°C to obtam second cracked hydrocarbon products, iv) separating the catalytically cracked products from the second nser ieactor alongwith cracked products compnsmg hydrocarbons having boiling points less than equal to 370°C. from the first nser reactor in a mam fractionating column to yield cracked products compnsmg dry gas. LPG, gasolme. middle distillates, heavy cycle oil and slurry oil. v) recycling the entire heavy cycle oil compnsmg hydrocarbons having boiling pomts m the range of 370°C to 450°C and full or part of the slurry oil having boiling points greater than or equal to 450°C mto the second nser reactor at a vertically displaced position lower than the position of introduction of the mam feed compnsmg bottom unconverted hydrocarbon fraction having boiling points gi eater than or equal to 370°C from the first πsei ieactoi to obtain middle distillate products compnsmg hydiocarbons havmg carbon atoms in the range of C₈ - C₂₄ ranging fiom about 50 to 65 wt of the feed stock iv ) Optionally recycling the fraction of unconverted hydrocarbons with boiling pomts greater than or equal to 370°C, obtained in step (v) m nser reactors by repeating steps (in) to (iv) to obtain substantially pure middle distillate products

In an embodiment, the feed stock is selected from petroleum based heavy feed stock, such as vacuum gas oil (VGO), visbreaker / coker heavy gas oil, coker fuel oil. hydrocracker bottom, etc

In another embodiment, mixed catalyst is obtamed from an intermediate vessel used foi mixing the spent catalyst from the common stnpper or preferably first stnppei with the regenerated catalyst from the common regenerator and charging the mixed catalyst with coke content m the range of about 0 2 to 0 8 wt% to the bottom of the first riser at a temperature of 450 - 575°C

In anothei embodiment, the exit hydrocarbon vapors from the first and second users aie quickly sepaiated from respective bpent catalysts using lespective cy clones and/oi othei conventional separating devices to minimize the ovei ciacking of middle distillate range products mto undesirable lighter hvdiocaibons In yet anothei embodiment, the spent catalysts from the first and second nser reactors are passed through respective dedicated catalyst stnppers or a common stπppei to lender the catalysts substantially free of entrained hydrocarbons

In a furthei embodiment, the legenerated catalyst with coke content of less than 0 4 wt° o is obtained by burning a portion of the spent cataly st from the first stnpper. the spent catalyst from the second stnpper or the common stnpper m a turbulent oi fast fluidized bed regenerator in the presence of air oi oxygen containing gases at a temperature ranging from 600°C to 750°C

In anothei embodiment, the catalyst between the fluidized bed nser reactors, stnppers and the common regenerator is contmuously circulated through standpipe and slide valves

In yet anothei embodiment, the cntical catalytic cracking conditions m the first reactor mcludmg mixed regenerated catalyst result m very high selectivity of middle distillate range products and conversion of hydrocarbon products of boiling point less than or equal to 370°C at lower than 50 wt% of the fresh feed

In another embodiment, the catalyst compnses of a mixture of commercial ReUSY zeolite based catalyst having fresh surface area of 110- 180 m^"'gm . pore v olume of 0 25-0 38 cc/gm and average particle size of 60-70 micron along with selective acidic bottom upgrading components in the range of 0-10 wt%

In still another embodiment, the unconverted heavy hydrocarbon fraction from second user recycled into the second nser ranges from about 0-50 wt°o of the mam feed rate to the second user, dependmg on the nature of the feedstock and operating conditions kept m the risers

In yet another embodiment, amount of steam for feed dispersion and atomization m the first and the second nser reactors is in the range of 1-20 wt% of the respective total hydiocarbon feed dependmg on the quality of the feedstock In further embodiment, the spent catalyst resides in the stnpper for a penod of upto 30 seconds

In anothei embodiment, pressure m the first and second nser reactors are m the range of 1 0 to 4 0 kg/cπr(g)

In yet another embodiment, the regenerated catalyst entering at the bottom of the second nsei reactor has coke of about 0 1-0 3 wt% at a temperature of about 600-750°C and is lifted bv catalytically inert gases

In a furthei embodiment, the combmed Total Cycle Oil (150-370°C) product which is a mixture of Heavy naphtha (150-216°C) and Light cycle oil (216- 370°C), has higher cetane number than that from conventional distillate mode FCC unit and other properties such as specific gravity, viscosity, pour pomt, etc are m the same ran ^lgce^ as that of commercial distillate mode FCC unit

In still another embodiment, changing the cut pomt of the TCO from the first user to 120-370°C. processing 370°C-r part of the first nser product m the second riser, and changing the cut pomt of TCO from second nser to 120-390°C. the yield overall combined TCO product mcreases by 8- 10 wt% and the combined TCO pioduct has same properties but improved cetane number as that of TCO from commercial distillate mode FCC unit

Brief description of the accompanying drawings

The invention is illustrated herembelow with reference to the following accompanying drawings, wherein

Fig 1 shows conventional fluid catalytic crackmg smgle nser system Fig 2 shows a fluidized catalytic crackmg two nser system of the present invention Fig 3 is a graph showing the ratio of TCO Yield / Yields of (Dry

Coke) Vs -370°C conversion with first nser feed at two different temperatures (425°C & 490°C) Fig 4 is a graph showing the ratio of TCO Yield ' Yields of (Dry gas-LPG-

Gasohne-t-Coke) Vs -370°C conversion with second nser feed at two different temperatures (490°C & 510°C)

Description of Fig.l

In the conventional Fluid Catalytic Crackmg (FCC) unit, fresh feed ( 1) is injected at the bottom of the nser (2) which comes mto contact with the hot iegenerated catalyst from the regenerator (3) The catalyst along with hydrocarbon product vapors ascends the nser and at the end of the nser spent catalyst is separated from the hydrocarbon vapor and subjected to steam snipping The hydrocarbon vapors from the nser reactor is sent to a mam fractionator column (4) for separating mto the desned products. The snipped catalyst is passed to the regenerator (3) where the coke deposited on the catalyst is burnt and the clean catalyst is circulated back to the bottom of the nser

The fluidized catalytic crackmg two riser system of the invention is schematically shown m Fig 2 and descnbed in detail herembelow

The fluidized bed catalytic crackmg system for the production of high yield of middle distillate products compnsmg hydrocarbons having carbon atoms m the iange of C₈ to C₂₄ from heavy petroleum feeds, by a process as defined in claim 1, said system compnsmg at least two nser reactors ( 1 and 2) wherem, a fresh feed is introduced into the first nser reactor (1), typically, at the bottom section above regenerated catalyst entry zone through a feed nozzle (3), and at the end of the first nsei ieactoi ( 1). the spent catalyst is quickly separated from hydrocarbon product vapors using separating devices (4) and subjected to multistage steam stnppmg to remove any entramed hydrocarbons, and a conduit (5) feeds a part of the said stnpped catalyst mto a regeneratmg apparatus (7) and the other part of the stnpped catalyst from the conduit (5) travels through anothei conduit (6) into a mixing vessel (10), and thereafter, the mixed catalyst from the mixing vessel ( 10) tiavels through a conduit ( 19) and is fed to the bottom of the first nser reactor ( 1), the hydrocarbon product vapors from the first nser leactoi ( 1) which are separated from the catalyst in the separating devices (4) are fed to a vacuum or atmosphenc distillation column (13) through conduit (12) whereby the first cracked hydrocarbon products are separated mto a first fraction compnsmg hydrocarbons having boiling points less than or equal to 370°C and a second fraction compnsmg uncracked hydrocarbons with boiling pomts greater than or equal to 370°C, the said second fraction compnsmg uncracked hydrocarbon products is fed through feed nozzle (16) mto the bottom of second nser reactor (2) above the regenerated catalyst entry zone, and the regenerated catalyst from the regenerating apparatus (7) is fed to the bottom of the second nser reactor (2) through a conduit (9), and subsequently, the hydrocarbon products of the second nser reactor (2) are separated from the catalyst m separating devices (11), and the cracked products of the second nser leactor (2) along with the products of the first fraction of the first nser reactor (1) compnsmg hydrocarbons with boiling pomts less than or equal to 370°C are fed to a mam fractionator column (15) which separates the said products mto dry gas. LPG, gasolme. heavy naphtha, light cycle oil. heavy cycle oil. and slurry oil. and the entire heavy cycle oil and full or part of the slurry oil consisting mainly of hydrocarbons with boiling pomts greater than or equal to 370°C are recycled back to the second nser reactor (2) through a separate feed nozzle ( 1 ") located at a point lower than the position of introduction of mam feed, and the feed and cracked product vapors travel along with the catalyst, mto the reactor wherein the spent catalyst separated from product vapors of the second nser reactoi (2) in separating devices and the spent catalyst is subjected to multistage steam stnppmg for removal of entramed hydrocarbons and the stnpped catalyst travels through a conduit (18) mto the regeneratmg apparatus (7). wherem the coke on catalyst is burnt m the presence of an and/or oxygen containing gases at high temperature, and the flue gas from regeneration is separated from the entrained catalyst fines in separating devices (23) and the flue gas leaves from top of the regeneratmg apparatus (7) through a conduit (22) for heat recovery and ventmg through stack, the hot regenerated catalyst is withdiawn from the legenerating apparatus (7) and divided mto two parts, one going to the mixing vessel ( 10) through the conduit (8) and the other directly to the bottom of the second nsei leactor (2), and the mixed catalyst from the mixing vessel ( 10) is fed through the conduit ( 19) to the inlet of the first nser reactor (1). controlling the catalyst bed level m the individual or common stnppei. the catalyst cnculation rate from the common regenerator and the quantity of the spent and regenerated catalyst entering mto the mixing vessel ( 10) usmg slide valves placed on the conduits and thereby producmg high yield of middle distillate products

At the bottom ^'Y^' section of both the nsers ( 1&2), steam is used to lift the catalyst m upward direction upto the feed entry zone .Also steam is used m the feed nozzles (3.16 & 17) for atomization and dispersion of the feed. The quantity of the steam flow mto the respective nsers (1&2) are vaned dependmg on the feedstock quality and the desired velocity m the nsers

As an example, the system designed to practice the process of the invention has been descnbed employing only two nser reactors It is pertinent to note that m practice, user reactois of desired number may be connected to the second user reactor so that the unconverted hydrocarbons obtamed from the second nser may be further treated m accordance with the process descnbed herein above and eventually, substantiallv the pure middle distillate products may be obtained m high vield from the oπgmal feed

In catalytic crackmg processes usmg zeolite based catalyst the reactions proceed sequentially High boiling large feed molecules first enter the catalyst through relatively laige poies which allows pre-crackmg to form intermediate middle distillate iange molecules which are further cracked to lighter molecules conespondmg to Dry gas. LPG and Gasolme Ideally, middle distillate yield can be mcreased. if it s ciackmg to lighter products is restncted Any attempt m this iegard is hkelv to ieduce the conversion, resultmg m higher yield of unconverted products Conventionally, iecyclmg of unconverted fraction has been practiced to impiove the overall conversion The seventy reqmred for crackmg of the unconverted iecycled fraction is adequate to produce sigmficant quantity of gasolme and LPG by over-cracking of middle distillate range product It also promotes hydrogen transfer reactions producmg aromatics m middle distillate range products and therefore, detenorates the cetane quality To summanze. it may be noted that maximization of intermediate product middle distillate is more challenging as compared to maximization of gasolme

In distinction to other pnor art processes, the present invention provides a process foi producing maximized quantity middle distillate through catalytic crackmg of heavy hydrocarbon fractions employing multiple nsers The applicants realized that the middle distillate selectivity is highei only at lower conversion In fact, the ratio of yield of Total Cycle Oil (TCO 150-370°C) to the sum of other products, (such as. dry gas, LPG. gasolme and coke) increases as the conveision i educes Moreover, nser temperature has dramatic impact on the selectivity At same conveision the applicants have found that middle distillate selectivity improves significantly as nser temperature is reduced The applicants have also investigated the role of coke on regenerated catalyst (CRC) and discovered that there is an optimum CRC for maximum yield of TCO (Ref Ind Chem Res . 32, 1081, 1993) Fmally. the applicants have arnved at some specific conditions (compnsmg of very low nser temperature, low contact time. low catalyst oil ratio, higher CRC, etc ) and type of the catalyst with which vield of TCO is maximized

Accoidmg to the piesent invention, petroleum feed stocks such as Vacuum Gas Oil (VGO). Coker fuel oil. CokerYisbreaker heavy gas oil. Hydiocracker bottom, etc is catalytically cracked m presence of solid zeolite catalyst with or without selective acidic bottom crackmg components m multiple nser-reactors The feed is first preheated at a temperature m the range of 150-350°C and then injected to pneumatic flow nser type crackmg reactor with residence time of 1-8 seconds and preferably of 2-5 seconds. At the exit of the nser, hydrocarbon vapors are quickly separated from catalyst for minimizing the over crackmg of middle distillate to lighter products

The product from the first nser is separated in a fractionator to at least two streams, one compnsmg hydrocarbons having boiling below 370°C and the other compnsmg hydrocaibons having boiling pomts greater than 370°C The removal of hydrocarbons having boiling pomts less than or equal to 370°C products reduces the chance of over-cracking of middle distillate range molecules to lighter products The unconverted fraction compnsmg hydrocarbons having boiling points greater than or equal to 370°C fraction f the first nser is pre-heated and then injected to the second riser reactor with residence tune of about 1-12 seconds and preferably in the range of about 4-10 seconds, through the feed nozzles located at a higher elevation In the second nser. the regenerated catalyst is contacted with the recycle stream of unconverted heavy hydrocarbons from the second nsei at a relatively lower elevation of the nser This allows preferential cracking of the recycle components under high severity conditions (e g , higher temperature, higher dynamic activity of the catalyst owing to low coke on iegenerated catalysts) at the bottom of second nser Typically, recycle iatio is maintained in the range of 0-50% of the feed throughput in the second riser

Steam and'oi watei. m the range of 1-20 wt% of feed is added for dispeision and atomization in both the risers dependmg on type of feedstock The desned velocity m the nsers. especially in the first nser is adjusted by addition of steam

The hydrocarbon product vapor from the second nser is quickly quenched with water/other hydrocarbon fraction and separated for minimizing the post nser non-selective crackmg The product from the second nser and the product boilmg below 370°C from the first nser are separated m a common fractionator mto several products, such as Dry gas, LPG, Gasolme. Heavy naphtha. Light Cycle Oil and cracked bottom Part of the unconverted bottom product (370°C^- fraction) from the second fractionator is recycled to the second nser and remaining part is sent to rundown after removal of catalyst fines

The spent catalyst with enti anted hydrocarbons from the nser exit is then passed through a common or separate stnppmg section where counter cirrrent steam stnppmg of the catalyst is earned out to remove the hydrocarbon vapors from the spent catalyst The catalyst residence time m the stnppers is reqmred to be kept m the lower side of preferably less than 30 seconds This helps to minimize undue thermal crackmg reactions and also reduces the possibility of over- cracking of middle distillate range products Stnpped catalyst is then passed to a common dense or turbulent fluidized bed regenerator where the coke on catalyst is burnt in presence of air and or oxygen contammg gases to achieve coke on regenerated catalyst (CRC) of lower than 0 4 wt% and preferably m the range of about 0 1 - 0 3 wt% A part of the regenerated catalyst is dnectly circulated to the second user reactor via standpipe / slide valve at a temperature of 600 -

As mentioned earliei. there is an optimum CRC at which maximum TCO yield is obtained In order to extract maximum TCO from the first nser. CRC is required lo be maintained at relatively higher level, m the range of 0 2- 0 8 wt% depending on catal st and operating conditions In the second nser. the desnable CRC is ielatively lower (m the range of 0 1 - 0 3 wt%) m order to utilize the full activity potential of the catalyst Also the temperature of the regenerated catalyst entenng to the two nseis are different The lower temperature and highei CRC of the catalyst enteimg to the first nser is achieved by mixing a part of the stnpped catalyst from the first nser / common stnpper with regenerated catalyst in a separate vessel equipped with fluidization steam and circulating the mixed catalyst to the bottom of the first nser via stand pipe / slide valve The mixed catalyst enters at the bottom of the first nser with a temperature m the range of 450 - 575°C (preferably in the range 475 - 550°C) and CRC of lower than 0 8 wt% (preferably in the range of 0 25 - 0 5 wt% dependmg on type of catalyst) Anothei option of controlling the catalyst return temperature m the first nser is to employ catalvst cooler so that catalyst/oil ratio could be controlled almost independently However, the mixing vessel is preferred smce it acts as second stage stnpper and helps to adjust the coke level on the catalyst

Pnor to the injection of the 370°C- fraction of the first nser product, the fresh regenerated catalyst is contacted with the recycle stream of unconverted hydrocarbons from the second nser at a relatively lower elevation of the nser The iecycle components are preferentially ciacked at the high seventy conditions pievailmg in the second nser bottom before the injection of 370°C- fraction of first nser product Typically recycle ratio is mamtamed m the range of 0 - 50% of the second reactor feed throughput dependmg on the type of the feed to be processed and the conversion level in both the reactors. If the recycle quantity is less, it may be injected along with the main feed i.e.. 370°C- fraction of first riser product.

In the present invention, the first nser operates m the range of 150 - 350 hr^"' weight hourly space velocity (WHSV). 2 - 8 catalyst to oil ratio, 400 - 500°C user top temperature to convert the feedstock to selectively cracked product including 35 - 45 wt% mm. TCO yield and 40 - 60 wt% 370°C- (bottom) yield. The second user operates in the range of 75 - 275 hr^"1 WHSV. 4 - 12 catalyst to oil ratio and 425 -525°C nser top temperature. The absolute pressure in both reactors are 1 - 4 kg/cm^" (g). Steam and / or water, in the range of 1 - 20 wt% of feed is added not only for dispersion and atomization of feed but also to attain the desired fluidization velocity in the risers, especially in the first riser bottom. It also helps in avoiding the coke formation or catalyst agglomeration.

Comparison of major process conditions of the process of the present invention with conventional FCC & multi stage process is shown below :

Table - 1

Multistage process of the present invention FCC

Process first reactor second reactor

Range Preferred , Range _, Preferred Range '• Range ' ' Range

WHSV. hi-^"' 150 - 350 1 200- 300 75 - 275 , 120 - 220 125 -200 Cataty st/Oil 2 - 8 3 - 5 - 4 -12 5 - 8 4 - 8 ratio (w/w) Riser temp.. C 400 - 500 - 425 - 475 425 - 525 ! 460 - 510.490- 540 Steam injection. 1-20 8-12 1-20 4-8 0-10 wt% of feed Lse of multiple riser concepts is not new, as each researcher has employed it for diffeient purposes The present invention utilizes dual or multiple nser systems for exclusive maximization of middle distillate products Bemg an intermediate pioduct middle distillate iange molecules have a tendency to undergo further crackmg There is always a trade off between maximization of an intermediate iange product and minimization of bottom unconverted part This invention includes the sequence of operation and operating conditions for control of over- ciacking of middle distillate m the first nser and upgradation of heaviei molecules to middle distillate in the second nser This invention provides a novel scheme for operation of two or multiple nsers at entirely different operating conditions with a common regenerator Lse of so much lower tempeiature ciackmg is unusual so far However, the applicants have found that ieaction tempeiature has a predominant effect on the over crackmg of middle distillate range products For example, at 40 wt% of 370°C- conversion, the wt% yield ratio of TCO and all other products, (I e., Dry Gas, LPG. Gasolme & Coke) except TCO and bottom (subsequently refened as TCO/Rest ratio) are m the range of about 3 0 - 3 5 and about 1 5 - 1 8 at reaction temperatures of 425°C and 490°C iespectively The difference m the above ratio is nanowed down as the conveision inci eases (Fιgure-3)

Therefore, for maximizing TCO. low reaction temperature and catalyst to oil iatio as well as low catalyst activity is desirable The applicants identified that lower catalyst / oil iatio ( 2 - 8) and higher WHSV of ( 150 - 350 hi ') along with lowei nsei temperature in the first nser of the process of the present invention aie very important to achieve very low degree of over cracking for producmg maximum middle distillate iange components The applicants also observed that the TCO/Rest ratio is significantly affected by the 370°C- conversion level Foi example, for a given catalyst and reaction temperature, if 370°C- conversion is 40%, the TCO/Rest iatio is as high as 3 2 which comes down to about 1 3 when 370°C - conversion is increased to 70% This shows that restnchng the conversion in the first stage nser upto 40 - 45% is very important to maximize the yield of middle distillate

In the second nsei. the operating conditions need to be diffeient for upgradation of relatively less crackable heavy matenal to lighter products However, undue increase in seventy parameters will lead to conversion to LPG and Gasolme The applicants have discovered that operation at an intermediate seventy as compaied to gasolme maximization mode FCC operation is absolutely necessary The applicants have also found that m order to reduce the yield of unconverted bottom and improve the middle distillate selectivity, recycle at a lower elevated entry point at the bottom of the second nser is very much effective This allows the crackmg of the recycled heaviest fraction m presence of regenerated catalyst at lelatively higher temperature and lower CRC which unproves the dynamic activity of the catalyst and offers maximum crackmg of the recycled feed After crackmg of the recycled part, the catalyst temperature comes down due to utilization of part of the heat for vaponzation and endo hermic crackmg reactions of the recycled feed Also, the coke on catalyst mcreases which essentially blocks some of the active sites and thereby reduces the dynamic activity of the catalyst The contacting of catalyst having relatively lower temperatuie and higher coke on catalyst with the mam feed compnsmg the fraction of the first user of hydrocarbons with boilmg pomts greater than oi equal to 370°C. assists to improve the selectivity of middle distillate range products out of the second riser This contactmg pattern is unique and highly effective m mci easing the overall yield of the middle distillate and reducing yield of the unwanted slurry oil

In the present invention, the delta coke (defined as the difference m coke content of spent and regenerated catalyst) is low due to lower coke make m the extremely low seventy crackmg in the first nser which is expected to keep the regenerator temperature at relatively lower level as compared to the conventional FCC operation using similar type of feedstocks However, overall lower catalyst oil ratio is likely to compensate this effect and thereby mamtam the regenerator temperature at least to the same level as that of conventional FCC as requued for burning of coke on catalyst

Furthei details of feedstock, catalyst, products and operatmg conditions of the process of the present invention are descnbed below

Feed Stock:

Feed stock foi the piesent invention mcludes hydrocarbon fractions starting from carbon no 20 to carbon no 80 The fraction could be straight run light and heavy Vacuum Gas Oil, Hydrocracker bottom. Heavy Gas Oil fractions from Hydrociackmg, FCC, Visbreaking or Delayed Cokmg The conditions m the process of the present invention are adjusted depending on the type of the feedstock so as to maximize the yield of middle distillate Details of the feedstock properties are outlmed m the examples given herembelow The above feed stock types are for illustration only and the invention is not limited in any manner to only these feed stocks

Catalyst:

Catalyst employed in the process of the present mvention predominantly consists of Y-zeohte m raie earth ultra-stabilized form Bottom crackmg components consisting of peptized alumina, acidic silica alumina or T- alumina or a mixture thereof are also added to the catalyst formulation to produce synergistic effect towards maximum middle distillate under the operating conditions as outlmed above It may be noted that both the first and second stage nsers are charged with same catalyst The pore size range of the active components namely, Re- USY zeolite and bottom selective active materials are in the range of 8 - 1 1 and 50 - 1000 angstrom respectively. The typical properties of the Y-zeolite based catalyst are given in Table-2. Table - 2

Surface .Area. m7g, Fresh ! 1 10 - - 180

Steamed 100 - - 140

% Crystallnity Fresh 10 - - 15

Steamed 8 - 12

Umt Cell Size. °A Fresh 24.35 - - 24.75

Steamed 24.2 - - 24.6

Micro-pore area. m^"/g, Fresh 65 - 100

Steamed 60 - - 90

Meso-pore area. m^"/g, Fresh 45 - • 80

1 Steamed 40 - ■ 50 ,

Pore volume, cc/gm 0.25 - - 0.38 i

The active components in the process of the present invention catalyst are supported on inactive materials of silica/alumina/silica-alumma compounds including kaolinites. The active components could be mixed together before spray diying or separately binded, supported and spray-dried using conventional spray drying technique. The spray-dried micro-spheres are washed, rare earth exchanged and flash dried to produce finished catalyst particles. The finished micro-spheres containing active materials in separate particles are physically blended in the desired composition. The prefened range of physical properties of the finished fresh catalyst as required for the process of the present invention:

Particle size range, micron 20-120

Particle below 40 microns, wt% . < 20 Average particle size, micron 50-80

Average bulk density, micron 0.6 - 1.0 Typically, the above properties and other related physical properties, e.g., attrition resistance, fludizability etc. are in the same range as used in the conventional FCC process.

Products:

The main products in the process of the present invention is the middle distillate components namely, Heavy Cracked Naphtha (HCN : 150 - 216°C) and Light Cycle Oil (LCO : 216 - 370°C). The sum total of these two fractions which is called as Total Cycle Oil (TCO : 150 - 370°C) is obtained with a yield upto 50 - 65 wt% of the feed. The other useful products of the process are LPG (5 - 12%) and Gasoline (15-25 wt%). Range of other product yields from first and second stage risers are summarized in Table - 3 :

Table -3

Dry Gas (d -r-C ) 0.1- -0.35 1-1.5 0.5- - 1.5

LPG (C₃ + C₄) 3 - 8-12 5 - 12

-«

Gasoline (C₅ -150°C) 10- -15 25-30 15- -30

Heavy Naphtha, ( 150-216°C) 8 - -10 10-13 10- - 15

Light Cycle Oil. (216-370°C) 35- -45 25-35 40- -50

Total Cycle Oil (150-370°C) 45- -50 30-40 50- -65

Bottom (370°C+) 40- -60 10-20 5 - 15

Coke 1- 2-5 2 - -4 The invention and its embodiments are described in further detail hereunder, with reference to the following examples, which should not be construed to limit the scope of the mvention in any manner. Various modifications of the invention that may be apparent to those skilled in the art are deemed to be included within the scope of the present invention.

Example-1

Yield of middle distillate at different conversions in conventional FCC operation

This example illustrates the change in yield of the middle distillate product (TCO) at different conversion levels under conventional FCC conditions. -216°C conversion is defined as the total quantity of products boiling below 216°C including Coke. Similarly -370°C conversion is defined as the total quantity of products boiling below 370°C including Coke. The experiments were conducted in standard fixed bed Micro Activity Test (MAT) reactor described as per ASTM D-3907 with minor modifications indicated subsequently as modified MAT. The catalyst to be used is first steamed at 788°C for 3 hours in presence of 100% steam. The physico-chemical properties of the feed used in the modified MAT reactor are given in the Table - 4 & 5.

Table - 4

Density @ 15°C, gπvcc 0.8953

CCR, wt% 0.32

^' Sulfur. wt% 1.12

Basic Nitrogen, PPM 366

Paraffins. wt% 44.4

Naphthenes, wt% 18.1

Aromatics. wt% 37.6 Nickel. PPM < 1 Vanadium. PPM < 1

The runs were taken at a reaction temperature of 495°C. feed mjection time of 30 seconds with WHSV in the range of 40 - 120 hr^"1 Catalysts used in this example are catalyst A & B which are commercially available FCC catalyst samples having properties as shown in the Table-6

Table -5

ASTM Distillation (Dl 160)

Volume % Temperature, C

IBP 299

5/ 12/15/20/30/40 342/358/371/381/401/418 50/60/70/80/ 90/95 432/444/458/474/497/515 FBP 550

Table - 6

Catalyst - A Catalyst - B

Surface Area. m^"/gm Fresh 170103 272 Steamed 208

Pore A^'olume. cc/gm 0.22 0.26

Cιystalιnιty,% Fresh 18.9 277 Steamed 23.2

UCS. ^UA Fresh 24.61 2456 Steamed 2432 24.31 Chemical Analysis. wt%

A1₂0₃ 56.5 30.85

Re₂0₃ 1.44 1.03

Fe 0.49 0.53

APS. microns 74 , 77

The product yields along with conversions are given in Table-7 wherein it is obseiyed that as in both -216°C and -370°C conversion increases. TCO yield increases upto an optimum value and thereafter, it reduces with mcrease m conversion. TCO bemg an intermediate product undergoes further cracking as reaction seventy mcreases. Therefore, m order to maximize TCO yield, the over- cracking is to be restricted.

Table - 7

Product Yield, Catalyst A Catalyst B wt% 1

W/F, Mm. 0 51 0 62 0.94 0.44 0.51 0.63 0.94

Hydrogen 0 018 0 021 0 041 1 0.025 0 025 ; 0.033 0.046

Dry gas 0.44 0.56 1 14 j 0.59 ¹ 0.64 0.86 , 1.46

LPG / . j j 8.82 13.61 6.18 6.97 , 10.09 ϊ 12.34 -

Gasoline 19 32 23.43 30.78 17.20 20.50 25.03 30.94

| 1

TCO 40.09 41.53 37.79 , 36.33 37.97 J 39.94 i 37.67

Bottom (370°C-) 31.81 24.52 14.25 1 38.73 32.82 ! 22.80 14.92 j

Coke 0 99 1.13 2.39 ' 0.95 1.08 1.25 2.61

-216°C 40 17 47 50 62.45 34 96 40.34 , 49 98 1 60.99

Conversion

-370°C 68 19 75 48 85.75 61.27 ¹ 67 18 ! 77.20 85.08

1 1

Conversion Example-2

Effect of reaction temperature on middle distillate yields at same conversion

This example illustrates the effect of reaction temperature on the yield of middle distillate at a given -216°C conversion. The experiments were conducted in the modified MAT reactor with the same feed as mentioned m Example- 1. at two different temperatures, v z.. 425 C and 495°C. Catalyst employed here is catalyst C which is commercially available FCC catalyst of following properties as shown in the Table - 8.

Table - 8

C atalyst - ^■ C

1 Surface Area, m^"/gm Fresh 172

Steamed , j 1 19 i

I Pore volume, cc/gm I 0.32

³ ;

Crystallinity, % Fresh 13.80 1 Steamed 10.20

UCS °A Fresh 24.55 Steamed 24.31

Chemical Analysis. wt%

^' RE₂0₃ 0 69

! A1₂0₃ 36.40

Na₂O 0.11

Particle size, micron / wt^c V

. -20 / -40 / -60 / -80 / - 105 I / -120 3 / 16 / 32 / 56 / 77 / 86

APS, micron 76 Table - 9

Temperature. 'C ^→ 425 495

-216°C conversion. \vt% 30 50 30 50

W/F. Mm 1 1 2 7 0 10 0 5

Yield Pattern. wt%

Dry gas 0 20 0.42 0 38 0 56

LPG 4 10 9 1 5 07 10 72

Gasoline 14 94 23.52 16 00 24 58

Heavy naphtha 9 50 14 27 7 1 1 1 1 20

The conversion was varied by changing W/F ratio. The product yields are compared at same -216°C conversion but at different temperatures. It is noted from Table-9 that at higher temperature. TCO yield and more importantly the TCO/Rest ratio (the ratio of TCO yield and yield of other products e.g., Dry gas, LPG. Gasoline and Coke except bottom and TCO) are much lower in case of higher reaction temperature. For example, at a given -216°C conversion. TCO yield at 425°C temperature is about 6 - 10% higher than that at 495°C. The other significant pomt is that at low temperature of 425°C, it has been possible to get 46% TCO yield (per pass) at 50% -216°C conversion. Similarly, there is a sigmficant improvement in TCO/Rest ratio for 425°C as compared to that of 495°C at same conversion. This clearly demonstrates that m order to conserve middle distillate range molecules, low reaction temperature is essential. Example-3

First stage riser cracking conditions

This example illustrates the sigmficance of first stage riser crackmg conditions e g . temperature catalyst/oil iatio and conversion on the yield of middle distillate and othei products while employmg commercially available FCC catalysts A and C properties of which aie descnbed in Example- 1 &. 2 iespectively The tests were conducted m modified fixed bed MAT umt with bame feed as descnbed m Example- 1 Yield data were generated at different conversion level for the catalysts as mdicated above and the yields of different products were obtained TCO/Rest ratios at different conversion levels are plotted in Fιgure-3 from which it is observed that for both the catalysts, the TCO/Rest iatio increases as the -370 C conversion is reduced Therefore, it is important to note that the per pass -370 °C conversion m the first stage nser should be kept below 45% and preferably below 40%

From Fιgure-3. it is also observed that the TCO/Rest ratio is a strong function of the reactor temperature for a given conversion and catalyst For example, with catalyst C while i educing reaction temperature from 490 to 425°C. the TCO/Rest ratio is increased from 3 4 to 3 75 at about -370 C conversion level of 40% This cleaily shows that for the first stage crackmg, the ieaction temperature should be kept lower preferably m the range of 425 - 450°C

Example - 4

Catalyst characteristics for middle distillate maximization

One of the important obseivation as illustrated in Example-3. is that for maximization of middle distillate yield, it is necessary to restnct the per-pass conversion withm 40 - 45% and operate the first stage nser at lowei reaction temperatuie The low reaction temperature coupled with high coke on regenerated catalyst leads to lower dynamic activity of the catalyst Therefore. the desired catalyst should have high intrinsic activity. However, the problem is that high active catalysts are not usually diesel selective. In this example, we illustrate the importance of catalyst characteristics to obtain higher yield of middle distillate out of the dual / multi - stage risers.

MAT activity is measured in ASTM MAT unit using a standard feedstock and defined as the wt% of products boiling below 216°C mcluding coke at ASTM conditions. .All other experiments were conducted at the temperature of 425°C in the modified MAT reactor with the same feed as described in Example- 1 and different catalysts. The important properties of the catalysts and the yield / conversion data are compared in Table- 10.

Table-10

Catalyst - A Catalyst-C Catalyst-D Catalyst-E

Surface Area. m^"/gm 103 119 1 10 20

Zeolite Area. m7gm 59 80 62 -

Rare earth content.

1.44 0 69 1 40 -

Matrix Area, rn^'/gm 44 39 48 -

Zeolite / Matrix ratio 1 34 2 05 1 29 -

MAT Activity 71.38 74 02 70 19 13 55

TCO Yield at 40% -370T 31.00 32.01 30.90 31.20

Conversion

TCO yield/Rest ratio at 40% 3 44 4 00 3 39 3 30

-370"C conversion

W/F for 40% -370°C 0 22 0.25 0 22 3 5 conversion

Table -11

Catalyst-A Catalvst- Catalvst- C D

TCO Yield at 80% -370 Conversion 58.45 54.78 43.0

TCO yield Rest ratio at 80% -370^υC 0.95 0.80 1.08 conversion It is seen that the zeolite/matnx ratio, TCO yield at 40% -370°C conversion. TCO Rest iatio aie in the order of C > A > D In catalyst C. the available active matrix is adequate to crack the large molecules which are crackable undei the prevailing operating conditions but it requires slightly higher 'F ratio Highei zeolite quantity is also synergistically taking part in the over all ciackmg activity but the conversion of middle distillate to lighter pioducts is not increasing corresponding to higher zeolite content due to lower temperature However foi catalyst-E whose activity is extremely low, at 40% of-370°C conversion, both TCO yield and TCO/Rest ratio is comparable to those with the higher active catalysts But W/F ratio required to achieve 40% -370°C conversion is much higher which is difficult to achieve At comparable W/F ratio, -370°C conversion will be very low, producmg very low amount of TCO Therefore, such low active catalyst is not useful for producmg maximum distillate

Expenments with catalysts A. C & D at a reaction temperature of 495°C conespondmg to the second nser conditions were taken and the TCO yield and TCO/ Rest ratio are compared at -370°C conversion of 80% m Table- 1 1 Both the TCO yield and TCO Rest ratio are found to be m the order of D > A > C It may be noted that the zeolite / matrix ratio is just m the reverse order 1 e . C > A > D The higher quantity of zeolite as well as the high zeolite/matnx iatio m catalyst C. is resulting m overcrack g of middle distillate range molecules mto lighter products For a given -370°C conversion, the -216°C conveision is much higher for catalyst C It is quite clear that the catalyst which is supposed to be the best m the first nsei conditions, may not be that much good for the second nser conditions as for as TCO maximization is concerned This demonstrates that m oider to achieve maximum TCO and minimum Bottom yield, some optimization of the catalyst properties is essential Example- 5

Impact of basic nitrogen compound on middle distillate yield

It is generally conceived that low activity of the catalyst is deshable for maximum distillate yield. Basic nitrogen compounds present in feed stock interact with the catalyst at reaction conditions leading to loss of the active acid sites and hence decrease of catalyst activity. Two feed stocks were prepared containing 200 and 700 PPM pyridine respectively. The experiments were conducted in the modified MAT reactor with catalyst C using the same feed stock as mentioned in Example- 1. but containing different PPM of pyridine, at the temperature of 425°C. The conversion and yield data are shown m Table- 12.

Table-12

Feed without Feed containing 200 Feed containing 700 Pyridine 1 PPM Pvridine PPM Pyridine

-216^UC conversion at 40% 13.50 i 14.15 13.99

1 -370°C conversion

TCO Yield at 40% 32.00 30.90 29.98 -370°C conversion

1

TCO/Rest ratio at 40% 4.00 3.39 3.00 -370°C conversion

W/F to achieve 40% - 0.25 0.30 0.38 370 C conversion

It is observed that both TCO and TCO/Rest ratio are decreasing as the feed basic nitrogen content is increasing. However, at 40% -370°C conversion, -216°C conversion is increasing with increase in basic nitrogen in feed upto 200 PPM after which it reduces marginally at 700 PPM of pyridine in feed. This is due to the irreversible adsorption of the nitrogenous basic compounds leading to preferential destruction/poisoning of the strong acid sites, which are responsible for heavy molecule cracking. This is reflected in the higher W/F requirement to achieve 40% -370°C conversion. However, so called relatively weaker acid sites which do not get affected by basic nitrogen, helps in cracking of middle distillate range molecules at higher W/F resulting higher -216°C conversion. In case of 700 PPM pyridine containing feed, even some of the relatively weaker acid sites are getting affected reducing both -216°C and -370°C conversion as compared to the 200 PPM pyridine containing feed case. This example demonstrates that just activity reduction may not lead to higher middle distillate yield.

Example - 6

Impact of cracking conditions for second stage riser operation

This example illustrates the significance of second stage riser cracking conditions e.g., temperature, catalyst/oil ratio and conversion on the yield of middle distillate. The tests were conducted in modified fixed bed MAT umt as described in Example-1, using catalyst C, at the temperature of 425, 490 and 510°C. The feed stock used is 370°C^" product obtained from first stage cracking in circulating riser FCC pilot plant, the properties of which is summarized in Table- 13. Product yields data were generated at different conversion levels at different temperatures for catalyst C and according the TCO/Rest ratios at different conversion levels are plotted in Figure-4.

Table - 13

Density, gm/cc @ 15°C i 0.903

CCR. wt% 0.43

Sulfur. wt% i 1.75

Olefrns, wt% Nil

Saturates. wt% > 59.0

Aromatics. wt% ; 41.0

From the Figure-4. it is observed that at a given temperature, the TCO/Rest ratio increases as the -370^ϋC conversion reduces. Also, at a given -370°C conversion, TCO/Rest ratio improves as the reaction temperature reduces. For example, at about -370°C conversion of about 55%. TCO/Rest ratio increases from 1 22 to 1 34 as the temperature is reduced from 510 to 490°C This clearly shows that even for the second stage crackmg, the reaction temperature should be kept prefeiably lowei However, it will also lead to generation of higher quantity of bottom at same T ratio At 425°C. W/F required to crack the 370°C- product from first stage crackmg along with the recycle stream (unconverted part from the second nsei) will be very high and hence difficult to achieve .Another important fact is that the mean average boilmg point (MeABP) of second nser combined feed is definitely higher than that of first nser Operation at lower temperature than the MeABP of the second nser combmed feed is not desnable as it will lead to non-selective thermal crackmg of the non-vaponzed feed producing higher quantity of Coke and Dry gas. Considering these, it has been established that m the second nser, the reaction temperature should be preferably kept m the range of 460 - 510°C

Example - 7

Combined effect of two stage cracking on middle distillate yield

In this example, the yields from two stage catalytic crackmg for maximization of middle distillate is demonstrated The expenments have been conducted using catalyst C m continuously circulating fluid bed pilot plant of feed rate 0 75 kg hr where both the nsei and regenerator are operated isothermally The feed is the same as mentioned m Example- 1 After first stage cracking at 425°C, the product is separated into 370°C- and 370°C- fractions In the second stage 370°C- fraction is cracked at 495°C usmg the same catalyst as used m the first stage The product yields from the first and second stage crackmg and also the combmed yields are given in Table- 14 D

Table-14

It is clearly seen that the ratio of yield of TCO and the sum of yields of Dry gas, LPG, Gasolme and Coke (TCO/Rest) is very high in case of the first stage crackmg which is essentially contnbuting higher TCO yield for the overall process For second stage crackmg, the TCO/Rest ratio is similar to that of conventional distillate mode FCC umt as the seventy required for minimizing the bottom yield is high enough to crack sigmficant portion of TCO produced from heavy molecule crackmg

The yield companson between smgle and dual nser cracking at similar -216°C conversion with same catalyst and feed is compared in Table- 15 It is seen that for same -216°C conversion. -370°C conversion is much higher resulting about 20% higher yield of TCO m case of two stage crackmg This establishes the workability of the concept of the present invention where process schemes, catalyst and operating conditions are such that TCO over-crackmg is restncted 36

with simultaneous the upgradation of heavy molecules to TCO range molecules. Here, the first riser operates to extract as much TCO as possible while minimizing the yields of lighter products and the second riser is operated to upgrade as much bottom as possible while maximizing the yield of TCO. This process overcomes the trade off between lower bottom yield and higher TCO yield.

Table- 15

Example - 8

Comparison of Micro-reactor & Circulating pilot plant data

This example shows the comparison of individual product yields obtained from Micro-reactor and circulating Pilot Plant using same catalyst and feedstock at similai" -216°C conversion range. From the data summarized in Table- 16. it is noticed that at similai- conversion, there is an excellent match in Gasoline, TCO and bottom yields. The main difference is coming in the yields of Dry gas, LPG and Coke. This is mainly due to the non-selective thermal cracking reactions occumng at the riser bottom as well as at the end of the riser in the pilot plant. This has resulted relatively higher yield of Dry gas and Coke in the pilot plant riser. This example demonstrates that so far the yields of TCO and un-reacted bottom are concerned, the inferences drawn based on either Micro-reactor or Pilot Plant data are going to be same.

Table-16

Pilot Plant data Micro-reactor data

Feed rate, gmymin 12.9 - -

CCR. gm/min 55.5 53.0 -

1

^' Cat/Oil (w/w) 4.29 3.98 -

1

' W/F. mm. - - 0.609 0.501

; Contact time, sec - 30 30

1 1

-216 Conversion, wt% 29.86 25.0 29.39 24.93

Product Yields. wt% . j

¹

Dry gas 0.62 0.36 0.17 0.13

LPG 8.28 6.29 9.96 8.61

Gasoline 11.82 10.7 12.00 10.65

Heavy Naphtha 7.15 5.92 5.80 4.62

LCO 27.3 26 28.29 26.61

TCO 34.45 31.9 34.09 31.23

370°C - 42.82 49 42.3 1 48.46

Coke 2.00 1.71 1.46 0.91

Example - 9

Comparison of the yields of present two stage process in present invention, commercial FCCTJ and two stage hydrocracker 38

The product yields of the present invention is compared with that of commercial distillate mode FCC and tvvo-stage hvdrocracker units m Table- 17. The data for the process of the present invention is the combined yield obtained fi-om two stage crackmg where the two nsers are operated at 425°C and 495°C respectively

Table - 17

Product yields. Distillate Present Yields, wt% of Distillate mode Present wt% of feed mode FCC process feed Hvdrocracker process

Diy gas 2.50 0.78 Dry gas 1.74 0 70

LPG 10.5 10.55 LPG 2.91 9.1 1

Gasoline 27.5 21.88 Gasoline 16.28 12.86

(C₅-150°C) (C₅-120°C)

Heavy Naphtha 12.5 14.28 (120-216°C) 18.41 a ( 150-216°C) (120-285°C) 27.91

LCO 30.0 40.73 (216-390^ϋC) 50.39

(216-370°C

TCO 42.5 55.01 (120-390^ϋC) 73.26 68.80

(150-370°C)

370^υC^" 12.75 8.82 370^υC 5.81 5.85

Coke 4.25 2.93 Coke 2.68

-216^υC conv 57.25 50.45 -216^υC conv.

-370^ϋC conv 87.25 91.18 -370^υC conv. 94.19 94.15

It is observed that m the process of the present invention, the TCO yield is higher by about 12 50% as compared to the commercial FCC umt. By varying the cut point of TCO from 150 - 370°C to 120 - 390°C as reported for Hvdrocracker unit, and processing the hydrocarbon product vapors having boiling points greater than or equal to 370°C of the first riser product in the second riser, the yield of TCO mcreases by about 14 wt% which is only about 5% less than that from the commercial Hvdrocracker unit Also, the conversion of hydrocarbon product vapors having boiling pomts less than or equal to 370°C is similar to hvdrocracker and better than distillate mode FCC unit. This demonstrates that without using external hydrogen and operating under very high pressure, it is possible to produce higher yield of middle distillate product which is close to that from a distillate mode two stage Hvdrocracker unit.

Example -10

Comparison of properties of TCO obtained in the process of the present invention with middle distillate products obtained from commercial FCCU and two stage Hvdrocracker

The properties of the TCO obtamed from the process of the present mvention is compared with TCO from commercial distillate mode FCC and Diesel from distillate mode two stage Hvdrocracker units which is given in Table- 18.

Table -18

Process of the : present Distillate Distillate invention mode FCC mode

Hvdrocracker

1 2 _^ j 4

TCO Middle TCO Diesel distillate

TBP cut pom ^ϋC 150 - 370 120 - 390 150 - 370 150 - 390

Density @ 15°C. 0.8793 0.8863 0.8654 0.835 gm cc

Pour point ^JC 0.7 36 0 - 2 6 - 10

Kinematic Viscosity 2.20 7.00 2.7 9.0 /3⁾. 50⁰ CST PONA .Analysis. wt% ' i

Olefrns 19.97 6.82 18.6 ! Nil

Sa rates 24.64 49.26 22.1 91

Aromatics 55.39 43.92 59.3 9

Cetane no. 36.22 38.39 28 - 30 63

Expectedlv. the quality of Diesel range product from Hvdrocracker is much superior in terms of cetane no., olefm and aromatics contents etc. than the cracked products without using hydrogen. Mainly, the high aromatics content in cracked middle distillate product contribute to poor cetane quality However, the viscosity and the pour point of Hvdrocracker Diesel is poor as compared to TCO from conventional FCC unit or the process of the present invention. From column 1 & 3, it is seen that the cetane no. of TCO obtained from the present process is higher by 6 units than TCO from conventional distillate mode FCCU. .All other properties including the pour point are almost in the same range. In column 2, the properties of the product fraction of 120-390°C range for the present process is listed. While cetane no. of this fraction is further higher, the pour point as well as the viscosity is very high. This has been mainly contributed by the hydrocarbon fraction of 370 - 390°C cut from the first riser product of the present process. The pour point as well as the viscosity of this product fraction is very high and hence its inclusion in the middle distillate product is not desirable. If we take the 120 - 370°C cut from the first riser product and the 120 - 390°C cut from the second riser (while processing the unconverted 370°C- part of the first riser product into the second riser), the pour point and the kinematic viscosity Λ> 50°C become 0.95°C and 2.44 CST respectively, which are almost same as that of 150 - 370°C product of the present invention as shown in the column 1 of Table- 18. Additionally, by this approach, the yield of the middle distillate increases from about 55 wt% to 63.6 wt% without any adverse impact on flash point.

Claims

im

A multi stage selective catalytic crackmg process for producmg high y ield ot middle distillate products having carbon atoms in the range of about C_s to C;₄. from heavy hydrocarbon feed stocks m the absence of added h diogen. said process compnsmg the steps of

i ) contacting pieheated feed stock with a mixed catalyst in a first nsei ieactoi undei catalytic cracking conditions including catalyst to oil iatio of 2 to 8 WHSV of 150-350 hr contact period of about 1 to 8 seconds and top tempeiature in the range of about 400°C to 500°C to obtam first cracked hydrocarbon products.

a ) separating the first cracked hydrocarbon products from the first nser reactor in a vacuum or atmosphenc distillation column mto a fust fraction compnsmg hydrocarbons with boilmg pomts less than or equal to 370°C and a second fraction compnsmg unconverted hydrocarbons with boilmg points greater than or equal to 370°C.

m i ciackmg the unconverted second fraction from the first nser ieactoi compnsmg hydrocarbons having boilmg points greatei than or equal to 370°C, m the presence of regenerated catalyst m a second user leactor operating under catalytic crackmg conditions including WHSV of 75-275 hr ' . catalyst to oil ratio of 4-12 and nsei top temperature of 425 - 525°C to obtam second ciacked hydrocarbon products,

iv ) sepaiating the catalytically cracked products from the second nsei ieactoi alongwith the cracked products compnsmg hydrocarbons having boilmg points less than or equal to 370°C. from the first nsei leactor m a main fractionating column to vield ciacked products compnsmg dry gas, LPG gasolme. middle distillates, heavy cycle oil and slurry oil.

v ) recycling the entire heavy cycle oil compnsmg hydrocaibons having boilmg pomts m the range of 3"0°C to 450°C and full or part of the slurry oil having boilmg points greater than oi equal to 450°C mto the second riser reactoi at a vertically displaced position lowei than the position of introduction of the mam feed comprising bottom unconverted hydrocarbon fraction havmg boilmg points greater than or equal to 3^"O°C from the first riser ieactoi to obtam middle distillate products compnsmg hydrocarbons having carbon atoms in the range of C₈ - C₂₄ ranging from about 50 to 65 wt % of the feed stock.

v ) optionally, recyclmg the fraction of unconverted hydrocarbons with boilmg points greater than or equal to 370°C. obtamed in step (iv) in nser reactors by repeating steps (in) to (iv) to obtam substantially pure middle distillate products

A process as claimed in claim 1 wherem, the feed stock is selected from petroleum based heavy feed stock such as vacuum gas oil (VGO), visbreaker/cokei heavy gas oil. coker fuel oil and hydrociacker bottom, etc

A process as claimed in claim 1 wherem the feed stock is preheated at a tteemmppeerraattuuiiee mm tthhee rraannggee ooff 11550-350°C and then injected to pneumatic flow user t e crackmg reactor

A process as claimed in claim 1 wherem. the mixed catalyst is obtamed from an intermediate vessel that mixes the spent catalyst from the common stnpper oi preferably first stnpper with the regenerated catalyst from the common regenerator and charges the mixed catalyst with coke content in the range of about 0 2 to 0 8 wt% of catalyst to the bottom of the fust nsei at a tempeiature of 450 - 575°C

A process as claimed in claim 1 wherem the ciacked hvdiocarbon vapoi products from the fust and second risers are quickly separated from iespective spent catalysts using separating devices to minimize the over crackmg of middle distillate range products mto undesirable lightei hvdiocaibons

A process as claimed in claim 1 wherem the spent catalysts from the first and second nser reactois are passed through respective dedicated catalyst stnppers or a common stnpper to render the catalysts substantially free θf from entramed hydrocarbons

A process as claimed m claim 1 wherem the regenerated catalyst with coke content of less than 0 4 wt% is obtained by burning a portion of the spent catalyst from the first stnpper. the spent catalyst from the second stnppei or the common stnpper in a turbulent or fast fluidized bed legeneratoi m the presence of an or oxygen containing gases at a temperature in the range of about 600°C to 750°C

A process as claimed m claim 1 wherem the cataly st between the fluidized bed nsei ieactors. stnppers and the common regenerator is continuously circulated through standpipe and slide alves

A process as claimed m claim 1 wherein the cntical catalytic crackmg conditions in the fust ieactor mcludmg mixed regenerated cataly st iesult in very high selectivity of middle distillate range products and conv ersion of hydrocarbon products of boilmg pomt less than or equal to 370°C at lowei than 50 wt% of the fresh feed.

A process as claimed m claim 1 wherem the catalyst compnses a mixture of commeicial ReL SY zeolite based catalyst having fresh surface aiea of 1 10- 180 πr/gm . pore volume of 0 25-0 38 cc gm and average particle size of 60-70 micron along with selective acidic bottom upgrading components m the range of about 0-10 wf%

A process as claimed in claim 1 wherem the unconverted heavy hydrocarbon fraction from second riser lecycled mto the second user ranges from about 0-50 wt% of the mam feed rate to the second nser. depending on the nature of the feedstock and operatmg conditions kept m the nsers

A process as claimed m claim 1 wherem amount of steam foi feed dispeision and atomization. catalyst lifting at the nser bottom in the first and the second nser reactors is in the range of 1-20 wt% of the respective total hydrocaibon feed dependmg on the quality of the feedstock

A process as claimed m claim 1 wherein the spent catalyst resides in the stnpper foi a period of upto 30 seconds

A process as claimed m claim 1 wherem the iegenerated catalyst entering at the bottom of the second nser reactor has coke of about 0 1-0 3 wt% at a temperatuie of about 600-750°C and is lifted by catalytically inert gases 43

A process as claimed m claim 1 wherem the combined Total Cycle Oil ( 150-370°C) product which is a mixture of Heavy naphtha (150-216°C) and Light c cle oil (216- 370°C) has higher cetane number than that from conventional distillate mode FCC unit and other properties such as specific giavity viscosity pour point etc aie in the same range as that of commeicial distillate mode FCC unit

A process as claimed m claim 1 wherem changing the cut pomt of the TCO from the first user to 120-370°C, processmg 370°C- part of the first nsei product m the second riser, and changing the cut pomt of TCO from second nsei to 120-390°C. the yield overall combmed TCO product mcreases by 8- 10 wt% and the combmed TCO product has the same properties but improved cetane number as that of TCO from commercial distillate mode FCC umt

A process as claimed in claim 1 wherem the Total Cycle Oil compnses a mixture of heavy naphtha hydrocarbons havmg boilmg pomts from about 150°C to 216°C and light cycle oil hydrocarbons havmg boilmg points from about 216 °C to 370 °C

A fluidized bed catalytic crackmg system for the production of high yield of middle distillate products compnsmg hydrocarbons having carbon atoms m the iange of C₈ to C₂₄ from heavy petroleum feeds, by a process as defined in claim 1 said system compnsmg at least two nser reactors ( 1 and 2) wheiem. a fresh feed is introduced into the first nser reactor ( 1), typically at the bottom section above regenerated catalyst entry zone through a feed nozzle (3), and at the end of the first nser reactoi ( 1), the spent catalyst is quickly separated from h drocarbon product vapors usmg separating devices (4) and subjected to multistage steam stnppmg to remove any entramed hydrocarbons, and a conduit (5) feeds a part of the said stripped catalyst mto a regeneratmg apparams (7) and the other part of the stripped catalyst from the conduit (5) travels through another conduit (6) mto a mixing vessel ( 10). and thereafter, the mixed catalyst from the mixing vessel ( 10) travels through a conduit ( 19) and is fed to the bottom of the first nser reactor (1), the hydrocarbon product vapois from the first user reactor ( 1) which are separated from the catalvst in the sepaiating devices (4) aie fed to a vacuum or atmospheiic distillation column ( 13) through conduit (12) whereby the first cracked hydrocaibon products aie sepaiated mto a first fraction compnsmg hydrocarbons having boilmg points less than or equal to 370°C and a second fraction compnsmg uncracked hydrocarbons with boilmg pomts greater than or equal to 370°C, the said second fraction compnsmg uncracked hydrocarbon products is fed through feed nozzle (16) mto the bottom of second nser reactor (2) above the regenerated catalyst entry zone, and the regenerated catalyst from the regenerating apparatus (7) is fed to the bottom of the second nser reactor (2) through a conduit (9). and subsequently the hydrocarbon products of the second nsei reactor (2) are separated from the catalyst m separating devices (11), and the cracked products of the second nser reactor (2) along with the products of the first fraction of the first nser reactor ( 1) compnsmg hydrocarbons with boilmg pomts less than oi equal to 370°C are fed to a mam fractionator column (15) which separates the said products mto dry gas. LPG. gasolme. heavy naphtha, light cycle oil. heavy cycle oil. and slurry oil. and the entire heavy cycle oil and full or part of the slurry oil consisting mamly of hydrocarbons with boilmg points greater than or equal to 370°C are iecycled back to the second nser reactor (2) through a separate feed nozzle ( 17) located at a point lower than the position of introduction of mam feed, and the feed and cracked product vapors travel along with the catalyst mto the reactor wherem the spent catalyst separated from product vapors of the second nser reactor (2) m separating devices and the spent catalyst is subjected to multistage steam stnppmg for removal of entrained hy diocarbons and the stnpped catalyst travels through a conduit ( 18) mto the iegeneiatmg apparatus (7). wherem the coke on catalyst is burnt m the piesence of air and/or oxygen containing gases at high temperatuie. and the flue gas from regeneration is separated from the entramed catalyst fines in separatmg devices (23) and the flue gas leaves from top of the regeneiatmg apparatus (7) through a conduit (22) for heat iecovery and venting through stack, the hot regenerated catalyst is withdrawn from the regeneratmg apparatus (7) and divided mto two parts, one going to the mixing vessel (10) through the conduit (8) and the other directly to the bottom of the second nser reactor (2), and the mixed catalyst from the mixing vessel (10) is fed through the conduit (19) to the inlet of the first user reactor (1), controlling the catalyst bed level m the individual oi common stnpper, the catalyst circulation rate from the common regenerator and the quantity of the spent and tegenerated catalyst entering into the mixing vessel (10) usmg slide valves placed on the conduits and thereby producmg high yield of middle distillate products

A system as claimed m claim 1 wherem the separatmg device includes cyclone separator

A process as claimed in claim 1 wherein pressure in the first and second user reactors aie in the range of 1 0 to 4 0 kg/cm^"(g)