CA1198129A - Thermal regenerative cracking (trc) process and apparatus - Google Patents

Abstract

ABSTRACT OF THE DISCLOSURE

An improved Thermal Regenerative Cracking (TRC) apparatus and process includes: (1) an improved low residence time solid-gas separation device and system; and (2) an improved solids feeding device and system; as well as an improved sequential thermal cracking process; an improved solids quench boiler and process; an improved preheat vaporization system;
and an improved fuel gas generation system for solids heated.
One or more of the improvements may be incorporated in a conventional TRC system.

Description

z~

This application is a d.ivision oE co-pending application Serial No. 361,734 filed September 30, 1980.
The present invention relates to improvernents in Thermal Regenerative Cracking (TRC) apparatus and process, as described in U.S. Letters Patent Nos. 4,061,562 and 4,097,363 to McKinney et al.
According -to one broad aspect, the presen-t invention relates to a TRC process wherein the telnperatUre in the cracking zone is between 1300 and 2500F. and wherein the hydrocarbon feed or the hydrosulfurization residual oil along with the entrained inert solids and the diluent gas are passed through the cracking zone for a residence time of 0.05 to 2 seconds, the improvelnents comprising the process o.r generating uel oil and removing coke deposits on said solids comprising the steps of generating a fuel gas from fuel and air;
delivering the fuel gas to a transfer line; mixing the particulate solids with the fuel gas in the transfer line to elevate the temperature of the solids; and combusting the fuel gas in the transfer line to elevate the temperature of the solids and remove the coke from the solids; and the process for separating by centrifugal force particulate solids from the dilute mixed phase stream of gas and solids comprising the steps of adding the mixed phase stream to a chamber having a flow path of essentially rectangular cross section from an inlet of inside diameter Di disposed normal to the flow path, said flow path having a height H equal to D1 or 4 inches, whichever is greater, and a width W greater than or equal to 0.75 Di bu~ less than or equal to 1.25 Di, disengaging solids from gas by centriEugal force within said chamber along a bed of solids found at a wall opposite to the inlet as the gas flows through said flow path, the gas changing direction 180, ~8~

and the solids being projected 90 toward a solids outlet, wi-thdrawing the gaseous portion o:E the inlet stream from a gas outlet, disposed 180 from the inlet, the gas portion containing about 20~ residual solids, said gas outlet located between -the solids outlet and inlet, the gas outlet being at a distance no grea-ter than ~ Di from the inlet as measured between respective centerlines, and withdrawing the solids by gravity through the solids outlet.
According to another broad aspect, the present invention relates to a TRC apparatus wherein the temperature in the reaction chamber is between 130a and 2500F and wherein the hydrocarbon fluid feed or the hydrosulfurization residual fluid feed oil along with the entrained inert solids and the diluent gas are passed through the reaction chamber for a residence time of 0.05 to 2 seconds, the improvements wherein the apparatus for admixing the inert solids rapidly and intimately with the fluid feed introduced simultaneously thereto comprises an upper reservoir containing the particulate solids; a conduit extending downwardly from the reservoir to the reaction chamber, said conduit being in open communication with the reservoir and reaction chamber: and a solids-gas separator designed to efEect rapid removal of particulate solids from a dilute mixed phase s-tream o~ solids and gas, said separator comprising a chamber for disengaging solids from the incoming mixed phase stream, said chamber having rectilinear or slightly arcuate longitudinal walls to form a flow path essentially .rec-tangular in cross section, said chamber also having a mixed phase inlet, a gas phase outlet, and a solids phase outlet, with the inlet at one end of the chamber disposed normal to the flow path, the solids outlet at the other end of the chamber, said solids outlet suitable for downflow of -2a-.~

discharged solids by gravity, and the gas outlet therebetween orien-ted to e.~fect a 180 change in direction o~ the gas.
According to yet another broad aspect, the present invention relates to a TRC apparatus wherein the temperature in -the reaction chamber cracking ~one is between 1300 and 2500F
and wherein the hydrocarbon fluid feed or the hydrosul~urized resiaual oil fluid ~eed along with the entrained inert par-ticulate solids and the diluent gas are passed through the cracking zone for a residence time of 0.05 to 2 seconds, the improvement comprising a solids-gas separator designed to effect rapid removal of particulate solids from a dilute m:ixed phase stream of solids and gas, said separator comprising a chamber for disengaging solids from the incoming mixed phase stream, said chamber hav.ing rectilinear or slightly arcuate longitudinal ~alls to form a flow path essentially rectangular in cross section, said chamber also having a mixed phase inlet, a gas phase outlet, and a solids phase outlet, with the inle~
at one end of the chamber disposed normal to the flow path, the solids outlet at the other end of the chamber, said solids outlet suitable for down~low of discharged solids by gra~ity, and the gas ou-tlet therebetween oriented to effec-t a 1~0 change in direction of the gas.
According to a still further broad aspect, the present invention relates to a TRC apparatus wherein the temperature in the reaction chamber is between 1300 and 2500F and wherein the hydrocarbon feed or the hydrosulfurized residual oil along with the en-trained inert particula-te solids and the diluent ~as are passed through the reaction zone for a residence time of 0.05 to 2 seconds, the improvement comprising a solids-gas separation system to separate a dilute mixed phase stream of gas and particula-te solids into an essen-tially solids free gas -2b-stream, the separa-tion system comprising a chamber for rapidly disen~aging about 80% of the par-ticula-te solids from the incoming dilute mixed phase stream, said chamber having approximately rectilinear or slightly arcuate longitudinal side walls to form a flow path of heiyht H and width W approxima-tely rectangular in cross section, said chamber also having a mixed p~ase inlet of inside width Di, a gas outlet, and a solids outlet, said inlet at one end of the chamber disposed normal to -t~e flow path whose height H is equal to at least Di or 4 inches, whichever i5 greater and whose width W is no less than 0.75 Di but no more than 1~25 Di, said solids outlet at the opposite end of the chamber and being suitable for downflow oF
discharged solids by gravity, and said gas outlet therebetween at a distance ~o greater than ~ Di from the inlet as measured between respective centerlines and oriented -to efEect a 180 change in direction of the gas whereby resultant centrifugal forces direct the solid particles in the incoming stream toward a wall o~ the chamber opposite to the inl~t forming thereat and maintaining an essentially static bed of solids, the surEace of the bed defining a curvilinear pa-th o~ approximately 90~ for the outElow of solids to the solids outlet, a secondary solids--gas separator, said secondary separator removing essen-tially all o~ the residual solids, a first conduit connecting the gas outlet from the charnber to the secondary separator, a vessel for the discharge of solids, a second conduit connecting said vessel and the chamber, and pressure balance means to maintain a height of solids in said second conduit -to provide a positive seal between the chamber and vessel.
According to at least one further broad aspect, the present invention relates to a TRC process wherein -the -2c-1, ~ ~, ~8~

temperature in the reac-tion chamber is between 1300 amd 2500F
and wherein the hydroc~rbon fluid feed or the hydrosulfurization residual oil along with the entrained inert solids and the diluent gas are passed through the reaction chamber for a residence time of 0.05 to 2 secondsl the improvemen-t comprising a rnethod Eor separating by centrifugal Eorce particulate solids from a dilute mixed phase s-tream oE
gas and solids, the method comprising the steps of adding the mixed phase stream to a chamber having a flow path of ess~ntially rectangular cross section from an inlet of inside diameter Di disposed normal to the flow path, said flow path having a height H equal to Di or 4 inches, whichever is greater, and a width W grea-ter than or equal to 0.75 Di but less than or equal to 1.25 Di, disengaging solids from gas by centrifugal force within said chamber along a bed of solids found at a wall opposite to the inlet as the gas flows through said flow path, the gas changing direction 180, and the solids being projected 90 toward a solids outlet, withdrawing -the ga~eOUS portion of the inlet s-tream from a gas outlet, disposed 180 from the inlet, the gas portion containing about 20%
residual solids, said gas outlet located between the solids outlet and inlet, the gas outlet being at a distance no greater than 4 Di frorn the inlet as measured between respective cen-terlines and withdrawing the solids by gravity through the solids outlet.
The invention will be more fully appreciated by reference to the accompanying drawings in w~ich:

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Figure 1 is a schematic diagram of a TRC system and process according to -the prior art.
Figure 2 is a schematic diagram of the fuel gas generation system and process of the subject inven-tion.
Figure 3 is an alternative embodiment wherein the fuel gas is burned to fuel gas to provide additional heat for tl-e particulate solids.
Figure 4 is a cross-sectional elevational view of the solids feeding device and system as applied to tubular reactors and for use with gaseous feeds.
Figure S is an enlarged view of the intersection of the solid and gas phases within the mixing zone of the reaction chamber.
Figure 6 is a top view of the preferred plate geometry, said plate serving as the base of the gas distributlon chamber.
Figure 7 is a graph of -the relationship between bed density, pressure drop, bed height and aera-tion gas velocity in a fluidized bed.
Figure 8 is a view through line 8-8 of Figure 5.
Figure 9 is an isometric view of the plug which extends into the mixing zone to reduce flow area.
Figure 10 is an alternative preferred embodiment of the control features of the present invention.
Figure 11 is a view along line 11-11 of Figure 10 showing the header and piping arrangements supplying aeration gas to the clean out and fluidization nozzles.
Figure 12 is an alternate embodiment of the preferred invention wherein a second feed gas is contemplated.

(S&W) 1 FIGURE 13 is a view of the apparatus of FIGURE 12 through line 13-13 of FIGURE 12.
3 FIGURE 14 is a schematic diagram of the sequential 4 thermal cracking process and system of the present invention~
S FIGURE 15 is a schematic flow diagram of the 6 separation system of the present invention as ap~ended to a 7 typical tu~ular reactor.
8 FIGURE 16 is a cross sectional elevational view of g the preferred embodiment of the separator~
FIGURE 17 is a cutaway view throu~h section 1~-17 11 of FIGURE 16.
12 FIGURE 1~ is a cutaway view through section 18-18 13 of FIGURE 16 showing an alternate geometric configuration oE the 14 separator shell.
FIGURE 19 is a sketch of the separation device of 16 the present invention indicating gas and solids phase flow 17 patterns in a separator not having a weir.
18 FIGURE 20 is a sketch of an alternate embodiment 19 of the separation device having a weir and an extended separatlon chamber.
21 FIGURE 21 is a sketch of an alternate er~odiment of 22 the separation device wherein a stepped solids outlet is employ~d, 23 said outlet having a section collinear with the flow path as well as a gravity flow section.
FIGU~E 22 is a variation of the embodiment of FIGVRE
26 21 in which the solids outlet of FIGURE 20 is used, but is not 27 stepped.
28 FIGURE 23 is a sketch of a variation of the 29 separation device of FIGURE 8 wherein a venturi restriction is incorporated in the collinear section of the solids outlet.

(S&W) 69Ç-147 1 FIGURE 24 is a variation of the embodimen~ of

2 FIGURE 23 oriented for use with a riser type reactor~

3 FIGURE 25 is a sectional elevational view of the

4 solids quench boiler using the quench riser;
S FIGURE 26 ls a detailed cross sectional elevational ~ view of the quench exchanger of the system;
7 FIGURE 27 is a cross sectional plan view taken 8 through line 27-27 of FIGURE 26;
9 FIGURE 28 is a detailed drawlng of the reactor outlet and fluld bed quench riser particle entry area.
11 FI~URE 29 is a schema~ic diagram of the system of 12 the invention for vaporizing heavy oil.

16 DESCRIPTION OF THE PREFERRED E~1BODIME~TS

18 The improvements of the subject invention are e~hodied in the environment of a thermal regeneration cracking reactor ~TRC) which is illustrated in FIGURE 1.

2~ , Referring to FIGURE 1, in the prior art TRC process~
~3 and system, thermal cracker feed o.il or residual oil, with or 24 I'without hlended distillate heavy gas, entering throu~h line 10 ~25 !`and hydrogen entering through line 12 pass through hydrodesulfurized 26 zone 14. ~ydrosulfurization effluent passes through line 16 an~
27 enters flash chamber 19 from which hydrogen and contamina~ing gases 28 ~including hydrogen sulfi.de and ammonia are removed overhead through 29 ~line 20, while flash liquid is removed through line 22. The flash . . :
S&W-1 : liquid passes through preheater 24, is adrnixed with dilutlon 2 l steam entering tllrough line 26 and then flo~s to the bottom 3 1 of thermal cracking reactor 28 through line 30O
4 1l A stream of hot regenerated solids is charged l throuyh line 32 and admixed with steam or other fluidizing gas 6 .~ entering through line 34 prior to entering the bottom of riser 7 ,l 28. The oil, steam and hot solids pass in entrained flow up-B , wardly through riser 28 and are discharged through a curved 9 seg~ent 36 at the ~op of the riser to induce centrifugal separ-atlon of solids from the effluent stream. A stream containing 11 most of ~he solids passes ~hrough riser discharge segment 38 and 12 can be mixed, if desired, with make-up solids entering through 13 ~ line 40 before or after enlering solids separator-stripper 42.
14 Another stream containing most of the cracked product is dis-~5 charged axially through conduit 44 and can be cooled by means of 16 a quench stream entering through line 46 in advance of solids 17 separator-stripper ~8.
18 " Stripper steam is charged to solids separators 42 19 `I and 48 through lines 50 and 52, respectively. Product streams 1 are removed from solids separators 42 and 48 through lines 54 21 ~ and 56, respectively, and then combined in line 58 for passage ~2 ' to a secondary quench and product recovery trainj not shown.
23 Coke-laden solids are remo~ed from so~ids separators 42 and ~8 24 through lines 60 and 62, respectively, and cornbined in line 64 for passage to coke burner 66. If required, torch oil can be 26 added to burner 66 through line 68 while stripping steam may be 27 ~~ added through line 70 to strip combustion gases from the heated 2B 1I solids. Air is charged to the burner through line 69. Com~ustlon 29 I gases are removed from the burner through line 72 for passage I to heat. and energy recovery systems, not shown, while regenerated S&W
6~147 1 ~ hot solids wh.ich are relatively free of coke are removed from 2 the burner tnrough line 32 for recycle to riser ~8. In order 3 to produce a cracked product containing ethylene and molecular 4 hydrogen, petroleum residual oil is passed through the catalytic I hydrodesulfurized zone in the~resence of hydrogen at a tem-6 perature between 650F and 900F, with the hydrogen being 7 chemlcally combined with theoil during the hydrocycling step.
8 The hydrosulfurization residual oil passes through the thermal 9 cracking zone together with the entrained inert hot solids . functioning as the heat source and a diluent gas at a temperature 11 l, between about 1300F and 2500F for a residual time between 12 ~ about 0.05 to ~ seconds to produce the cracked product and 13 ethylene and hydrogen. For the production of ethylene ~y 14 thermally cracking a hydrogen feed at least 90 volume percent ~ of which comprises light gas oil fraction of a crude oil 1~ boiling between 400F and 650F, the hydrocarbon feed, along 17 1 with diluent gas and entrained inert hot gases are passed 18 i through the cracking zone at a temperature between 1300F and 19 ~500F for a residence time of 0.05 to 2 seconds. The weight I ratio of oil gas to fuel oil is at least 0.3, while the cracking 21 ! severity corresponds to a methane yield of at least 12 weight 22 percent based on said feed oil. Quench cooling of the produc~
immediately upon leaving the cracked zone to a temperature ~4 below 1300F ensures that the ethylene yield is greater than 25 : the methane yield on a weight basis.

~9 ~

_ _ / _ __ S&W ~ 2~
69~ 7 1 (a) Improved Fuel Gas Generation For Solids 2 Heatins.

4 FIGURE 2 illustrates the improved process and system cf the invention as may be e~bodied in a prior art TRC system, in lieu 6 ,of the coke burner 66 (FIG~ 1). Particulate solids and hyd ~ arbon feed gas 7 ,'enter a tubular reactor 13A throuah ~nes llA and 12A respec~vely. Ihe cracked 8 1 effl~lent from the tubular reactor 13A is separated from the pa~c~ate solids 9 in a separator 14A and quenched in line 15A by quench material injected from line 17A. The solids separated fxom the effluent 11 are delivered through line 16A to a solids separator. The residual 12 solids are removed from the quenched product gas in a secondary 13 separator 18A and delivered to the solid stripper 22A. The solids-14 ~Ifree product ~as is taken overhead from the secondary separator ~,18A through line l9A.

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s~w 1 I The particulate solids in the solid stripper 22A, 2 having delivered heat during the thermal cracking in the tubular 3 I reactor 13A, must be reheated and returned to the tubular reactor 4 1 13A to continue the cracking process.
' The particulate solids prior to being reheated~ are 6 ~stripped of gas in the solid stripper 27A by steam delivered to 7 l the solid stripper 22A through line 23A.
8 After the particulate solids have been stripped of g gas impurities in the solid s-tripper 22A, the particulates solids ~t ' Y 10 are at a temperature of about 1,450F.
11 ~ The fuel gas generation apparatus of the inven~ior.
12 ~ consists of a combustion vessel 30A, and pre-heat equipment for 13 1 fuel/ air (or 2) and steam which are ~elivered to the combustion 14 , vessel 30A. Pre-heaters 32A, 34A, and 36 are shown in fuel llne ! 38A, air line 40A, and steam line 42A respectively, i6 , 'l'he system also includes a transfer line 44A intc 17 I which the combusted fuel gas from the combustion vessel 30A ar.d 18 the stripped particulate solids from the solid stripper 22A aLe 19 , mixed to heat and decoke the particulate solids. The transfef ,~ line 44A is sized to afford sufficient residence time for the 21 1 steam emanating from the combustion vessel 30A to decompose by 22 !, the reaction with carbon in the presence of hydrogen and to remove the net carbon from the solids-gas mixture. In the preferred 24 , embodiment the transfer line 44 will be about 100 feet long~ A
,25 1 line 26A is provided for pneumatic transport gas if necessary~ ;
26 A separator, such as a c~clone separa~or 46~ is 27 I provided to separate the heated decoked particulate solids from 28 ¦I the fuel gas. The particulate solids from the separator 46A are 29 ,~ returned through line 48A to the hot solids hold vessel 27A
',~ and the fuel gas is taken overhead through line 50A.

., il ' _g_ S&W ~ 2~

1 In the process, fuel, air and steam are delivered 2 through lines 38A, 40A and 42A respectively to the combustion vessel 30A and combusted therein to a temperature of about 4 2,300F. to produce a fuel gas having a high ratio of CO to CO2 S and at least an equivalent molal ratio of H20 to H2. The H20 6 to H2 ratio of the fuel gas leaving the combustion vessel 30A
7 ; is above the ratio required to decompose steam by reaction witn 8 carbon in the presence of hydrogen and to insure that the net 9 carbon in the fuel gas-particulate solids will mix will be removed before reaching the separator 46A.
11 ~he fuel gas from combustion vessel 30A at a 12 ; temperature of about 2,300F. is mixed in the tubular vessel 44A
13 l~ with stripped particulate solids having a temperature of about 14 1,450F. The particulate solids and fuel gas rapidly reach an equilibrium temperature of 1,780F. and continue to pass through 16 the tubular vessel 44A. During the passage through the tubulac 17 vessel 44A the particulate solid-fuel gas mixture provide the 18 ~' heat necessary to react the net coke in the mixture with steam.
19 ll As a result, the particulate solid-fuel gas mixture is cooled by about 30F. i.e., from 1,780F. to 1,750F.
21 I The particulate solid-fuel gas mixture is separated 22 in the separator 4ÇA and the fuel gas is taken at 1,750~F. through line 50A. The particulate solids are delivered to the hot soli~s 24 l, hold vessel 27A at 1,750F. and then to the tubular reactor 13A.
,25 In the alternative embodiment of the invention 26 1 illustrated in FIGURE 3, only fuel and air are delivered to the 27 combustor 30 and burned to a temperature of about 2,300F. to 28 ' provide a fuel sas. The fuel gas at 2,300F. and particulate 29 1I solids at about 1,450~. are mixed in the transfer line 44A to a temperature of about 1,486F. Thereafter air is delivered I to the transfer line 44A throug~ a line 54A. The fuel gas in ! .

S&W~

1 ~he line 94A is burned to elevate the temperature of the partlcu-2 ~late solids to about 1,750F. The resultant flue gas is separated 3 Ifrom the hot solids in the separator 46A and discharged through 4 I the line 52A. The hot particulate solids are returned to ~he ! system to provide reaction heat.
6 'I An example of the system and process of FIG~RE 3 7 ~ follows. 7,000 pounds per hour of fuel pre-heated to 600F. in 3 the preheater 32A and 13 MM SCFD of air heated to 1,000F. are 9 ;burned in the combustor 30A to 2,300F. to produce 15.6 M~l SCFD
' of fuel ~as.
11 The 15 MM SCFD of fuel ~as at 2,300F. is mixed in 12 the transfer line 44A with 1 ~ pounds per hour of stripped pdr~iCU-i3 late solids from the solids stripper 22A. The particulate sollds 4 ;have 1,600 pounds per hour of carbon deposited thereon~ The com-i5 posite fuel gas-particulate solids gas mixture reaches an 16 equilibrium temperature of 1,480F. at 5 psig in about 5 milll-17 ,'seconds. Thereafter, 13 ~l SCFD of air is delivered to ~he 18 'transfer line 44A and the 15.6 ~M SCFD of fuel gas is burned 19 !with the air to elevate the solids temperature to 1,750F. and buxn the 1,600 pounds per hour of carbon from the particulate 21 ,solids.
22 The combusted ~as from the transfer line 44A is 23 separated from the solids in the separator 46A and discharged 24 as flue gas.

%9 (b) Improved Solids Feeding Device and System~

Again referring to Figure 4 in lieu of the system of the prior art (See Figure 1) wherein the stream of solids plus Eluidizing gas contact the flash liquid-dilution steam mixture entering reactor 28~ structurally the apparatus 32B of the subject invention comprises a solids reservoir vessel 33B and a housing 34B for the internal elements described belowO The housin~ 34B is conically shaped in the embodiment of Figure 4 and serves as a transition spool piece between the reservoir 33B and the reactor 32B to which it is flangeably connected via flanges 35B~ 36B, 37B and 38Bo The particular geometry of the housing is functional rather than critica].~ The housing is itself comprised of an outer metallic shell 39B~ preferably of steel ? and an inner core 40B of a castable ceramic material. It is convenient that the material of the core 40B
forms the base 41B of the reservoir 33B~ /
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~96-147 Set into and supported by the inner core 40B is a 2 I gas distribution chamber 42B, said chamber being supplied with 3 gaseous feed from a header 43B. While the chamber 42B may be 4 ' of unitary construction, it is preferred that the base separating the chamber 42B from reaction zone 44B be a removable plate 45B.
6 One or more conduits 46B extend downwardly from the reservoir 7 33B to the reaction zone 44B, passing through the hase 41B, 8 and the chamber 42B. The conduits 46B are in open communication 9 with both the reservoir 33B and the reaction zone 44B providing thereby a path for the flow of solids from the reservoir 33B ~D
11 ' the reaction zone 44B. The conduits 46B are supported by the 12 material of the core 40B, and terminate coplanarly with a pla~e 13 45B, which has apertures 47B to receive the conduits 46B. The 14 region immediately below the plate 45B is hereinafter referred to as a mixing zone 5~ which is also part of the reaction zone 16 44.
17 As shown in FIGURE 5, an enlarged partial view of ~he 18 intersection of the conduit 46~ and the plate 45B, the apertures 19 47B are larger than the outside dimension of conduits 46B, for~ning ,therebetween annular orifices 48B for the passage of gaseous ~eed ~1 from the chamber 42B. Edges 49B of the apertures 47B are pre--22 ferably convergently beveled, as are the edges 50B, at the tip of the conduit wall 51B. In this way the gaseous stream from the charnber 4~B is angularly injected into the mixing zone53B
1l and intercepts thesolids phase flowing from conduits 46B. A
26 Iprojectlon of the gas flow would form a cone shown by dotted lines 27 1l 52B the vertex of which is beneath the flow path of the solids.
28 IBy introducing the gas phase angularly, the two phases are mixed 29 I rapidly and uniformly, and form a homogeneous reaction phase.
The ~.~ixing of a solid phase with a gaseous phase is a function of 1~
., ,, 2~3 the shear surface between the solids and gas phases, and the flow area. A ratio of shear surface to flow area ~S/A) of infinity defines perfect mixing; poorest mixing occurs when the solids are introduced at the wall of the reaction zone. In the system of the present invention, the gas stream is introduced annularly to the solids which ensures high shear surface. By also adding the gas phase transversely through an annular feed means, as in the preferred embodiment, penetration of the phases is obtained and even faster mixing results. By using a plurality of annular gas feed points and a plurality of solid feed conduits, even greater mixing is more rapidly promoted, since the surface to area ratio for a constant solids flow area is increased. Mixing is also a known function o~ the L/D of the mixing zone. A plug creates an effectively reduced diameter D in a constant L, thus increasing mixing.
The Plug 54B, which extends downwardly from plate 45B, as shown in Figures 4 and 5, reduces the flow area, and forms discrete mixing zones 53B. The combination of annular gas addition around each solids feed point and a confined discrete mixing zone greatly enhances the conditions for mixing. Using this preferred embodiment, the time required to obtain an essentially homogeneous reaction phase in the reaction zone 44B
is quite low. Thus, this preferred method of gas and solids addition can be used in reaction systems having a residence time below 1 second, and even below 100 milliseconds. In such reactions the mixing step must be performed in a fraction of the total rssidence time, generally under 20% thereofO If this criteria is not achieved, localized and uncontrolled reaction occurs which deleteriously affects the product yield and distribution. This is caused by the maldistribution of solids Y~`i' 2~
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1 I normal to the flow through the reaction zone 44B thereby creating 2 temperature and or concentration gradients therein.
3 The flow area is further reduced b~ placing the 4 1 apertures 47s as close to the walls of the mixing zone 53B as ;i possible. FIGURE 6 shows the top view of plate 45B having in 6 complete circular aper~ures 47B symmetrically spaced along the 7 circumference, The plug 54B, shown by the dotted lines and 8 in FIGURE g, is below the plate, and establishes the discrete 9 ; mixing zones 53B described above. In this embodiment~ the apertures 47B are completed by the side walls 55B of gas 11 distribution chamber 42B as shown in FI~URE 5. In order to 12 prevent movement of conduits 46B by vibration and to retain t~e 13 uniform width of the annular orifices 48B, spacers 56B, are 14 1 used as shown in FIGURE ~. However, the conduits 46B are pr marily supported within the housing 3~B by the material of the 16 core 40B as stated above.
17 Referring to FIGURE 9, the plug 54B serves to 13 reduce the flow area and define discrete mixing zones 53B~
i9 ~he plug 54B may also be convergently -tapered so that there is a gradual increase in the flow area of the mixing zone 53s 21 until the mixing zone merges with remainder of the reaction 22 1 zone 44B. Alternatively, a plurality of plugs 54B can be use~
23 Ij to obtain a mixing zone 53B of the desired geometric con-24 figuration~
25 I Referring again to FIG~RE 4, the housing 34B may ~6 preferably contain a neck portion 57B with corresponding lining 27 1 58B of the castable ceramic material and a flange 37B to cooperate 28 with a flange 38B on the reaction chamber 31B to mount the neck 29 ~I portion 57B~ This neck portion 57B defines mixing zone 53B, ' Il . , ;! -15-and allows complete removal of the housing 34B without dis-2 i assembly of the reactor 313 or the solids reservoir 33B~ Thus, 3 1 installation, removal and maintenance can be accomplished 4 ' easily. Ceramic linings 60B and 62B on the reservoir 33B
S ~ and the reactor walls 613 respectively are provided to prevent 6 erosion.
7 , The solids in reservoir 33B are not fluidized 8 ; except solids 63B in the vicinity of conduits 46B. Aeration 9 gas to locally fluidize the solids 63B is supplied by nozzles 64B syrnrnetrically placed around the conduits 46B. Gas to 11 I nozzles 64Bis supplied by a header 65B, Preferably, the header 12 65Bis set within the castable material of the core 40B, but 13 this is depend~nt on whether there is sufficient space in the 14 housing 34B~ A large mesh screen 66B is placed over the inlets ~ of ~he conduit 64B to prevent debris and large particles from 16 entering the reaction zone 4~B or blocking the passage of the 17 particulate solids through the conduits 46B.
18 By locally fluidizing the solids 63B, the solids 19 ~63B assume the characteristics of a fluid, and will flow thrGugh j the cond~lits 96B. The conduits 46B have a fixed cross sectional 21 1l area, and serve as orifices having a specific response to a 22 ,change in orifice pressure drop. Generally, the flow of 23 ; fluidized solids through an orifice is a function of the pressilre 24 drop through the orifice. That orifice pressure dropl in turn, 'isa function of bed height, bed density~ and system pressure~
26 , However, in the process and apparatus of this 27 , invention the bul)c of the solids in reservoir 33B are not 28 I fluidized. Thus, static pressure changes caused by varia~ions 29 in ~ed height are only 510wly communicated to the inlet of the coriduit 46B. Also the bed density remains approximately cons-tant S&~ g ,1 ntil the polnt of incipient fluldization is reached, that is, 2 point "a" of FIGURE 7. In the present invention, however, it 3 l is essential that the amount of aeration gas be below that 4 , amount. Any aeration gas flow above that at point "a" on S 1 ~IGUR~ 7 will effectively provide a fluidized bed and thereby 6 lose the benefits of this invention. By adjustment of the 7 aeration gas flow rate, the pressure drop acroCs the non-8 ; fluidized bed can be varied. Accordingly, the pressure drop 9 ; across the orifice is re~ulated and the flow of solids thereby ~ regulated as shown in FIGURE 7. As gas flow rates below 11 incipient flui~ization, significant pressure increases 12 above the orifice can be obtained without fluidizing the bul~ of 13 the solids. Any effect which the bed height and the bed denslty 1~ variations have on mass flow are dampened considerably by the presence of the non-fluidized reservoir solids and are essentidlly 16 eliminated as a signiricant factor. Further the control provided 17 by this invention affords rapid response to changes in solids 18 mass flow regardless of the cause.
~ " , 19 Together with the rapid mixing features descr-bed above, the present invention offers an integrated system for 21 , feedin~ particulate solids to a reactor or vessel, especially ~2 to a TRC tubular reactor wherein very low reaction residence 23 ~ times are encountered.
24 ' FIGURES loandll depict an alternate preferred embodi I ment of the control features o~ the present invention. In this 26 embodiment the reservoir 33B extends downwardly into the core material ~OB to form a secondary or control reservoir 71B~ The ~8 ~I screen 66B is positioned over the entire control reservoir 71B.
~9 ' The aeration nozzles 64B project down~Jardly to fluidize ~ssentially ~i these solids 63B beneath the scree~ 66~o The bot~om 41B of the l!

2~

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696-1~7 , 1 ; reservoir 33B is again preferably formed of the same material 2 i as the core 40B.
3 A plurality of clean out nozzles 72B are preferably 4 , provided to allow for an intermittent aeration gas discharge which removes debris and large particles that may have accumulated 6 on the screen 66B. Porous stone filters 73B prevent solids from 7 entering the nozzles 72B. Headers 65B and 74B provide the gas 8 supply to nozzles 64B and 7 2B respectively.
9 . The conduits 46B communicate with the reservoir 71B
' through leading section 4 6'B, The leading sections 461B are 11 formed in a block 75B made of castable erosion resistent ceramic 12 material such as Carborundum Alfrax 201. The block 75Bis 13 ~ removable, and can be replaced if eroded~ The entrance 75B to 14 each section 46'B can be sloped to allow solids to enter more , easily. In addition to being erosion resistent, the block i6 75B provides greater longevity because erosion may occur without 17 . loss of the preset response function. Thus, even if the conduit 18 leading sections 46'B erode as depicted by dotted lines ~7B~
19 the remaining leading section 46'Bwill still provide a known ' orifice size and pressure drop response. The conduits 46B
21 l, are completed as before using erosion resistent metal tubes 22 !~51B, said tubes being set into core material 40B and affixed ~3 l~ to the block 75B.
24 I FIG~RE llis a plan view of FIGURE ~o along section 1 9-9 showing an arrangement for the no2zles 64B and 72B, and the 26 I heade~s 65B and 74B. Gas is supplied to the headers 65B and 74B.
~; through feed lines 79B and 80B respectively, which extend out~
~8 ~ beyond the shell 34B. It is not necessary ~hat the headers be 29 set into the material of the core 40B, although this is a convenience from the fabxication standpoint. Uniform flow !

S&W ~!
696-147 ', ., I
1 j distribution to each of the nozzles is ensured by the hydraullcs 2 of the nozzles themselves, and does not require other devices 3 such as an orifice or venturi. The gas supplied to feed lines 4 1 79B and 80B is regulated via valve means not shown.
~IGURES 12 and 13 show the pextinent parts of an 6 al~ernate embodiment of the invention wherein a second gas dis-7 tribution assembly for feed gas is contemplated. As in the other 8 embodiments, a gas distribution chamber 42B terminating in annular orifice 48B surrounds each solids delivery conduit 46B. Howe~er, ; rather than a common ~IJall between the chamber 47B and the cor,duit 46B, a second annulus 83B is formed between the chamber 42B
12 ; and the conduit 46B. ~7alls 81~ and 51B define the chambers 13 83B. Feed is introduced through both the annular opening 48B
14 in the chamber 42B and the annular opening 84B in the annulus 83B at an angle to the flow of solids from the conduits 46i3.
16 The angular entry of the feed gas to the mixing zone 53Bis 17 provided by beveled walls 49B and 85B, which define the openirgs 18 il 48B and beveled walls 50B and 89B which define the openings 19 'I84B. Gas is introduced to the annulus 83B through the header ~l86B~ the header being set into the core 40Bi~ convenient.
21 ,, FIGURE i2 is a plan view of the apparatus of FIG~lRE
22 1 13 through section 11-11 showing the conduit openings and the ~3 annular feed openings 48B and 84B. Gas i5 supplied through feed 2~ lines 87Band88Bto the headers 43B and 86B and ultimately , to the mixing zones through the annular openings. Uniform flow 26 from the chambers 42B and 83B is ensured by the annular orifices 27 j, 48B and 84B. Therefore, it is not essential that flow dis-28 1l~ tribution means such as venturis or orifices be included in 29 !~ the header 43B. The plug 54B is shaped symmetrically to define discrete mixing zones 53B.

,, --1 9--CLW 'I
696-1471, .1 .

; Mixing efficiency is also dependent upon the velocities 2 of the gas and solid phases. The solids flow throu~h the conduits 3 I`46B in dense phase flow at mass velocities from preferably 200 4 to 500 pounds/sq. ft.~sec) although mass velocities between 50 and l'1000 pounds/sq. ft./sec./ may be used depending on the character-6 istics of the solids used. The flow pattern of the solids in the 7 absence of gas is a slowly diverging cone. With the introduction 8 of the gas phase through the annular orifices 48B at velocities 9 , between 30 and 800 ft./sec., the solids develop a hyperbolic flow pattern which has a high degree of shear surface. Preferably, the 11 ,gas velocity through the orifices 48B is between 125 and 250 fto/
12 secO Higher velocities are not preferred because erosion is 13 accelerated; lower velocities are not preferred because the hyper-14 ;bolic shear surface is less developed.
~ The initial superficial velocity of the two phases in 16 the mixing zone 53B is preferably about 20 to 80 ft./sec., 17 'although this velocity changes rapidly in many reaction sys~ems~
18 'such as thermal cracking, as the gaseous reaction products are 19 formed. The actual average velocity through the mixing ~one 53B
l and the reaction zone ~4B is a process consideration, the velocity 21 j being a function of the allowed residence time therethroughO
22 ! By em~loying the solid feed device and method of tne 23 present inventions, the mixing length to diameter ratio necessary 24 Ito intimately mix the two phases is greatly reduced~ This ratio is used as an informal criteria ~hich defines good mixing. Gen-26 lerally, an L/D(lenyth/dia.~ ratio of from 10 to 40 is required~
27 IUsing the device disclosed herein, this ratio fs less than 5, wi~h 28 Iratios less than 1.0 being possible~ Well designed mixing devices 29 jlOf the present invention may even achieve essentially complete Imixing at L/D ratios less than 0.5.

~_1 7 1 (c) Improved Sequential Thermal Cracking 2 Process.

urning now to the sequential cracking process 2C
of the subjeet inven.ion, as illustrated in YIGURE 1~, in lieu of reactor 28 (see FIGUR~ 1) of the prior art, the system of the , invention includes a solids heater 4C, a primary reactor 6C, a 8 1' secondary reaetor 8C and downstream equipment. ~he downstream 9 1.
e~ui,oment is c~mprised essentially of an indireet heat exchanger lCC,a 1 1 \

1~ \

2`7 ~9 -21- \

s~w ` ~ 25~
696-1~7 , .

l I fractionation tower 12C, and a recycle line 14C from ~he 2 fractionation tower 12C to the entry of the primary reactor 3 6C.
4 I The system also includes a first hydrocarbon feed ' line 16C, a second hydrocarbon feed-quench line 18C~ a transfer line 20C and an air delivery line 22C.
7 The first hydrocarbon feed stream is introduced 8 into the primary reactor 6C and contacted with heated solids 9 from the solids heater 4C. The first or primary reactor 6C
in which the first feed is cracked is at high severity conditions~
ll The hydrocarbon feed, from line 16C~ may be any hydrocarbon gas 12 or hydrocarbon liquid in the vaporized state which has been used i3 heretofore as a feed to the conventional thermal cracking process.
l~ Thus~ the feed introduced into the primary reactor 6C may be lS ~ selected from the group consisting of low molecular weight hydro-16 carbon gases such as ethane, propane, and butane, light hydro-17 carbon liquids such as pentane, hexane, heptane and octane, low la boiling point gas oils such as naphtha having a boiling range l9 between 350 to 650F, high boiling point gas oils having a I boiling range between 650 to 950F and compatible combinations 21 of same. These constituents may be introduced as fresh feed 22 or as recycle streams through the line 14C from downstream 23 , puri~ication facilities e.g., fractionation to~er 12C. DilutiGri steam may also be delivered with the hydrocarbon through lines l 16C an~ 14C. The use of dilution steam reduces the partial 26 l' pressure, improves cracking selectivity and also lessens the 27 ~' tendency of high boiling aromatic components to form coke.

28 i The preferred primary feedstock for the high 29 severity reaction is a light hydrocarbon material selected from I the group consisting of low molecular weiqh~, hydrocarbon gases~

S&W
696-147 l, ~ ~ ~ w ~ ~ ~

1 light hydrocarbon liquids, light gas oils ~oiling between 350 2 and 650F, and combinations of same. These feedstocks offer the , 3 ~ greatest increase improvement in selectivity at high severity 4 and short residence times.
The hydrocarbon feed to the first reaction zone is 6 ; preferably pre-heated to a temperature of between 600 to 1200F
7 before introduction thereto. The inlet pressure in the line 16C
8 is 10 to 100 psig. The feed should be a gas or gasified liquid 9 The feed increases rapidly in tem2erature reaching thermal equi-librium with the solids in about 5 milliseconds. As mixing of 11 the hydrocarbon with the heated solid occurs, the final tem-12 perature in the primary reactor reaches about 1600 to 2000F. At 13 these temperatures a high severity thermal cracking reaction takes lq place. Tlle residence time maintained within the primary reactor is about 50 milliseconds, preferably between 20 and 150 milli-16 seconds, to ensure a high conversion at high selectivity. Typl-~7 cally, the ~SF (Kinetic Severity Function) is about 3~5 (97%
18 conversion of n-pentane). Reaction products of this reaction ~-lre 19 olefins, primarily ethylene with lesser amounts of propylene and butadiene, hydrogen, methane~ C4 hydrocarbons, distillates suc~
21 as gasoline and gas oilsl heavy fuel oils, coke and an acid gas.
22 ; Other products may be present in lesser quentities. Feed con-23 ! version in this first reaction zone is about between 95 to 100~ by 24 weight of feed, and the yield of ethylene for liquid feedstocks is about 25 to 45~ by weight of the feed, with selectivities o~
26 about 2.5 to 4 pounds of ethylene per pound of methane.
27 A second feed is introduced through the line 18C
28 ~ and combines with the cracked ~as from the primary reactor 6C
29 between the primary reactor 6C and the secondary reactor 8C~ The ' combined stream comprising the second unreacted feed, and the i S&W
6 9 ~ -1 a 7 1 first reacted feed passes ~hrough the secondary reactor 8C under 2 , low severity reaction conditions. The second feed introduced 3 l! through the line 18C is preferably virgin feed stock but may 4 11 also be comprised of the hydrocarbons previously mentioned~
including recycle streams containing low molecular weight 6 ; hydrocarbon gases, light hydrocarbon liquids, low boiling 7 point, light compatible gas oils, high boiling point gas oilst 8 and combinations of same.
9 i Supplemental dilution steam may be added with the ~ secondary hydrocarbon stream entering through s-tream 18C. ~owever, 11 ; in most instances the amount of steam initially delivered to the 12 primary reactor 16C will be sufficient to achieve the requisite 13 , partial pressure reduction in the reactors 6C and 8C. It should 14 be understood that the recycle stream 14C is illustrative, and not specific to a particular recycle constituent.
16 The hydrocarbon feed delivered through the line 18C
17 ' is preferably virgin gas oil 400-650F. The second feed is pr-~
18 !, heated to between 600 to 1200F. and upon entry into the seco!dary lg I reactor 8C quenches the reaction products from the primary re-actor to below 1500F~ It has been found that in general 100 21 pounds of hydrocarbon delivered through the line 18C will quench 22 l~ 60 pounds of effluent from the primary reactor 6C. At this ter"-~
23 l, perature level, the cracking reactions of the first feed are 2~ ~l essentially terminated. However, coincident with the ~uenching ~l of the effluent from the prlmary reactor, the secondary feed 26 ~I entering through line 18C is thermally cracked at this tempera~ure 27 ! (1500 to 1200F) and pressures of 10 to 100 psig at low severity 28 by providing a residence time in the secondary reactor between 29 1 150 and 2000 milliseconds, preferably between 250 to 500 milll-seconds. Typically, the KSF cracking severity in the secondary , I .

S&~
69fi-147 I;
. .

1 I reactor is about O.S at 300 to 400 milliseconds.
2 I The inlet pressure of the second feed in line 18C is 3 between 10 and 100 psig, as is the pressure of the first feed.
4 Reaction products from the low severity reaction zone comprise , ethylene with lesser amounts of propylene and butadiene, hydro-6 gen, methane, C4 hydrocarbons, petroleum distillates and gas 7 oils, heavy fuel oils, coke and an acid gas. Minor amounts of 8 other products may also be produced. ~eed conversion in this 9 second reaction zone is about 30 to 80% by weight of feed, ~ and the yield of ethylene is about 8 to 2G% by weight of feed, 11 ! with selectivities of 2.5 to 4.0 pounds of ethylene per pound 12 of methane 13 Although the products from the high severity reac~ion 14 are combined with the second feed, and pass through the second ~ reaction zone, the low severity conditions in the second reaction zone are insufficient to appreciably alter the product dis-17 I tribution of the primary products from the high severity react~on 18 ` ~one. Some chemical changes will occur, however these reaction ~ ' 19 products are substantially stabilized by the direct quench provided by the second feed.
~1 The virgin gas oils normally contain aromatic 22 molecules with paraffinic hydrocarbon side chains. For some 23 gas oils the nurnber of carbon atoms associated with such 24 ~ paraffinic side chains will be a large fraction of the total number of carbon atoms in the molecule, or the gas oil will ~6 have a low "aromaticity".

27 ; In the secondary reactor~ these molecules will 2~ I undergo dealkylation - splitting of the paraffin molecules/
29 ll leaving a reactive residual methyl aromatic, which will tend to react to form high boilers. The paraffins in the boiling .1, S&w ~ z 696-1~7 1 l~ range 400 to 6S0F are separated from the higher boiling 2 ; aromatics in column 1? and constitute tlle preferred recycle 3 ~ to the primary reactor.
4 , Other recycle feed stocks can include propylene, butadiene, butenes and the C5 - 400F pyrolysis gasoline.
6 ~ The total effluent leaves the secondary reactor 7 and is passed through the indirect quench means 10C to generate 8 ~ steam for use within and outside the system. The effluent is 9 then sent to downstream separation facilities 12C via line 24C~
10 i The purification facilities 12C employ conventional 11 separation methods used currently in thermal cracking processes~
12 FIGURE 2 illustrates schematically the products obtained. Hydro-13 gen and methane are taken overhead throush the line 36C. C4 and 14 lighter olefins, C5 - 400F and 400-650F fractions are removed from the fractionator 12C through lines 26C, 28C and 30C re-16 spectively. Other light paraffinic gases of ethane and propane 17 are recycled through the line 14C to the high severity primary 18 reactorO The product taken through line 28C consists of liquid ~ ' l9 ' hydrocarbons boiling between C5 and 400F, and is preferably exported although such material may be recycled to the primary 21 ~ reac~or 6C if desired. The light gas oil boilinq between 400 22 to 650F is the preferred recycle feed, but may be removed through 23 l line 30C. The heavy gas oil which boils between 650-950F is 2q exported through stream 32C, while excess residuim, boiling ~ above 950F is removed from the battery limits via stream 34C
~6 The heavy gas oil and residuim may also be used as fuel within 27 the system.
2~ ~ In the preferred embodiment of the process, the 29 I second feed would be one which is not recommended for high ~ severity operation~ Such a feed would be a gas oil boiling 696-147 1 ~

. . .

~ above 400F whlch COTItalns a slgniflcant amount of high molecular ,~ weiyht aromatic components. Generally, these components have 3 ; paraffinic side chains which will form olefins under proper conditions. However, even at moderate severity, the dealkylated aromatic rings will polymerize to form coke deposits. By pro-6 cessing the aromatic gas oil feed at low severity, it is possib1e 7 to dealkylate the rings, but also to prevent subsequent poly~
8 merization and coke Eormation. As a consequence of the low 9 severity, however, the yield of olefins is low, even though selectivity as previously defined is high. Hence, low severity 11 reaction effluents often have significant amounts of light 12 paraffinic gases and paraffinic gas oils. These light gases 13 ~ and paraffinic gas oils are recycled preferably to the high 14 I severity section, such compounds being the preferred feeds i5 ll thereto. The aromatic components of the effluent are removed 1~ I from the purification facilities 12C as part of the heavy gas 17 ! oil product, and either recycled for use as fuel within the 18 system, or exported for further purification or storage.
19 An illustration of the benefits of the process of the invention is set forth below wherein feed cracked and the 21 I resultant product obtained under conventional high severity 22 cracking and quenching conditions is compared with the same feed 23 sequentially cracked in accordance with this invention.

-25 , \

28 , i ,",. , _ S&~

1 (d) Improved Residerlce Time Solid~Gas Separation 2 Device and System.

- Referring to FIGURE 15 in the subject invention, in lieu separation zone or curved segment region 36 and the quench area 6 44 of the prior art TRC system (see FIGUR~ 1), solids and gas enter 7 the tubular reactor 13D through lines llD and 12D respectively.
8 The reactor effluent flows directly to separator 14D where a 9 separation into a gas phase and a solids phase stream is effected.
The gas phase is removed via line 15D, while the solid phase is 11 sent to the s-tripping vessel 22D via line 16D. Depending upon 12 the nature of the process and the degree of separation, an in line 13 quench of the gas leaving the separator via line 15D may be made 14 by injecting quench material from line 17D. Usually, the product sas contains residual solids and is sent to a secondary separator 16 ~18D, preferably a conventional cyclone. Quench material should 17 be introduced in line 15D in a way that precludes back flow of 18 quench material to the separator. The residual solids are removed from separator 18D via line 21D, while essentially solids free product gas is removed overhead through line l9D. Solids from lines 16D and 21D are stripped of gas impurities in 22 fluidized bed stripping vessel 22D using steam or other inert 23 fluidizing gas a~nitted via line 23D. Vapors are removed from 24 the stripping vessel through line 24D and, if economical or if 26 need be, sent to down-stream purification units~ Stripped solids 27 ~ _ -,1 s ~w jl, 1 ll removed from the vessel 22D through line 25D are sent to re-2 ,I generation vessel 27D using pneumatic kransport gas from line 3 , 26D. Off gases are removed from the regenera-tor through line 28D.
4 1 After regeneration the solids are then recycled to reactor 13D
, via line llD.
6 , The separator 14D should disenga~e solids rapidly 7 ;~ from the reactor effluent in order to prevent product degradation 8 i and ensure optimal yield and selectivity of the desired products.
9 l Further, the separator 14D operates in a manner that eliminates j or at least significantly reduces the amount of gas entering the 11 stripping vessel 22D inasmuch as this portion of the gas product 12 would be severely degraded by remaining in intimate contact with 13 the solid phase. This is accomplished with a positive seal which 14 has been provided between the separator 14D and the stripping ' vessel 22D. Finally, the separator 14D operates so that 16 erosion is minimized despite high temperature and high velocity 17 , conditions that are inherent in many of these processes~ The 18 separator system of the present invention is designed to meet 19 each one of these criteria as is descri~ed below.
FIGUR~16 is a cross sectional elevational view 21 1I showing the preferred embodiment of solids-gas separation devlce 2 ~ j 14D of the present lnvention. The separator 14D is provided 23 with a separator shell 37D and is com?rised of a solids-gas ~4 , disengaging chamber 31D having an inlet 32D for the mixed phas2 ~ stream, a gas phase outlet 3i3D, and a solids phase outlet 34D.
26 The inlet 32D and the solids outlet 34D are preferably locate~
27 ' at opposite ends of the chamber 31D. While the gas outlet 33~
28 , lies at a point therebetween. Clean-out and maintenance manways 29 35D and 36D may be provided at either end of the chamber 31D.

; The separator shell 37D and manways 35D and 36D preferably are ,j ~&W
6~6-la7 , ~ :

1 l lined w~ eroslon resi~tent linings 3~D, 39D and 41D re-2 spectively which may be required if solids at hi~h velocities 3 j are encountered. Typical commercially available materials 4 for exosion resistent llning include Carborundum Precast Carbofrax D, Carborundum Precast Alfrax 201 or their equivalent 6 h thermal insulation lining 4OD may be placed between shell 37D
7 ~ and lining 38D and between the manways and their respective 8 1 erosion resistent linings when the separator is to be used 9 in high temperature service. Thus, process te~peratures above 0 1500F. (870C.) are not inconsistent with the utilization of 11 this device.
12 FIGURE 17shows a cutaway view of the separator 13 along section ~4. For greater strength and ease of construction 14 the separator 14D shell is preferably fabricated from cylindrical sections such as pipe SOD, although other materials may, of 16 ~ course, be used. It is essential that longitudinal side walls 17 51D and 52D should be rectilinear, or slightly arcuate as in-18 , dicated by the dotted lines 51D and 52D. Thus, flow path 31D
19 through the separator is essentially rectangular in cross ~ section having a height H and width W as shown in FIGURE 17.
21 The embodiment shown in FIGURE 17 defines the geometry of the 22 flow path by adjustment of the lining width for walls 51D and 23 j 52D. Alternatively, baffles, inserts, weirs or other means 24 ~I may be used. In like fashion the configuration of walls 53D
~5 `~ and 54D transverse to the flow path may be similarly shaped, 26 although this is not essential. FIG~RE 18is a cutaway view along Section 4-4 of FIGURE16 wherein the separation shell 37~
28 1 is fabricated from a rectangular conduit. Because the shell 37D
29 has rectilinear walls 51D and 52D it is not necessary to adjust the width of the flow path with a thickness of lining. Linings ' -30-`~_ -38D and 40D could be added for erosion and thermal resistence respectively.
Again referring to Figure 16 inlet 32D and outlets 33D are disposed normal to flow path 31D (shown in Figure 17) so that the incoming mixed phase stream from inlet 32D is re~uired to undergo a 90 change in direction upon entering the chamber. As a further requirement, however, the gas phase outlet 33D is also oriented so that the gas phase upon leaving the separator has completed a 180 change in direction.
Centrifugal force propels the solid particles to the wall 54D opposite inlet 32D of the chamber 31D, while the gas portion, having less momentum, flows through the vapor space of the chamber 31D. Initially, solids impinge on the wall 54D, but subseguently accumulate to form a static bed of solids 42D, which ultimately form in a surface configuration having a curvilinear arc 43D of approximately 90. Solids impinging upon the bed are moved along the curvilinear arc 43D
to the solids outlet 34D which is preferably oriented for downflow of solids by gravity. The exact shape of -the arc 43D
is determined by the geometry of the particular separator and the inlet stream parameters such as velocity, mass flowrate, bulk density, and particle si~e. Because the force imparted to the incoming solids is directed against the static bed 42D
rather than the separator 14D itself, erosion is minimal.
Separator efficiency, defined as -the removal of solids from the gas phase leaving through outlet 33D, is, therefore, not affected adversely by high inlet velocities up to 150 ft./sec., and the separator 14D is operable over a wide range of dilute phase densities, preferably between 0.1 and 10.0 lbs./ft3. The separator 14D of the present invention achieves efficiencies of about ~0%, although the preferred embodiment, discussed below, can obtain over 90~ removal of solidsO

., ~

It has been found that separator efficiency is dependent upon separator geometry inasmuch as the flow path must be essentially rectangular and the relationship between height H, and the sharpness of the U-bend in the gas flows.
Referring to ~igures 16 and 17 we have found that for a given height H of chamber 31D, efEiciency increases as the 180 U-bend between inlet 32D and outle-t 33D becomes progressively sharper; that is, as outlet 33D is brought progressively closer to inlet 32D. Thus, for a given H the efficiency of the separator increases as the Elow path and, hence, residence time decreases. Assuming an inside chamber di of inlet 32D, the preferred distance CL between the centerlines of inlet 32D and outlet 33D is less than 4.0 di~
while the most preferred distance between said centerlines is between 1. 5 and 2. ~ di. Below 1. 5Di better separation is obtained but difficulty in fabrication makes this embodiment less attractive in most instances. Should this latter embodiment be desired, the separator 14D would probably require a unitary casting design because inlet 32D and outlet 20 33D would be too close to one another to allow welded fabrication.
It has been found that the height of flow path ~1 should be at least equal to the value of Di or 4 inches in height, whichever is greater. Practice teaches that if ~ is less than Di or 4 inches the incoming stream is apt to disturb the bed solids 42D, thereby re-entraining solids ln the gas product leaving through outlet 33D. Preferably H is on the order of twice Di to obtain even greater separation efficiency. While no-t otherwise limi-ted, it is apparent that 30 too large an H eventually merely increases residence time wi-thout substantive increases in efficienc~. The width W of the flow path is s ~

1 preferably between 0.75 and 1.25 times Di, most preferably between 2 I 0.9 and 1.10Di.
3 Outlet 33D may be of any inside diameter. However~
4 , velocities greater than 75 ft./sec. can cause erosion because ', of residual solids entrained in the gas. The inside diameter 6 ~ o outlet 34D should be sized so that a pressure differential 7 1 between the stripping vessel 22D shown in EIGURE1s and the 8 , separator 14D exist such that a static height of solids is g I formed in solids outlet line 16D. The static height of solids in line 16D forms a positive seal which prevents gases from 11 l entering the stripping vessel 22D~ The magnitude of the 12 pressure differential between the stripping vessel 22D and the 13 separator 14D is determined by the force required to move the 14 ;' solids in bulk flow to the solids outlet 34D as well as the 1 ehight of solids in line 16D. As the differential increases 16 , the net flow of sas to the stripping vessel 22D decreases.
17 Solids, having ~ravitational momentum, overcome the differential, 18 while gas preferentially leaves through the gas outlet 33D.
19 sy reaulating the pressure in the stripping vessel ' 22D it is possible to control the amount of ~as going to the 21 stripper. The pressure regulating means may include a check 22 or "flapper" valve 29D at the outlet of line 16D, or a pressure 23 l control 29D device on vessel 22Do Alternatively, as suggested 24 j~ above, the pressure may be re~ulated by selecting the size of ! the outlet 34D and conduit 16D to obtain hydraulic forces 26 ~I acting on the system that set the flow of gas to the stripper 27 1 32D. ~hile su~h gas is degraded, we have found tha-t an increase 28 ' in separation efficiency occurs with a bleed of ~as to the 29 ' stripper of less than 10~, preferably between 2 and 7~. Economic ~ and process considerations would dictate whether this mode of ,~, S&W , ~ g 696-1~7 ~
,1 .

1 1 operation should be used. It is also possible to design the 2 system to obtain a net backflow of gas from the stripping 3 ~I vessel. This ~as flow should be less than 10~ of the total 4 ~ feed gas rate.
sy establishing a minimal flow path, consistent 6 ~ with the above recor~nendations, residences times as low as 7 0.1 seconds or less may be obtained, even in separators 8 ,I having inlets over 3 feet in diameter. Scale-up to 6 feet 9 l' in diameter is possible in rnany systems where residence times ' approaching 0.5 seconds are allowable.
11 In the preferred embodiment of FIGURE 16, a weir 44D
12 ' is placed across th~ flow path at a point at or just beyond the 13 i gas outlet to establish a positive height of solids prior to 14 li solids outlet 34D. By installing a weir (or an equivalent I restriction) at this point a more stable bed is established 16 thereby reducing turbulence and erosion~ Moreover, the weir 17 , 44D establishes a bed which has a crescent shaped curvilinear 18 arc 43D of slightly more than 90~0 An arc of this shape 19 diverts gas towards the gas outlet and creates the U-shaped I gas glow pattern illustrated diagrammatically by line 45D in 21 FIGURE 16. Without the weir 44D an arc sor,lewhat less than or 22 equal to 90 would be formed, and which would extend asymptoti-23 cally toward outlet 34D as shown by dotted line 60D in the 24 'i schematic diagram of the separator of FIGURE 19. While neither efficiency nor gas loss (to the stripping vessel) is affected 26 ! adversely, the flow pattern of line 61D increase~ residence timet 27 ~ and more importantly, creates greater poten~ial for erosion at 28 l areas 62D, 63D and 64D.
29 The separator of FIGURE 20 is a schematic diagram of ! another embodiment of the separator 14D, said separator 14D

-3~-!

696-1~7 ~9~2~

1 1 having an extended separation chamber in the lon(Jitudinal 2 1; dimension. Here, the horizontal distance L between the gas 3 outlet 34D and the weir 44D is extended to establish a solids 4 ! bed of greater leng-th. L is preferably less than or equal I to 5 Di. Although the gas flow pattern 61D does not develope 6 I the preferred V-shape, a crescent, shaped arc is obtained 7 which limits erosion potential to area 64D. Embodiments 8 11 shown by FIGURES 19 and 20are useful when the solids loadinq 9 j of the incoming stream is low. The e~bodiment of FIGURE 19 also has the minimum pressure loss and may be used when the 11 velocity of the incoming stream is low.
12 As shown in FI5UR~21 it is equally possible to use a 13 , stepped solids outlet 65D having a section 66D collinear with 14 the flow path as well as a gravity flow section 67D. Wall 68D

-5 replaces weir 44D, and arc 43D and flow pattern 45 are similar 16 ; to the preferred embodiment of FIGURE 16. Because solids accumu-17 i late in the restricted collinear section 66D, pressure losses 18 are ~reater. This embodiment, then, is not preferred where the 19 incoming stream is at low velocity and cannot supply sufficien-t force to expel the solids through outlet 65D. However~ because 21 of the restricted solids flow path, better deaeration is obtained 22 l and gas losses are minimal.
23 j FIGURE 22 illustrates another embodiment of the 24 ~ separator 14D of FIGURE 21wherein the solids outlet is steppeG.
l` Althoush a weir is not used, the outlet restricts solids flow ,26 which helps from the bed 42D. As in FIGURE 20, an extended L
27 distance between the gas outlet and solids outlet may be used~
28 The separator of FIGURE21 or22 may be used in 29 I conjunction with a venturi, an orifice~ or an equivalent flow restriction device as shown in FIGURE 28. The venturi 69D having S&W ~ 2~
696-147 !, 1 ~, dimensions Dv (diameter at venturi inlet), Dv~ (diameter of 2 I venturi tl-roat), and ~ (angle of cone formed by projec~ion 3 , of convergent venturi walls) is placed in the collinear section 4 66D of the outlet 65D to greatly improve deaeration of solids.
The embodiment of FIGURE 24 is a variation of the separator

6 shown inFIGURE 23. Here, inlet 32 and outlet 33D are oriented

7 for use in a riser type reactor. Solids are propelled to the

8 wall 71D and the bed thus formed is kept in place by the force g of the incoming stream. As before the gas portion of the feed , follows th~ U-shaped 2attern of line 95D. However, an asymptotic 11 bed will be formed unless there is a restriction in the solids 12 outlet. A weir would be ineffective in establishing bed height~
13 and would deflect solids into the gas outlet. For ~his reason 14 ' the solids outlet of FIG~RE 23 is preferred. ~Sost preferably, ; the venturi 69D is placed in collinear section 66D as shown in 16 ; FIGURE24 to improve the deaeration of the solids. Of course, 17 each of these alternate embodiment may have one or more of ~he 1~ optional design features of the basic separator discussed in 19 relation to FIGURES 1~ 17 and 18.
The separator of the present invention is rnore 21 clearly illustrated and explained by the examples which follow.
22 In these examples, which are basedon data obtained during 23 experimental testiny of the separator design, the separator 24 has critical dimensions specified in Table I. These dimensions (in inches except as noted) are indicated in the various drawlng 26 figures and listed in the Nomenclature below.
Distance between inlet and gas outlet centerlines 28 ' Di Insi~e diameter of inlet 29 ~ Dog Inside diameter of gas outlet 30 , Dos Inside diameter of solids outlet Dv Diameter of venturi inlet Dvt Diameter of venturi t~roat ;

~ ~ ~ ~4 H Height of flow p~th Hw Height of weir or step L Length from gas outlet to weir or step as indicated in Figure 6 W Width of flow path Angle of cone formed by projection of convergent venturi walls, degrees Table 1 Dimensions of Separators in Examples 1 to 10, i.nches*
Example Dimension 1 2 3 4 5 6 7 8 9 10 CL 3.875 3.875 3.875 3.875 3.8753.87511 11 3.5 3.5 Di 2 2 2 2 2 2 6 6 2 2 Dog1.751.75 1.75 1.75 1.75 1.75 4 4 Dos 2 2 2 2 2 2 6 6 2 2 Dv ~ ~ _ 2 Dvt ~ - ~ ~ ~ ~ ~ 1 H 4 4 4 4 4 4 12 12 7.5 6 Hw 0~750.75 0.75 0.75 0.75 0.75 2.252.25 0 4 W 2 2 2 2 2 ~ 6 6 2 2 0,degrees - - - - - - - - - 28 * Except as noted Example 1 In this example a separator of the preferred embodiment of Figure 16 was tested on a feed mixture of air and silica alumina. The dimensions of the apparatus are specified in Table I. Note that the distance L from the gas outlet to the weir was zero.

S&~

1 ll The inlet stream was comprised of BS ft, 3/min.
2 1l of air and 52 lbs./min. of silica alumina having a hulk densit~
3 ~, of 70 lbs./ft3 and àn average particle size of 100 microns.
4 ~ The stream density was 0.612 lbs./ft.3 and the operation was ;, performed at ambient te~perature and atmospheric pressure.
6 I The velocity of the incoming stream through the 2 inch inlet 7 ~ was 65.5 ft./sec., while the outlet gas velocity was 85.6 ft./sec.
8 l' through a 1.75 inch diameter outlet. A positive seal of solids

9 ' in the solids outlet prevented gas from being entrained in the solids leaving the separator. Bed solids were stabilized by 11 placing a 0.75 inch weir across the flow path.
12 I The observed separation efficiency was 89.1%, 13 and was accomplished in a yas phase residence tirne of approximately 1~ 1 0.008 seconds. Efficiency is defined as the percent removal of I solids from the inlet stream.
16 Example 2 17 ' The gas-solids mixture of Example 1 was processed 18 in a separator having a configuration illustrated by FIGURE 20.
19 I In the example the L dimension is 2 inches; all other dimensions are the same as Example 1. By extendin~ the separation chamber 21 along its longitudinal dimension, the flow pattern of the gas 22 ' began to deviate from the U-shaped discussed above. As a result 23 , residence time was longer and turbulence was increased. Separa,tion 24 ,l efficiency for this example was 70.8~.
1l Example 3 26 1 The separator of Example 2 was tested with an inlet 27 ~ stream comprised of 85 ft.3/min. of air and 102 lbsi/min. of 28 I silica alumina which gave a stream density of 1.18 lbs./ft.3 29 1 or approximately twice that of Example 2. Separation efficiency l improved to 83.8%.

il -3~-;

2g~

Example 4 The preferred separa-~or of Example 1 was tested at the inlet flow rate of Example 3. Efficiency increased slightly to 91.3%.
Example 5 The separator of Figure 16 was tested at the conditions of Example 1. Although the separation dirnensions are specified in Table I note that the distance CL between inlet and gas outlet centerlines was 5.875 inches, or about three times the diameter of the inlet. This dimension is outside the most preferred range for CL which is between 1.50 and 2.50 Di.
Residence time increased to 0.01 seconds, while efficiency was 73.0%.
Example 6 Same conditions apply as for Example 5 except that the solids loading was increased to 102 lbs./min. to give a strea~m density of 1.18 lbs./ft.3. As observed previously in Examples 3 and 4, the separator efficiency increased with higher solids loading -to 90.6%.
Example 7 The preferred separator configuration of Figure 16 was tested in this Example. However, in this example the apparatus was increased in size over the previous examples by a factor of nine based on flow area. A 6 inch inlet and 4 inch outlet were used to process 472 ft.3/min. of air and 661 lbs./min of silica alumina at 180F. and 12 psig. The respective velocities were 40 and 90 ft./sec. The solids had a bulk density of 70 lbs./ft3 and the stream density was 1.37 lbs./f-t.3 Distance CL between inlet and gas outlet centerlines was 11 inches, or 1~83 times the inlet diameter; distance L was zero. The bed was ~&w , 69~-147 1 stabllized by a 2.25 inch weir, and gas loss was prevented 2 by a positive seal of solids. However, the solids were 3 ; collected in a closed vessel, and the pressure differential 4 I was such that a positive flow of displaced gas from ~he j collection vessel to the separator was observed. This volume 6 ; was approximately 9.4 ft.3/min. Observed separator efficiency 7 was 90.0~, and the gas phase residence time approximately 8 0.02 seconds.
9 Example 8

10 , The separator used in Example 7 was tested with

11 an ldentical feed of gas and solids. However, the solids

12 collection vessel was vented to the atmosphere and the pressure

13 differential adjusted such that 9% of the feed gas, or 42.5 fto /

14 , min. exited through the solids outlet at a velocity of 3.6 , ft./sec. Separator efficiency increased with this positive 16 bleed through the solids outlet to 98.1%.
17 Example 9 18 The separator of FIG~RE 22was tested in a unit 19 I having a 2 inch inlet and a 1 inch gas outlet. The solids out-let was 2 inches in diameter and was located 10 inches away 21 from tne ~as outlet (dlmension L)o A weir was not used. The 22 feed was comprised of 85 ft.3/min. of air and 105 lbs./min. of spent fluid catalytic cracker catalyst having a bulk density 24 of 45 lbs./ft.3 and an average particle si~e of 50 microns. This ~ave a stream density of 1.20 lbs./ft.3 Gas inlet velocity was . 26 65 ft./sec. while the gas outlet velocity was 262 ft./sec. As ~7 in Example 7 there was a positive counter-current flow of 28 ' displaced gas from the collection vessel to the separator.
29 I Th1s flow was approximately 1.7 fto3/ min. at a velocity of 1.3 ft./sec. Operation was at ambient temperature and atmos~

31 pherlc pressure. Separator efficiency was 95.0~.

, 696-147 l . I , 1 ,,Example 10 ;!
2 'The separator of ~IGURE23 was tested on a feed 3 comprised of 85 ft.3/ min. of air and 78 lbs./minO of spent 4 il Fluid Catalytic Cracking catalyst. The inlet was 2 inches in , diameter which resulted in a velocity of 65 ft./sec., the gas 6 , outlet was 1 inch in diameter which resulted in an o~tlet 7 I velocity of 262 ft ./sec. This separator had a stepped 8 solids outlet with a venturl in the collinear section of the 9 outlet. The venturi mouth was 2 inches in diameter, whlle 10the throat was 1 inch. A cone of 281.1~ was formed by pro-11 jection of the convergent walls of the venturi. An observed 12 efriciency of 92.6~ was measured, and the solids leaving the i3 ~ separator were completely deaerated except for interstitial gas 14 remaining in the solidsl voids.

8 ~ \
1 9 , \
20 ;

~3 , \

27 ' \

29 ' \
30 ,, -,;~

~11 1 (e) ImDroved Solids Quench Boiler and 2 Process.

As see~n in FIGURE 25,in lieu of quench zone 44, 46 ~see 6 E'IGURE 1) of the prior art, the composi~e solids quench boiler 2E

7 ~ of the subject invention is comprises essentially of a quench ex~

8 changer 4E, a fluid bed-quench riser 6E, a cyclone seoarator 8E

9 with a solids return line lOE to the fluid bed-riser 6E and a line 1~ 36E for the delivery of gas to the fluid bed-quench riser.
11 The quench exchanger 4E as best seen in FIGURES 26 an~ ~7 ~2 ls formed with a plurality o~ concentrically arranged tubes ex-13 tending parallel to the longitudinal axis of the quench exchanger 14 4E. The outer circle of tubes 16E form the outside wall of the ~4 \

-~2--5&~1 6n6-147 1 quench exchanger 4E. The tubes 16E are joined together, pre 2 ~ ferably by welding, and form a pressure-tight mem~rane wall which 3 is in effect, the outer wall of the quench exchanger 4E. The 4 inner circles of tubes 18E and 20E are spaced apart and allo~
for the passage of effluent gas and particulate solids there-6 around. The arrays of tubes 16E, 18E and 20E are manifolded 7 to an inlet torus 24E to which boiler feed water is delivered 8 and an upper discharge torus 22E from which high pressure steam 9 is discharged for system service. The quench exchanger 4Eis provided with an inlet hood 26E and an outlet hood 28E, to 11 insure a pressure tight vessel. The quench exchanger inlet hood 12 26E extends from the quench riser 6E to the lower torus 24E.
13 The quench exchanger outlet hood 28E extends from the upper 14 ; torus 22E and is connected to the downstream piping equipment lS by piping such as an elbow 30E which is arranged to deliver tr.e 1~ cooled effluent and particulate solids to the cyclone separator 17 ~ 8E.
18 The fluid bed quench riser 6Eis essentially a sealed 19 vessel attached in sealed relationship to the quench exchanger 4E.The fluid bed-quench riser 6Eis arranged to receive the 21 reactor outlet tube 36E which is preferably centrally disposec~ at 22 the bottom of the fluid quench riser 6E. A slightly enlarged 23 , centrally disposed tube 38E is aligned with the reactor outlet.
24 ;36E and extends from the fluid bed-quench riser 6E into the quench exchanger 4E. In the quench exchanger 4E~ the centrally `26 disposed fluid bed-quench riser tube 38E terminates in a conical 27 opening 4UE. The conical opening 40E is provided to facilitate 28 nonturbulent transition from the quench riser tube 38E to the 29 enlarged opening of the quench exchanger 4E. It has been round that the angle of the cone ~, best seen in FIGURE26, should 31 be not greater --han 10 degrees.

~r-S&W~ 3~2~

1 ¦ The fluid bed 42E contained in the fluid bed quench 2 riser 4E is maintained at a level well above the bottom of the 3 quench riser tube 38E. A bleed line 50E is provided to bleed 4 , solids from the bed 42E~ Although virtually any particulate ~ solids ean be used to provide the quench bed 42E, it has been 6 `, found in practice that the same solids used in the reactor are 7 ' preferably used in the fluidized bed 42E. Illustrations of 8 , the suitable particulate solids are FCC alumina solids.
9 , As best seen in FIG~RE 28,the opening 48E thxough ~ which the fluidized particles from the bed 42E are drawn into 11 ' the quench riser tube 38E is defined by the interior of a coné
12 ' 44E at the lower end of the quench riser tube 38E and a refractory 13 ' cone 46E located on the outer surface of the reactor outlet 14 ' tube 36E. In practice, it has been found that the refractory ~ cone 46E can be formed of any refractory rnaterial. The opening 16 ' 48, defined by the conical end 44E of the quench riser tube 38E
17 and the refractory cone 46E, is preferably 3-4 square feet for 18 a unit of 50 ;~IBTU/HR capacity. The opening is sized to insure 19 I penetration of the cracked gas solid mass velocity of 100 to ' 800 pounds per second per square foot is required. The amoun~
21 l, of solids from bed 42E delivered to the tube 38E is a functior, 22 l' Of the velocity of the gas and solids entering the tube 38E
23 j from the reactor outlet 36E and the size of the opening 48E~
24 In practice, it has been found that the Thermal ll Regenerative Cracking (TRC) reactor effluent will contain ' 26 ~' approximately 2 pounds of solids per pound of gas at a tem-~7 1l perature of about 1,400F to 1,600F.
28 , The process of the solids quench boiler 2E of 29 FIGURES 25-2~ is illustrated by the following example. Fffluent I from a TRC outlet 36E at about 1,500F is delivered to the quenc~

~ -~4-~i S&W
6~6-1~7 ~

1 I riser tube 3~E at a velocity of approximately 40 to 100 feet 2 l, per second. The ratio of particula~e solids to cracked effluent 3 j entering or leaving the tube 36E is approximately two pounds o 4 ', solid per pound of gas at a temperature of about 1,500F. At , 70 to 100 feet per second the particulate solids entrained into 6 , the effluen~ stream by the eductor effect is between twenty five 7 and fifty pounds solid per pound of gas. In 5 milliseconds the 8 ' addition of the particulate solids from the bed 42E which is 9 at a temperature of 1,000F reduces the temperature of the ~ composite effluent and solids to 1,030F. The gas-solids mixture 11 is passed from the quench riser tube 38E tO the quench exchanger 12 4E wherein the temperature-is reduced from 1,030F to 1,000F
13 by indirect heat exchange with the boiler feed water in ~he tubes 14 ; 16E, l~E, and 20E. With 120,000 pounds of effluent per hour~

15 ' 50 MMBTUs per hour of steam at 1,500 PSIG and 600F will be

16 generated for system service. The pressure drop of the gas

17 , solid mixture passing through quench exchanger 4E is 1.5 PSI. The

18 l, cooled gas-solids mixture is delivered through line 30E to the

19 I cyclone separator 8E wherein the bulk of the solids is removed frbm the quenched-cracked gas and returned through line lOE
21 I to the quench riser 6E~

23 l' ~ v 24 ~ \

'26 l \
~7 1, ~
28 '~ ~ \

30 l, \

.
, -45-s~w 1 (f) Improved Preheat Vaporization System.

3 Again referring to FI~29,in lieu of preheat zone 24 (FI~ 1) 4 of ~e system 2F of the subject invention is e~bodied in a TRC system and is co~prised of essentially a liquid feed heater 4F, a mixer 8F for flashing 6 steam and the heated feedstock, a separator lOF to separate 7 the flashed gas and liquid, a vapor feed superheater 12F, and 8 a second mi~er 14F for flashing. The system also prefer~ntially includes a knockout drum 16F for the preheated vapor.
The liquid feed heater 4F is provided for heating the hydrocarbon feedstock such as desulfurized Kuwait ~GO to 13 initially elevate the temperature of the feedstock.
The initial mixer 8F is used in the system 2F to initially flash superheated steam from a steam line 6F and the heated feedstock delivered from the llquid feed heater 4F by 16 a line 18F.
17 The system separator lOF is to separate the liquid and 18 ' vapor produced by flashing in the mixer 8F. Separated gas is ~'1 \ ., -636-147 1~ 29 1 discharged through a line 22F from the separator overhead and 2 the remaining liquid is discharged through a line 26F.
3 l, A vapor feed superheater 12F heats the gaseous overhead 4 'from the line 22F to a high temperature and discharges the ~;heated vapor through a line 24F.
6 'I The second mixer 14F is provided to flash the vaporized ;
7 `gaseous discharge from the vapor feed superheater 12F and the 8 , liquid bottoms from the separator lOF, thereby vaporizing the 9 composite steam and feed initially delivered to the system 2F.
~ knockout drum 16F is employed to remove any liquid 11 from the flashed vapor discharged from the second mixer 14F
12 ; through the line 28F. The liquid-free vapor is delivered to a 13 reactor through the line 30F.
14 , In the subject process, the heavy oil liquid hydro-,carbon feedstock is first heated in the liquid feed heater 4F
16 to a tem~erature of about 440 to 700F. The heated heavy 17 , oil hydrocarbon feedstock is then delivered through the line 18 ~18F to the mixer 8F. Superheated steam from the line 6F if 19 mixed with the heated heavy oil hydrocarbon feedstock in the ,Imixer 8F and the steam-heavy oil mixture is flashed to about 21 ~,700 to 800F. For lighter feedstock the flashing temperature 22 will be about 500 to 600F., and for heavier feedstock the ~3 flashing temperature will be about 700 to 900F.
24 ,I The flashed mixture of the steam and hydrocarbon is ` sent to the system separator lOF wherein the vapor or gas is 26 taken overhead through the line 22F and the liquid is 27 discharged through the line 26F. Both the overhead vapor and 28 ' liquid bottoms are in the temperature range of about 700 to 29 ~800F. The temperature level and percent of hydrocarbon ' vaporized are determined within the limits of equipment fouling ~ 47-s~w 1 `criteria. The vapor stream in the line 22F is comprised of 2 essentially all of the steam delivered to the system 2F and 3 a large portion of the heavy oil hydrocarbon feedstock.
4 Between 30% and 70% of the heavy oil hydrocarbon feedstock supplied to the system will be contained in the overhead 6 leaving the separator 10F through the line 22F.
7 The steam-hydrocarbon vapor in the line 22F is delivered 8 to the system vapor feed superheater 12F wherein it is heated to 9 about 1,030F. The heated vapor is taken from the vapor feed superheater 12F throush the line 24F and sent to the second mixer 11 14F. Liquid bottoms from the separator 10F is also delivered 12 to the second mixer 14F and the vapor-liquid mix is flashed in 13 the mixer 14F to a temperature of about 1,000F.
14 The flashed vapor is then sent downstream through the ' line 28F to the knockout drum 16F for removal or any liquid 16 from the vapor. Finally, the vaporized hydrocarbon feed is 17 ~ sent through the line 30F to a reactor.
18 An illustration of the system preheat process is 19 seen in the following example.
A Nigerian Heavy Gas Oil is preheated and vaporized in 21 ~e system 2F prior to delivery to a reactor. The Nigerian Heavy 22 j Gas Oil has the following composition and properties:
~3 l ~4 Flemental Analysis, Wt.% ProPerties Carbon 86.69 Flash Point, F. 230.0 Hydrogen 12.69 Viscosity, SUS 210 F 44.2 '26 Sulfur .10 Pour Point, F ~90.0 Nitrogen .047 Carbon Residue, Ramsbottom .09 27 Nickel .10 Aniline Point, C 87.0 ~ Vanadium .10 S&W

l Distillation 2 I Vol. $
3 1 IsP
10669 ~ 2 4 30755 ~ 6 50820 ~ 4 70874 ~ 4 9094~6 6 ~P1~005~8 8 3 ~ 108 pounds per hour of the Nigerian Heavy Gas Oil is 9 heated to 750F~ in the liquid feed heater 4F and delivered at a pressure of 150 psia to the mixer 8F. 622 pounds per hour of ll superheated steam at l,100F. is simultaneously delivered to t~e 12 mixer 8F~ The pressure in the mixer is 50 psia.
13 The superheated steam and Heavy Gas Oil are flashed in 14 , the mixer 8F to a temperature of 760F~ wherein 60 of the Heavy Gas Oil is vapori7ed.
16 The vapor and liquid from the mixer 8F are separated 17 in the separator lOF ~ 622 pounds per hour of steam and 1, 864 ~ 8 18 pounds per hour of hydrocarbon are taken in line 22F~ as overhead l9 ; vapor. 1~243~2 pounds per hour of hydrocarbon are discharged through the line 26F as liquid and sent to the mixer 14~
21 The mixture of 622 pounds per hour of steam and 22 1~864~8 pour-ds per hour of hydrocarbon are superheated in the 23 ,vapor superheater 12F to 1~139QF~ and delivered through line 24 24F to the mixer 14Fo The mixer 14F is maintained at 4S psia.
~5 The 1~243~2 pounds per hour of liquid at 760F~ and 26 the vaporous mixture of 622 pounds per hour of steam and 27 1~864~8 pound per hour of h~drocarbon are flashed in the mixer 28 14F to 990F.
29 The vaporization of the hydrocarbon is effected with a ,Isteam to hydrocarbon ratio of 0.2. The heat necessary to vaporize _~9_ 8~29 S&W

,, .

1 the hydrocarbon and generate the necessary steam is 2.924 MM
2 ,BTU/hr.
3 ~I The same 3,108 pounds per hour of Nigerian Heavy Gas 4 Oil feedstock vaporiæed by a conventional flashing operation 'Irequires steam in a 1~1 ratio to maintain a steam temperature 6 of 1,434F. The composite heat to vaporize the hydrocarbon and 7 l generate the necessary steam is 6.541 MM BTU/hr. In order to 8 reduce the input energy in the conventional process to the same 9 level as the present invention, a steam temperature of 3,208F~
is requiredt which temperature is effectively beyond design 11 ;llimitations.

13 !1 SU~ARY
14 With reference to the new and improved separation , (see FIGURES 15-24), it is noted that short residence time 16 favors selectivity in C2H4 production. This means that the 17 reaction must be quenched rapidly. When solids are used, they 18 must be separated from the gas rapidly or quenched with the gas.
19 I~If the gases and solids are not separated rapidly (but separated) as in a cyclone, and then quenched, product 21 degradation will occur. IE the total mix is quenched, to avold 22 rapid separationr a high thermal inefficiency will result since 23 lall the heat of the solids will be rejected to some lower 24 llevel heat recovery. Thus, a rapid high efficiency separator, ,according to the subject invention, is optimal for a TRC process~
26 Similarly, in connection with the subject solids 27 feed device (see FIGURES 4-13~, it is noted that in order to 28 feed ~olids to an ethylene reactor, the flow must be controlled 29 jto within +2 percent or cracking severity oscillations will be I greater than that presently experienced in coil cracking. The 1, -50-S&~ 2~

1 subject feed devlce (local fluidization) minimizes bed height 2 as a variable and dampens the effect of over bed pressure fluct-3 uations, both of which con-tribute to flow fluctuations. It is a thus uniquely suited to short residence time reactions. Further, for short residence time reactions, the rapid and intimate 6 mixing are critical in obtaining good selectivity (minimize 7 mixing time as a percentage of total reactlng ti~e). Both of 8 the features permit the TRC to move to shorter residence times 9 which increase selectlvity. Conventional fluid bed feeding devices are adequate for longer ti~e and lower temperature 11 reactions (FCC) especi211y catalytic ones where minimal reaction 12 occurs if the solids are not contacting the gas (poor mixing).
13 In connection with the solids quench boiler 14 (see FIGURES 25-28), in the current TRC concept, a 90 percent separation occurs in the primary separator. This is followed 16 by an oil quench to 1300, and a cyclone to remove the remainder 17 of the solids. The mix is then quenched again with liquid to 18 600F. Thus, all the available heat from the reaction outlet 19 temperature to 600F is rejected to a circulating oil stream.
Steam is generated from the oil at 600 psig, 500F. This 21 scheme is used to avoid exchanger fouling when cracking heavy 22 feeds at low steam dilutions and high severities in the TRC~
23 ~owever, instead of an oil quench, a circulating solids stream 2~ could be used to quench the effluent. As in the reaction itself, the co~e would be deposited preferentially on the solids thus 26, avoiding fouling. These solids can be held at 800F or above, 27 thus permitting the generation of high pressure ~team (1500 psig~) 28 which increased the overall thermal efficiency of the process.
~9 The oil loop can not operate at these temperatures due to instabilities (too many light fractions are boiled off, yieldin~3 ~ -"
S&~
696-1~7 1 an oil that is too viscous). The use OL solids can be done for 2 both TRC or a coil, but it is especially suited to a T~C since 3 it already uses solids. ~uring quenchin~, the coke accumulates 4 on the solid. ~t must be burned off. In a coil application, it would have to be burned off in a separate vessel while in a 6 T~C it could use the regenerator that already exists.
7 ~A7ith reference to the preheat vaporization system of 8 the s-~ject invention ~see ~IGURE 29), it is noted that 9 the TRC has maxirum economic advantA~es when cracking heavy feedstocks (650F+ boiling ~oint) at low steam dilutions.
11 Selectivity is favored by rapid and intimate mi~ing. ~apid 12 and intirr~ate mixinq is best accomplished with a vapor feed 13 rather than a li~uid feed.
1~ Finally, with reference to the sequential crackin~
system of the invention (see FIGURE 14), it is clear that 16 sequential cracking represents an alternative way of utilizing 17 the heat available in the quench (as opposed to the solids 18 quench boiler) in addi-tion to any yield advantages. It can a be applied to both T~C and a coil. Its synergism with TRC
is that it permits the use of lon~er solids/gas separation times 21 if the second feed is added prior to any separation. The high 22 amount of heat available in the solids permits the use of lowe~
23 temperatures compared to the coil case.
2~ ~hile there has been described what is considered to be preferred embodiments of the invention, variations and ~odif-26 ications therein will occur to those skilled in the art once 2; they become acquainted with the basi_ concepts of the inventior 2~ Therefore, it is intended that the appended claims shall be '9 construed to include not only the disclosed embodiments but all such variations and modifications that fall within the true 31 soirit and scope of the invention.

Claims (70)

The embodiments of the invention in which an exclusive property or privilege is claimed are defined as follows:

1. In a TRC process wherein the temperature in the cracking zone is between 1300° and 2500°F. and wherein the hydrocarbon feed or the hydrosulfurization residual oil along with the entrained inert solids and the diluent gas are passed through the cracking zone for a residence time of 0.05 to 2 seconds, the improvements comprising:
(1) the process for generating fuel oil and removing coke deposits on said solids comprising the steps of:
(a) generating a fuel gas from fuel and air;
(b) delivering the fuel gas to a transfer line;
(c) mixing the particulate solids with the fuel gas in the transfer line to elevate the temperature of the solids; and (d) combusting the fuel gas in the transfer line to elevate the temperature of the solids and remove the coke from the solids, and (2) the process for separating by centrifugal force particulate solids from the dilute mixed phase stream of gas and solids comprising the steps of:
(a) adding the mixed phase stream to a chamber having a flow path of essentially rectangular cross section from an inlet of inside diameter Di disposed normal to the flow path, said flow path having a height H
equal to Di or 4 inches, whichever is greater, and a width W greater than or equal to 0.75 Di but less than or equal to 1.25 Di, (b) disengaging solids from gas by centrifugal force within said chamber along a bed of solids found at a wall opposite to the inlet as the gas flows through said flow path, the gas changing direction 180°, and the solids being projected 90° toward a solids outlet, (c) withdrawing the gaseous portion of the inlet stream from a gas outlet, disposed 180° from the inlet, the gas portion containing about 20% residual solids, said gas outlet located between the solids outlet and inlet, the gas outlet being at a distance no greater than 4 Di from the inlet as measured between respective centerlines, and (d) withdrawing the solids by gravity through the solids outlet.

2. In a TRC process as in Claim 1, the further improvement wherein the process for mixing the solids rapidly and intimately with the fluid feed in the reaction chamber mixing zone comprises the steps of a. delivering fluidized solids through a conduit to a mixing chamber; and b. introducing fluid feed into the stream of solids at an angle.

3. In a TRC process as in Claim 2, the further improvement wherein the step of cracking hydrocarbon feed to produce olefins comprises:
a. delivering hydrocarbon feed to a first zone;
b. thermally cracking the hydrocarbons in the first zone at temperatures above 1,500°F.;
c. discharging the cracked effluent from the first zone to a second zone;
d. delivering a second hydrocarbon feed to the entry of the second zone; and e. mixing the cracked effluent from the first zone and the second hydro-carbon feed in the second zone;
whereby the cracked effluent from the first zone is quenched and the second hydrocarbon feed is cracked at low severity.

4. In a TRC process as in Claim 1 further comprising the improvement wherein the process for quenching the reactor effluent comprises the steps of:
a. introducing particulate solids into the effluent stream; and b. passing the composite stream of effluent and particulate solids in indirect heat exchange relationship with a coolant.

5. In a TRC process as in Claim 2, the further improvement wherein the process for quenching the reactor effluent comprises the steps of:
a. introducing particulate solids into the effluent stream; and b. passing the composite stream of effluent and particulate solids in indirect heat exchange relationship with a coolant.

6. In a TRC process as in Claim 3, the further improvement wherein the process for quenching the reactor effluent comprises the steps of:
a. introducing particulate solids into the effluent stream; and b. passing the composite stream of effluent and particulate solids in indirect heat exchange relationship with a coolant.

7. In a TRC process as in Claim 1, the further improvement in the process for preheating the heavy oil hydro-carbon feedstock comprising the steps of:
a. heating the liquid heavy oil hydrocarbon feedstock;
b. initially flashing the heated liquid heavy oil hydrocarbon feedstock with steam;
c. separating the vapor and liquid phases of the flashed liquid heavy oil hydrocarbon feedstock-steam mixture;

d. superheating the vapor phase of the flashed liquid heavy oil hydrocarbon feedstock-steam mixture; and e. flashing the superheated vapor and the liquid phase of the originally flahsed liquid heavy oil hydrocarbon feedstock-steam mixture.

8. In a TRC process as in Claim 2, the further improvement in the process for preheating the heavy oil hydro-carbon feedstock comprising the steps of:
a. heating the liquid heavy oil hydrocarbon feedstock;
b. initially flashing the heating liquid heavy oil hydrocarbon feedstock with steam;
c. separating the vapor and liquid phases of the flashed liquid heavy oil hydrocarbon feedstock-steam mixture;
d. superheating the vapor phase of the flashed liquid heavy oil hydrocarbon feedstock-steam mixture; and e. flashing the superheated vapor and the liquid phase of the originally flashed liquid heavy oil hydrocarbon feedstock-steam mixture.

9. In a TRC process as in Claim 3, the further improvement in the process for preheating the heavy oil hydro-carbon feedstock comprising the steps of:
a. heating the liquid heavy oil hydrocarbon feedstock;
b. initially flashing the heated liquid heavy oil hydrocarbon feedstock with steam;
c. separating the vapor and liquid phases of the flashed liquid heavy oil hydrocarbon feedstock-steam mixture;
d. superheating the vapor phase of the flashed liquid heavy oil hydrocarbon feedstock-steam mixture; and e. flashing the superheated vapor and the liquid phase of the originally flashed liquid heavy oil hydrocarbon feedstock-steam mixture.

10. In a TRC process as in Claim 4, the further improvement in the process for preheating the heavy oil hydro-carbon feedstock comprising the steps of:
a. heating the liquid heavy oil hydrocarbon feedstock;
b. initially flashing the heated liquid heavy oil hydrocarbon feedstock with steam;
c. separating the vapor and liquid phases of the flashed liquid heavy oil hydrocarbon feedstock-steam mixture;
d. superheating the vapor phase of the flashed liquid heavy oil hydrocarbon feedstock-steam mixture; and e. flashing the superheated vapor and the liquid phase of the originally flashed liquid heavy oil hydrocarbon feedstock-steam mixture.

11. In a TRC apparatus wherein the temperature in the reaction chamber is between 1300° and 2500°F and wherein the hydrocarbon fluid feed or the hydrosulfurization residual fluid feed oil along with the entrained inert solids and the diluent gas are passed through the reaction chamber for a residence time of 0.05 to 2 seconds, the improvements:
(1) wherein the apparatus for admixing the inert solids rapidly and intimately with the fluid feed introduced simultaneously thereto comprises:
a. an upper reservoir containing the particulate solids;
b. a conduit extending downwardly from the reservoir to the reaction chamber, said conduit being in open communication with the reservoir and reaction chamber; and (2) a solids-gas separator designed to effect rapid removal of particulate solids from a dilute mixed phase stream of solids and gas, said separator comprising a chamber for disengaging solids from the incoming mixed phase stream, said chamber having rectilinear or slightly arcuate longitudinal walls to form a flow path essentially rectangular in cross section, said chamber also having a mixed phase inlet, a gas phase outlet, and a solids phase outlet, with the inlet at one end of the chamber disposed normal to the flow path, the solids outlet at the other end of the chamber, said solids outlet suitable for downflow of discharged solids by gravity, and the gas outlet therebetween oriented to effect a 180° change in direction of the gas.

12. In a TRC apparatus as in Claim 11, the further improvement comprising a system for heating and removing coke from the particulate solids comprising:
a. means for generating fuel gas having a high molal ratio of H2O to H2 from fuel, air and steam;
b. a transfer line; and c. means to mix the fuel gas and particulate solids in the transfer line, whereby the fuel gas elevates the temperature of the particulate solids by intimate contact therewith and the steam in the fuel gas removes the coke from the solids during travel through said transfer line.

13. In a TRC apparatus as in Claim 12, the further improvement wherein the apparatus for quenching effluent comprises:
a. an indirect heat exchanger formed for an outer wall of longitudinally extending tubes joined together to form a pressure-tight membrane wall;
b. a reactor effluent outlet tube extending into the heat exchanger in communication with the hot side of the heat exchanger;

c. means to deliver particulate solids into the effluent discharging from the reactor effluent outlet tube; and d. means to deliver steam to the tubes forming the outer wall of the heat exchanger.

14. In a TRC apparatus as in Claim 11, the further improvement of providing a system for preheating the heavy oil hydrocarbon feedstock comprising:
a. means for preheating the liquid heavy oil hydrocarbon;
b. a first mixer for flashing the heated liquid heavy oil hydrocarbon and steam;
c. a vapor feed superheater for heating the vapors from the first mixer to about 1,030°F.; and d. a second mixer for flashing the super-heated vapor and the liquid from the first mixer.

15. In a TRC apparatus as in Claim 12, the further improvement of providing a system for preheating the heavy oil hydrocarbon feedstock comprising:
a. means for preheating the liquid heavy oil hydrocarbon;
b. a first mixer for flashing the heated liquid heavy oil hydrocarbon and steam;
c. a vapor feed superheat for heating the vapors from the first mixer to about 1,030°F;
and d. a second mixer for flashing the super-heated vapor and the liquid from the first mixer.

16. In a TRC apparatus as in Claim 13, the further improvement of providing a system for preheating the heavy oil hydrocarbon feedstock comprising:
a. means for preheating the liquid heavy oil hydrocarbon;
b. a first mixer for flashing the heated liquid heavy oil hydrocarbon and steam;
c. a vapor feed superheater for heating the vapors from the first mixer to about 1,030°F.;
and d. a second mixer for flashing the super heated vapor and the liquid from the first mixer.

17. The process of Claim 1 further comprising the step of further separating residual solids from the gaseous portion of the inlet stream removed via the gas outlet in a secondary separator.

18. The process of Claims 1 or 17 further comprising the step of stripping solids withdrawn from the solids outlet with the inert gas or steam.

19. The apparatus of Claim 11 having a mixed phase inlet of inside diameter Di and which is further characterized by a flow path with a preferred height H equal to at least Di or 4 inches, whichever is greater and with a preferred width W between 0.75 and 1.25 Di, and having a gas outlet located between the mixed phase inlet and solids outlet at a preferred distance from the inlet which is no greater than 4.0 Di as measured between their respective centerlines.

20. The apparatus of Claim 19 wherein the most pre-ferred distance between inlet and gas outlet centerlines is no less than 1.5 Di but no greater than 2.5 Di.

21. The apparatus of Claim 20 wherein the most preferred height H of the flow path is twice Di.

22. The apparatus of Claim 21 wherein the most preferred width W of the flow path is no less than 0.9 Di but no greater than 1.10 Di.

23. The apparatus of Claim 19 having a solids removal outlet the first section of which is collinear with the flow path and the second section normal to the first section, and aligned for downflow of solids by gravity.

24. The apparatus of Claim 23 further comprised of a flow restriction placed within the colinear section of the solids removal outlet.

25. The apparatus of Claim 24 wherein the flow restriction is an orifice.

26. The apparatus of Claim 24 wherein the preferred flow restriction is a venturi.

27. The apparatus of Claim 11 or 19 further comprising a weir placed across the flow path beyond or at the gas outlet but before the solids outlet.

28. The apparatus of Claim 19, 23 or 24 having a chamber whose longitudinal dimension is extended beyond the gas outlet by a length L.

29. The apparatus of Claims 19, 23 or 24 having a chamber whose longitudinal dimension is extended beyond the gas outlet by a length L which is less than or equal to 5 Di.

30. The apparatus of Claim 11 wherein the fluid feed is introduced angularly to the flow of solids such that the projected flow of feed intercepts the discharge flow of solids leaving each conduit.

31. The apparatus of Claim 30 wherein the means for introducing aeration gas is a plurality of nozzles spaced symmetrically around the inlets of said conduits.

32. The apparatus of Claim 31 further comprising a header to supply aeration gas to the aeration nozzles; and wherein the fluid feed is a gas, the means for introducing said gaseous feed being annular orifices around each conduit, said apparatus further comprising a gaseous feed distribution chamber above the reaction chamber and in communication therewith through the annular orifices, the conduits passing through said distri-bution chamber and terminating coplanarly with the base thereof, said base of the distribution chamber having holes therein, the holes receiving the conduit and being larger than the outside dimension of the conduits forming the annular orifices; said gaseous feed to the distribution chamber being supplied by a header; and wherein said apparatus further comprises a section of the reaction chamber, said section being in open communica-tion with the conduits and constituting a mixing zone for the gaseous and solid feed introduced thereto; with a plug extending downwardly from the base of the distribution chamber into the mixing zone to form discrete mixing zones.

33. The apparatus of Claim 32 wherein the base of the distribution chamber is a removable plate having holes larger than the outer dimension of and receiving the conduits, said conduits having convergently beveled outside walls as the outlet end which terminate coplanarly with said plate and within the holes forming thereby the annular orifices, the holes further having convergently beveled edges whereby the direction of gas flow is angled toward the conduit outlet, the projection of which forms a cone the vertex of which is beneath said conduit outlet.

34. The process as in Claim 1 further comprising the steps of passing the composite quenched effluent from the second zone through the hot side of an indirect heat exchanger and passing steam through the cold side of the indirect heat exchanger.

35. The process as in Claim 1 further comprising the steps of fractionating the cracked effluent and returning a portion of the fractionated cracked effluent to the first zone.

36. A process as in Claim 1 wherein the first zone is operated at high severity short residence cracking conditions.

37. A process as in Claim 1 wherein the feed delivered to the second zone is virgin gas oil 400° to 650°F.

38. A process as in Claim 35 wherein the fraction returned to the first zone is light paraffinic gases of ethane and propane.

39. A process as in Claim 1 wherein the hydrocarbon delivered to the first zone is pre-heated to a temperature between 600° and 1,200°F.

40. A process as in Claim 1 wherein the hydrocarbon delivered to the second zone is pre-heated to a temperature between 600° and 1,200°F.

41. A process as in Claim 36 wherein the kinetic severity function in the first zone is about 3.5.

42. A process as in Claim 36 wherein the kinetic severity factor is about 0.5 at about 300 to 400 milliseconds.

43. A process as in Claim 1 wherein 100 pounds of hydrocarbon are delivered to the second reaction zone as quench for every 60 pounds of effluent from the primary zone.

44. An apparatus as in Claim 11 further comprising a diverging cone at the termination of the riser tube terminating at the entry end of the hot side of the indirect heat exchanger.

45. An apparatus as in Claim 44 wherein the diver-gent cone on the riser tube is at an angle to the axis of the riser tube of less than 10°.

46. An apparatus as in Claim 11 further comprising:
a circulary array of tubes joined together to form a pressure-tight outer surface for the indirect heat exchanger;
a first torus to which one end of the circular array of tubes are manifolded;
a second torus to which the other end of the circular array of tubes are manifolded; and means for delivering coolant to the first torus.

47. An apparatus as in Claim 46 wherein the additional inner cooling tubes are arranged in two concentric circles,

48. An apparatus as in Claim 11 further comprising a diverging cone at the lower end of the riser tube and a converg-ing cone formed on the outer surface of the reactor outlet tube.

49. The process as in Claim 1 wherein the coolant is steam and comprising the further step of generating high pressure steam from the coolant during the heat exchange with effluent and particulate solids.

50. The process as in Claim 1 wherein the ratio of particulate solids introduced into the effluent stream to the gas in the stream is 25 to 50 pounds of solid per pound of gas.

51. The process as in Claim 49 further comprising the steps of separating the quenched cracked gas from the particulate solids and returning the particulate solids to the heat exchanger hot side.

52. The process as in Claim 1 wherein the initial flashing of the steam and the liquid heavy oil hydrocarbon is at a temperature of 500° to 900°F., the vapor from the initial flashing is superheated to about 1,100°F. and the superheated vapor and liquid from the initial flashing step is again flashed to about 1,000°F.

53. The process as in Claim 52 wherein the liquid heavy oil is preheated to 440° to 700°F.

54. The process as in Claim 1 wherein the particulate solids are decoked by the passage with the fuel gas in a vessel at about 100 feet per second and the steam decoking reaction reduces the particulate solids-fuel gas temperature.

55. The process as in Claim 1 further comprising the step of combusting the fuel gas in the transfer line to further heat the solids and remove the coke from the solids.

56. The apparatus as in Claim 11 further comprising means to separate the fuel gas from the heated particulate solids after the solids have been heated and cleaned of coke.

57. In a TRC apparatus wherein the temperature in the reaction chamber craking zone is between 1300° and 2500°F and wherein the hydrocarbon fluid feed or the hydro-sulfurized residual oil fluid feed along with the entrained inert particulate solids and the diluent gas are passed through the cracking zone for a residence time of 0.05 to 2 seconds, the improvement comprising a solids-gas separator designed to effect rapid removal of particualte solids from a dilute mixed phase stream of solids and gas, said separator comprising a chamber for disengaging solids from the incoming mixed phase stream, said chamber having rectilinear or slightly arcuate longitudinal walls to form a flow path essentially rectangular in cross section, said chamber also having a mixed phase inlet, a gas phase outlet, and a solids phase outlet, with the inlet at one end of the chamber disposed normal to the flow path, the solids outlet at the other end of the chamber, said solids outlet suitable for downflow of discharged solids by gravity, and the gas outlet therebetween oriented to effect a 180° change in direction of the gas.

58. The separator of Claim 57 having a mixed phase inlet of inside diameter Di and which is further characterized by a flow path with a preferred height H equal to at least Di or 4 inches, whichever is greater and with a preferred width W between 0.75 and 1.25 Di, and having a gas outlet located between the mixed phase inlet and solids outlet at a preferred distance from the inlet which is no greater than 4.0 Di as measured between their respective centerlines.

59. The separator of Claim 57 having a solids removal outlet the first section of which is collinear with the flow path and the second section normal to the first section, and aligned for downflow of solids by gravity.

60. The separator of Claim 59 further comprised of a flow restriction placed within the collinear section of the solids removal outlet.

61. The separator of Claim 60 wherein the flow restriction is an orifice.

62. The separator of Claim 60 wherein the preferred flow restriction is a venturi.

63. The separator of Claims 57 or 58 further com-prising a weir placed across the flow path beyond or at the gas outlet but before the solids outlet.

64. In a TRC apparatus wherein the temperature in the reaction chamber is between 1300° and 2500°F and wherein the hydrocarbon feed or the hydrosulfurized residual oil along with the entrained inert particulate solids and the diluent gas are passed through the reaction zone for a residence time of 0.05 to 2 seconds, the improvement comprising:
a solids-gas separation system to separate a dilute mixed phase stream of gas and particulate solids into an essentially solids free gas stream, the separation system comprising:
a chamber for rapidly disengaging about 80% of the particulate solids from the incoming dilute mixed phase stream, said chamber having approximately rectilinear or slightly arcuate longitudinal side walls to form a flow path of height H and width W approximately rectangular in cross section, said chamber also having a mixed phase inlet of inside width Di, a gas outlet, and a solids outlet, said inlet at one end of the chamber disposed normal to the flow path whose height H is equal to at least Di or 4 inches, whichever is greater and whose width W is no less than 0.75 Di but no more than 1.25 Di, said solids outlet at the opposite end of the chamber and being suitable for downflow of discharged solids by gravity, and said gas outlet therebetween at a distance no greater than 4 Di from the inlet as measured between respective centerlines and oriented to effect a 180° change in direction of the gas whereby resultant centrifugal forces direct the solid particles in the incoming stream toward a wall of the chamber opposite to the inlet forming thereat and maintaining an essentially static bed of solids, the surface of the bed defining a curvilinear path of approximately 90° for the outflow of solids to the solids outlet, a secondary solids-gas separator, said secondary separator removing essentially all of the residual solids, a first conduit connecting the gas outlet from the chamber to the secondary separator, a vessel for the discharge of solids, a second conduit connecting said vessel and the chamber, and, pressure balance means to maintain a height of solids in said second conduit to provide a positive seal between the chamber and vessel.

65. The separation system of Claim 64 wherein the pressure balance means is the hydraulic forces acting on the chamber, second conduit and vessel, the second conduit being sized for sufficient pressure loss to provide the height of solids.

66. The separation system of Claim 64 wherein the pressure balance means is a check valve at the outlet end of the second conduit.

67. The separation system of Claim 64 wherein the pressure balance means is a pressure control valve on the solids discharge vessel.

68. In a TRC process wherein the temperature in the reaction chamber is between 1300° and 2500°F and wherein the hydrocarbon fluid feed or the hydrosulfurization residual oil along with the entrained inert solids and the diluent gas are passed through the reaction chamber for a residence time of 0.05 to 2 seconds, the improvement comprising a method for separating by centrifugal force particulate solids from a dilute mixed phase stream of gas and solids, the method com-prising the steps of:
adding the mixed phase stream to a chamber having a flow path of essentially rectangular cross section from an inlet of inside diameter Di disposed normal to the flow path, said flow path having a height H
equal to Di or 4 inches, whichever is greater, and a width W greater than or equal to 0.75 Di but less than or equal to 1.25 Di, disengaging solids from gas by centrifugal force within said chamber along a bed of solids found at a wall opposite to the inlet as the gas flows through said flow path, the gas changing direction 180°, and the solids being projected 90° toward a solids outlet, withdrawing the gaseous portion of the inlet stream from a gas outlet, disposed 180°
from the inlet, the gas portion containing about 20% residual solids, said gas outlet located between the solids outlet and inlet, the gas outlet being at a distance no greater than 4 Di from the inlet as measured between respective centerlines, and withdrawing the solids by gravity through the solids outlet.

69. The method of Claim 68 further comprising the step of further separating residual solids from the gaseous portion of the inlet stream removed via the gas outlet in a secondary separator.

70. The method of Claims 68 or 69 further comprising the step of stripping solids withdrawn from the solids outlet with inert gas or steam.