CA1210726A - Pre-heat vaporization system - 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

lZ~7~6 This ~lpplicatioll is a divis.ion of co-pending applica-tion Serial No. 361,734 filed September 30, 19~0.

The present invention relates to lrnprovernents in Thermal Regenerative Cracking (TRC) apparatus and proces.s, as described in U.S. Letters Patent Nos. 4,061,562 and ~,097,363 to McKinney et al.

According to one broad aspect, the presen-t invention relates to a TRC process wherein the temperature i.n the cracking zone is between 1300~ and 2500F and wherein hydrosulfurized residual oil along with the entrained inert solids and the diluent gas are passed through a cracking zone for a residence time of 0.05 to 2 seconds, the improvement in the process for preheating the heavy oil hydrocarbon feedstock comprising the steps of heating the liquid heavy oil hydrocarbon feedstock; initially flashing the heated liquid heavy oil hydrocarbon feedstock with steam; separating the vapor and liquid phases of the flashed liquid heavy oil hydrocarbon feedstock-steam mixture; superheating the vapor phase of the flashed liquid heavy oil hydrocarbon feedstock-steam mixture; and flashing the superheated vapor and the liquid phase of the originally flashed liquid heavy oil hydrocarbon feedstock-steam mixture.

According to another broad aspect, the present invention relates to a TRC system wherein the temperature in the cracking zone is between 1300 and 2500F and wherein the hydrosulfurized residual oil along with the entrained inert solids and the diluent gas are passed ~hrough a cracking zone for a residence time of 0.05 to 2 seconds, the improvement of 72~

providiny a sys-tern for prehea-ting -the heavy oil hydrocarbon feedstock comprising means for preheating the liquid heavy oil hydrocarbon; a Eirst mixer for Elashing the heated liquid heavy oil hydrocarbon and steam; vapor feed superheater Eor heating the vapors from the first mixer to about 1,030F.; and a second mixer for flashing the superheated vapor and the liquid from the first mixer.

According to a s-till further broad aspect, the present invention relates to a process for pre-heating heavy oil hydrocarbon feedstock comprising heating the liquid heavy oil hydrocarbon feedstock; initially flashing the heated liquid heavy oil hydrocarbon feedstock with steam; separating the vapor and liquid phases of the flashed liquid heavy oil hydrocarbon feedstock-steam mixture; superheating the vapor phase of the flashed liquid heavy oil hydrocarbon feedstock-steam mixture; and flashing the superheated vapor and the liquid phase of the originally flashed liquid heavy oil hydrocarbon feedstock-steam mixture.

According to yet another broad aspect, the present invention relates to a system for pre-heating heavy oil hydrocarbon feedstock comprising means for preheating the liquid heavy oil hydrocarbon; a first mixer for flashing the heated liquid heavy oil hydrocarbon and s-team; a vapor feed superheater for heating the vapors from the first mixer to about 1,030~F.; and a second mixer for flashing the superheated vapor and the liquid from the first mixer.

The invention will be more fully appreciated by referring to the accompanying drawings in which:

-2a-, ~
j, ;;
. , , ~L2~(-7~6 i Figure l is a schematic ~iagram of cl Tr~c system and process according to the prior art.
Figure 2 is a schematic dia~rarn of the Luel gas generation syste~m and process of the subject invention.
Figure 3 is an alternative embodiment wherein the fuel gas is burned to flue gas to provide additional heat for the 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 5 is an enlarged view of the intersection of the solid and gas phases within the mixing zone of the reaction chamber.
Figure ~ is a top view of the preferred plate geometry, said plate serving as the base of the gas distribution chamber.
Figure 7 is a graph of the relationship between bed density, pressure drop, bed height and aeration 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 lO is an alternate preferred embodiment of the control features of the present invention.
Figure ll is a view along line ll-ll of Figure lO
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.

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7;26 (S~
6 9 6 - l 4 7 1 FIGURE 13 is a view of the apparatus of FIGVRE 12

2 through line 13-13 of FlGURE 12.

3 FIGVRE 14 is a schematic diagram of the sequential

4 thermal cracking process and system of the present invention.
FIGURE 15 is a schematic flow diagram of the 6 separation svstem of the present invention as ap~ended to a 7 typical tubular reactor.
8 FIGVRE 16 is a cross sectional elevational view of g the preferred e~bodiment of the separator.
FIGURE 17 is a cutaway view throu~h section 17-17 11 of FIGURE 16.
12 FIGURE 18 is a cutaway view through section 18-18 13 of FIGURE 16 showing an alternate geometric configuration of 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 separation cha~ber.
21 FIGURE 21 is a sketch of an alternate er.lbodiment 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 24 as a gravity flow section.
FIGURE 22 is a variation of ~he embodiment of FIGUiRE
26 21 in which the solids outlet of FIGVRE 20 is used, but is not 27 stepped.
28 FIGURE 23 is a sketch of a variation of the 29 separation device of FIGVRE 8 wherein a venturi restriction is incorporated in the collinear section of the solids outlet.

_4_ (S&W) 1 FIG~RE 24 is a variation of the embodi~ent of 2 FIGURE 23 oriented for use with a riser type reactor.
3 FIGURE 25 is a sectional elevat-onal vie~ of the 4 solids quench boiler using the quench riser;
FIGURE 26 is a detailed cross sectional elevatior.al 6 view of the quench exchanger of the system;
7 E`IGURE 27 is a cross sectional plan view taken 8 through line 27-27 of FIG~RE 26;
9 FIG~RE 28 i,s a detailed drawing of the reactor outlet and fluid bed quench riser particle entry area.
11 FIGURE 29 is a schematic diagram of the system of 12 the invention for vaporizing heavy oil.

17 - ~
18 The improvements of the subject invention are 19 em~odied in the environment of a thermal regeneration cracking reactor (TRC) which is illustrated in FIGURE 1.

22 ~,, Referring to FIG~RE 1, in the prior art TRC process~
23 jiand system, thermal cracker feed oil or residual oil, with or 24 I~ithout blended distillate heavy gas, entering through line 10 'and hydrogen entering through line 12 pass through hydrodesulfurized 26 jlzone 14. ~ydrosulfurization effluent passes through line 16 and 27 ~enters flash chamber 19 from which hydrogen and contaminating gases 28 llncluding hydrogen sulfide and ammonia are removed overhead through 29 line 20, while flash liquid is removed through line 22. The flash 1l _5_ S&W ~Z~7~6 I . , i 1 l liquld passes throu~h preheater 24, is admixed with dilution 2 1 steam entering through line 26 and then flows to the bottom 3 1 of thermal cracking reactor 28 through line 30.
4 I, A stream of hot regenerated solids is charged ' through line 32 and admixed with steam or other fluidizing gas 6 ,1 entering through line 34 prior to entering the bottom of riser 7 il 28. The oil, steam and hot solids pass in entrained flow up-8 I wardly through riser 28 and are discharged through a curved 9 ,, segr,lent 36 at the top of the riser to induce centrifugal separ-1 ation of solids from the effluent stream. A stream containing most of the solids passes through riser discharge segment 38 and 12 I can be mixed, if desired, with make-up solids entering through 13 ' line 40 before or after entering solids separator-stripper 42.
14 Another stream containing most of the cracked product is dis-!I 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 48.
18 I' Stripper steam is charged to solids separators 42 19 l and 48 through lines 50 and 52, respectively. Product streams 20 ll are removed from solids separators 42 and 48 through lines 54 21 1! and 56, respectively, and then combined in line 58 for passage 2 2 'li to a secondary quench and product recovery train, not shown.
23 .ll Coke-laden solids are removed from solids separators 42 and 48 24 ~ll through lines 60 and 62, respectively, and combined in line 64 ll for p~ssage to coke bur~er 66. ~f re~uired, torch oil can be 26 1l added to burner 66 through line 68 while stripping steam may be ~7 l¦ added through line 70 to strip combustion gases from the heated 2~3 ll solids. Air is charged to the burner through line 690 Combustio~, 29 '¦ gases are removed from the burner through line 7~ for passage I to heat and energy recovery systems, not shown, while regenerated ,l -6 z~
S&W
69~1~7 ot solids which are relatively free of coke are removed from 2 the burner tnrough line 32 for recycle to ~iser 28. In Grder 3 1 to produce a cracked product containing ethylene and ~ole~ular 4 I 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 chemically combined with theoil during the hydrocycling step.
8 The hydrosulfurization residual oil passes through the tnermal g ' cracking zone together with the entrained inert hot solids , functioning as the heat source and a diluent gas at a tempera~ure between about 1300F and 2500F for a residual time between 12 1 about 0.05 to 2 seconds to produce the cracked product and 13 i ethylene and hydrogen. For the production of ethylene by 14 , thermally crackinq a hydrogen reed at least 90 volume percent ' of which comprises light gas oil fraction of a crude oil 16 boiling between 400F and 650F, the hydrocarbon feed, along 17 I with diluent gas and entrained inert hot gases are passed 1~ , through the cracking zone at a tem?erature between 1300F and 19 ~ 2500F 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 crackl,.y 21 j severity corresponds to a methane yield of at least 12 weight 22 I percent based on said feed oil. Quench cooling of the product 23 ~l immediately upon leaving the cracked zone to a temperature 24 I below 1300F ensures that the ethylene yield is greater than ~25 l the methane yield on a weight basis.

28 ~
29 ~ `

, _ , S&W ~L21~7Z6 1 (a) Improved ~uel Gas Generation For Solids 2 Heating.

4 I FIGURE 2 illustrates the improved process and system of I the invention as may be embodied in a prior art TRC system, in lieu 6 ,of the coke burner 66 (FIo~ 1). Particulate solids and hydrocarbon feed gas 7 1 enter a tubular reactor 13A throuah ~nes llA and 12A re~ectively. The cracked 8 I effluent from the tubular reactor 13A is separated fr~m the partic~ate soiids 9 in a separator 14A and quenched in line 15A by quench material injected from line 17A. The solids separated from 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 I free product gas is taken overhead from the secondary separator l18A through line l9A.
16 ` -~
17 \ - -S&~

., The particulate solids in the solid stripper 22A, 2 `having delivered heat during the thermal cracking in the tubular 3 , reactor 13A, must be reheated and returned to the tubular reactor 4 , 13A to continue the crac~ing process.
' The particulate solids prior to being reheated, are 6 strlpped 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 ot 9 gas impurities in the solid stripper 22A, ~he particulates soli~s are at a temperature of about 1,450F.
11 ~ The fuel gas generation apparatus of the invention 12 consists of a combustion vessel 30A, and pre-heat equipment for 13 I fuel, air (or 2) and steam which are delivered to the combus~ion 14 vessel 30A. Pre-heaters 32A, 34A, and 36 are shown in fuel llne 1 38A, air line 40A, and steam line 42A respectively.
16 ` rllhe system also includes a transfer line 44A intc 17 , which the combusted fuel gas from the combustion vessel 30A and 18 the stripped particulate solids from the solid stripper 22A are 19 mixed to heat and decoke the particulate solids. The transfer 2~ ,~ line 44A is sized to afford sufficient residence time for the 21 ; steam emanating from the combustion vessel 30A to decompose bv 22 lil the reactio~ with carbon in the presence of hydrogen and to remove 23 1I the net carbon from the solids-gas mixture. In the preferred 24 l; embodiment the transfer line 44 will be about 100 feet long~ A
25 ~ e 26A is provided for pneumatic transport gas if necessary.
26 ll A separator, such as a c~-clone separator 46A is 27 ~ provided to separate the heated decoked particulate solids from 28 !I the fuel gas. The particulate solids from the separator 46A are 29 ,I returned through line 48A to the hot solids hold vessel 27A
, and the fuel gas is taken overhead through line 50A.

_9_ 1.`

S &W , ~Z1~7Z6 1 In the process, fuel, air and steam are delivered 2 through lines 38A, 40A and 42~ respectively to the combustion 3 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 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 with 8 carbon in the presence of hydrogen and to insure that the net carbon in the fuel gas-particulate solids will mix will be ~ removed before reaching the separator 46A.
ll ~ The 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 15 equilibrium temperature of 1,780F. and continue to pass through 16 the tubular vessel 44A. During the passage through the tubular 17 vessel 44A the particulate solid-fuel gas mixture provide the 18 , heat neeessary to react the net coke in the mixture with steam.
19 l, 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 separa~ed 22 in the separator 46A and the fuel gas is taken at l,750F. through 23 ' line SOA~ The particulate solids are delivered to the hot solias 24 1' hold vessel 27A at 1,750F. and then to the tubular reactor 13A.
,25 In the alternative embodiment of the invention 26 l 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 gas. The fuel gas at 2,300F. and particulate 29 l~ solids at about 1, 450F. are mixed in the transfer line 44A to a temperature of about 1,486F. Thereafter air is delivered l to the transfer line 44A through a line 54A. The fuel gas in 7~Z6 S&W

1 l the line 44A is burned to elevate the temperature of the particu-2 1~ late solids to about 1,750F. The resultant flue gas is separated 3 l,from the hot solids in the separator 46A and discharged through 4 I the line 52A. The hot particulate solids are returned to the ~,system to provide reaction heat.
6 il An example of the system and process of FIGURE 3 7 follows: 7,000 pounds per hour of fuel pre-heated to 600F. in 8 ; 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 ~1 SCFD
' of fuel gas.
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 particu-13 "late solids from the solids stripper 22A. The particulate solids 14 have 1,600 pounds per hour of carbon deposited thereon. The cGm-Iposite fuel gas-particulate solids gas mixture reaches an 16 lequilibrium temperature of 1,480F. at 5 psig in about 5 milli-17 ,Iseconds. Thereafter, 13 ~ SCFD of air is delivered to the 18 ll, transfer line 44A and the 15.6 l~M SCFD of fuel gas is burned 19 Iwith the air to elevate the solids temperature to i~750F. and 20 ll burn the 1,600 pounds per hour of carbon from the particulate 21 ,Isolids, 22 ¦I The combusted ~as from the transfer line 44A is 2~ 1I separated from the solids in the separator 46A and discharged 24 ¦l as flue gas.

26 !~

28 jl \

.1 .

iZ~726 tb) 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 fluidizing gas contac-t 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 below. The housing 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 flageably connected via flanges 35B 36B, 37B and 38B. The particular geometry of the housing is functional rather than critical. 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~

S&W ~ 7Z6 ~96-147 I;
Set into and supported by the inner core 40B is a 2 gas distribution chamber 42B~ sald 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 separatiny 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 base 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 t~

11 ! the reaction zone 44B~ The conduits 46B are supported by the 12 material of the core 40B~ and terminate coplanarly with a plate 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 53 which is also part of the reaction zone 16 44~
17 ~ As shown in FIGllRE 5, an enlarged partial vlew of the 18 intersection of the conduit 46B and the plate 45B~ the apertures 19 47B are larger than the outside dimension of conduits 46B~ forming I therebetween annular orifices 48B for the passage of gaseous reed 21 ~ from the chamber 42B~ Edges 49B of the apertures 47B are pre-22 ferably convergently beveled, as are the edges 50B~ at the tip 23 ~ of the conduit wall 51Br In this way the gaseous stream from 24 1I the chamber 42B is angularly injected into the mixing zone 53B
1l and intercepts thesolids phase flowing from conduits 46~3~ A
26 l projection of the gas flow would form a cone shown by dotted lines 27 ,i 52B the vertex of which is beneath the flow path of the solids.
28 ' By introducing the gas phase angularly, the two phases are mixed 29 ll rapidly and uniformly, and form a homogeneous reaction phase.
~ The mixing of a solid phase with a gaseous phase is a function of i:

, s~w ~ ` lZ1~726 1 the shear surface between the solids and ~as phases, and the 2 flow area. A ratio of shear surface to flow area (S/A) of 3 ; infinity defines perfect mixin~; poorest mixing occurs when 4 i the solids are introduced at the wall of the reaction zone. In the system of the present invention, the gas stream is intro-6 duced annularly to the solids which ensures high shear surface.
7 By also adding the gas phase transversely through an annular 8 feed means, as in the preferred embodiment, penetration of the 9 phases is obtained and even faster mixing results. sy using a plurality of annular gas feed points and a plurality of solid 11 feed conduits, even greater mixing is more rapidly promo~ed, 12 since the surface to area ratio for a constant solids flow area 13 is increased. Mixing is also a known function of the L/D of lq the mixing 70ne. A plug creates an effectively reduced diameter D in a constant L, thus increasing mixing.
16 The Plug 54B, which extends downwardly from plate 17 ~5B, as shown in FIG~RES 4 and 5, reduces the flow area, and 18 forms discrete mixing zones 53B. The combination of annular gas 19 addition around each solids feed point and a confined discrete , mixing zone greatly enhances the conditions for mixing. Using 21 ~l this preferred embodiment, the time required to obtain an 22 1 essentially homogeneous reaction phase in the reaction zone 23 1 44B is ~uite low. Thus, this preferred method of gas and sollds 24 ' addition can be used in reaction systems having a residence time ,25 il below 1 second, and even below 100 milliseconds. In such re-26 actions the mixing step must be performed in a fraction of the 27 l, total residence time, generally under 20~ thereof. If this 28 criteria is not achieved, localized and uncontrolled reactlon 29 occurs which deleteriously affects the product yield and dis-,! tribution. This is caused by the maldistribution of solids , i ~4~

~ZlC~7Z6 s~w ,.
69~-,47 . ~ .

1 I normal to the flow through the reaction zone 44B thereby creating 2 ternperature and or concentration gradients therein.
3 The flow area is further reduced b~ placing the 4 i apertures 47B 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 apertures 47B symmetrically spaced along the 7 circumference. The plug 54B, shown by the dotted lines and 8 in FIGURE 9, 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 FIGVRE 5. In order to 12 prevent movement of conduits 46B by vibration and to retain the 13 uniform width of the annular orifices 48B, spacers 56B, are 14 ~ used as shown in FIGURE 8. However, the conduits 46B are pri-marily supported within the housing 34B by the material of the 16 ; core 4OB as stated above.
Referring to FIGURE 9, the plug 54B serves to la reduce the flow area and define discrete mixing zones 53B.
19 The plug 54B may also be convergently tapered so that there ~ is a gradual increase in the flow area of the mixing zone 53B
21 until the mixing zone merges with remainder of the reaction 22 i zone 44B. Alternatively, a plurality of plugs 54B can be used 23 l to obtain a mlxing zone 53B of the desired geometric con-24 1 figuration.
1' Referring again to FIGURE 4, the housing 34B may 26 l preferably contain a neck portion 57B with corresponding lining 27 '' 58B of the castable ceramic material and a flange 37B to coop-rate 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 53Bt 'j ' -15-6s6-I47 ,' ~Z1~7~6 I' .

1 , and allows complete removal of the housing 34B without dis-2 I assembly of the reactor 31B or the solids reservoir 33B. Thus, 3 installation, removal and maintenance can be accompllshed 4 ' easily. Ceramic linings 60B and 62B on the reservoir 33s and the reactor walls 61B respectively are provided to preven~
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 10 6~B sy~metrically placed around the conduits 46B. Gas to 11 I nozzles 64B is supplied by a header 65B. Preferably, the header 12 65B is set within the castable material of the core 40B, but 13 , this is dependent on whether there is sufficient space in the 14 ' housing 34B. A large mesh screen 66B is placed over the inlets ~ of the conduit 64B to prevent debris and large particles from 16 entering the reaction zone 44B or blocking the passage of the 1~ ; particulate solids through the conduits 46B.
lB By locally fluidizing the solids 63B, the solids 19 ~ 63B assume the characteristics of a fluid, and wi~l flow through j the conduits 46B. The conduits 46B have a fixed cross sectional 21 ll area, and serve as orifices having a specific r~sponse to a 22 change in orifice pressure drop. Generally, the flow of 23 ;, fluidized solids through an orifice is a function of the pressure 24 1l drop through the orifice. That orifice pressure drop, in tu~nf ~ is a function of bed height, bed density, and system pressure~
2~ I However, in the process and apparatus of this 27 I invehtion the bulk of the solids in reservoir 33B are not 2a I fluidized. Thusl static pressure changes caused by variations 29 1 in bed height are only slowly communicated to the inlet of the ¦ conduit 46B. Also the bed density remains approximately constant !l -16-I
,, .

696-147 ~Z~7Z6 ,1 .

1 li until the point of incipient fluidization is reached, that is, 2 point "a" of FIG~RE 7. In the present invention, however, it 3 ' is essential that the amount of aeration gas be below that 4 , amount. Any aeration gas flow above that at point "a" on I FIGURE 7 will effectively provide a fluidized bed and thereby 6 lose the benefits of this invention. By adjustment o~ the 7 aeration gas flow rate, the pressure drop across the non-8 fluidized bed can be varied. Accordingly, the pressure drop 9 across the orifice is regulated and the flow of solids thereby regulated as shown in FIGURE 7. As gas flow rates below 11 incipient fluidization, 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 density 14 variations have on mass flow are dampened considerably by the presence of the non-fluidized reservoir solids and are essentlally 16 eliminated as a significant factor. Further the control provided 17 by this invention affords rapid response to changes in solids 18 mass flow regardless of the cause.
~ ' ' lg Together with the rapid mixing features described above, the present invention offers an integrated system for 21 !l feedin~ particulate solids to a reactor or vessel, especially i! ~
22 ~ to a TRC tubular reactor wherein very low reaction residence 23 I times are encou~texed.
24 , FIGURES loandll depict an alternate preferred embodi-25 1l ment of the control features of the present invention. In ~hls 26 l embodiment the reservoir 33B extends downwardly into the core 27 I material 40B to form a secondary or control reservoir 71B. The 28 ~i screen 66B is positioned over the entire control res~rvoir 71B.
29 1I The aeration nozzles 64B project downwardly to fluidize essentially l' these solids 63B beneath the scxeen 66B. The bottom 41B of the 1, .

~.Z~(~726 S&W
696-1~7 1 reservoir 33B is again preferably formed of the same material 2 1 as the core 40B.
3 A plurality of clean out nozzles 72B are preferably 4 ~ provided to allow for an intermittent aeration gas dischar~e ~ 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 7413 provide the gas 8 ` supply to nozzles 64B and 72B respectively.
9 ~ The conduits46B communicate with the reservoir 71B
I through leading section 46'B~ The leading sections 46'B 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 75Bto 14 I each section 46'B can be sloped to allow solids to enter more easily. In addition to bein~ erosion resistent, the block 16 75B provides greater longevity because erosion may occur ~ithout 17 loss of the preset response function. Thus, even if the conduit 18 leading sections 46'B erode as depicted by dotted lines 77B, 19 the remaining leading section 46'Bwill still provide a known ~ orifice size and pressure drop response. The conduits 46B
21 l,l are completed as before usin~ erosion resistent metal tubes 22 l' 51B, said tubes being set lnto core material 40B and affixed 23 1I to the block 75B.
24 I FIGURE ilis a plan view of FIGURE ~o along section .25 1, 9-9 showing an arrangement for the nozzles 64B and 72B, and the 26 1l headers 65B and 74B. Gasis supplied to the headers 65B and 74B
27 j 'chrough feed lines 79~ and 80B respectively, which extend out 2B ~ beyond the shell 34B. It is not necessary that the headers be 2g I set into the material of the core 40B, although this is a ~0 1~ convenience from the fabrication standpoint. Uniform flow 1., 696 147 l ~2~726 i j distribution to each of the nozzles is ensured by the hydraulics 2 of the nozzles themselves, and does not re~uire other devices 3 such as an orifice or venturi. The gas supplied to ~eed lines 4 1 79B and 80B is regulated via valve means not shown.
FIGURES 12 and 13show the pertinent parts of an 6 alternate 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 9 orifice 48B surrounds each solids delivery conduit 46B. However, rather than a comlmon ~all between the chamber 47B and the conciuit 11 ;46B, a second annulus 83B is formed between the chamber 42s 12 and the conduit 46B. Walls 81B and 51B define the chambers 13 83B. Feed is introduced through both the annular opening 48B
14 in the chamber 4 2B and the annular opening 84B in the annulus ;83B at an angle to the flow of solids from the conduits 46P~.
6 The angular entry of the feed gas to the mixing zone 53Bis 17 I provided by beveled walls 49B and 85B, which define the openings 18 ~, 4aB and beveled walls 50B and 89B which define the openings 19 ! 84B. Gasis introduced to the annulus 83B through the header ~l86B, the header being set into ~he core 40B if conv2nient~
21 FIGURE i2 is a plan view of the apparatus of FIGV~;
22 1l 13 through section 11-11 showing the conduit openings and the 23 1~ annular feed openings 48B and 84B. Gas is supplied through feed 24 ! lines 87B and B8B to the headers 43B and 86B and ultimately 25 l, to the mixing zones through the annular openings. Vniform fIow 26 ¦ from the chambers 42B and 83Bis ensured by the annular orifices 27 1, 48B alld 84B. Therefore, it is not esser.tial that flow dis-2a l tribution means such as venturis or orifices be included in 29 1I the header 43B. The plug 54B is shaped symmetrically to 30 1I define discrete mixing zones 53B.
I~, ol9 _ :, S&W ~ t~6 696-147ll !
1, , 1 IMixing efficiency is also dependent upon the velocities 2 Of the gas and solid phases. The solids flow through 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 ll1000 pounds/sq. ft./sec., may be used depending on the character 6 listics 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 ft./
12 sec. 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.
15 iThe initial superficial velocity of the two phases in 16the mixing ~one 53B is preferably about 20 to 80 ft./sec., 17 'although this velocity changes rapidly in many reaction systems, 18 Isuch as thermal cracking, as the gaseous reaction products are 19 formed. The actual average velocity through the mixing ~one '`3B
1 and the reaction zone 44B is a process consideration, the velocity 21 I being a function of the allowed residence time therethrough.
22 !I By employing the solid feed device and method of the 23 ¦llpresent inventions, the mixing length to diameter ratio necessary 24 !¦to intimately mix the two phases is greatly reduced. This ratio l is used as an informal criteria which defines good mixing. Gen-26 I,erally, an L/D~lenqth/dia.) ratio of from 10 to 40 is required.
27 li Using the device disclosed herein, this ratio is less than 5, with 28 Iratios less than 1.0 being possible. Well designed mixing devices 2~ If the present invention may even achieve essentially complete 30 ' mixing at L/D ratios less than 0. 5.
, I
, -20-.

~&~1 ~96- 7 ~Z1~7~6 l (c) Improved Sequential Thermal Cracking 2 Process.

4 ,1 .
Turnlng now to the se~uentlal cracking process 2C
of the subject invention, as lllustrated in FIGURE la, i~ lieu of , reactor 28 (see FIGUR~ l) of the prior art, the system of the 7 ~' invention includes a solids heater 4C, a primary reactor 6C, a secondary reactor 8C and downstream e~uipment. The downstream ec,ui~ment is ~prised essentially of an indirect heat exchanger l~C,a 1~

1~

2'7 s ~
696-147 I ~Z1~7Z~
"

1 I fractionation tower 12C, and a recycle line 14C from the 2 fractionation tower 12C to the entry of the primary reactor 3 6C.

4 l The system also includes a first hydrocarbon feed ' line 16C, a second hydrocarbon feed-quench line 18C, a transfer 6 line 20C ~nd an air delivery line 22C.
7 The first hydrocarbon feed stream is introduced 8 into the primary reactor 6C and contacted with heated solids from the solids heater 4C. The first or primary reactor 6C
in which the first feed is cracked is at high severity conditions.
11 Tne hydrocarbon feed, from line 16C, may be any hydrocarbon gas 12 vr hydrocarbon liquid in the vaporized state which has been used 13 j heretofore as a feed to the conventional thermal cracking process.
14 Thus, the feed introduced into the primary reactor 6C may be ' 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 18 boiling point gas oils such as naphtha having a boiling range 19 between 350 to 650F, high boiling point gas oils having a I boiling range between 650 to 950F and compatible combinations 21 1 of same~ These constituents may be introduced as fresh feed 22 or as recycle streams through the line 14C from downstream 23 ;' purification facilities e.g., fractionation tower 12C. ~ilution 24 ~I steam may also be delivered with the hydrocarbon through lines ~5 ll 16C ana 14C. The use of dilution steam reduces the partial 26 ~' pressure, improves cracking selectivity and also lessens the 27 `` tendency of high boiling aromatic components to form coke.

28 , The preferred primary feedstock for the high 29 l severity reaction is a ligh~ hydrocarbon material selected fron-i 1, the group consisting of low molecular weight, hydrocarbon gasei, s~w 696-147 l i2~`7Z6 ight hydrocarbon liquids, light gas oils boiling between 350 2 and 650F, and combinations of same. These feedstocks offer tne 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. ~he inlet pressure in the line 16C
8 is 10 to 100 psig. The feed should be a gas or gasified liquld.
g The feed increases rapidly in temperature reaching thermal equi-librium with the solids in about 5 milliseconds. As mixing Or 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 14 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. Typi~
17 cally, the ~SF (Kinetic Severity Function) is about 3.5 (97~
18 conversion of n-pentane). Reaction products of this reaction are 19 olefins, primarily ethylene with lesser amounts of propylene and butadiene, hydrogen, methane, C4 hydrocarbons, distillates such 21 as gasoline and gas oils, heavy fuel oils, coke and an acid gas~
~2 ~ Other products may be present in lesser quentities. Feed con--23 ! version in this first reaction zone is about between 95 to 100~ by24 ~ weight of feed, and the yield of ethylene for liquid feedstocks is about 25 ~o 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
2~ ' and combines with the cracked gas from the primary reactor 6C
29 between the primary reactor 6C and the secondary reactor 8C. '~he i combined stream comprising the second unreacted feed, and the '~ '' S&W ll ~
696-1-. lZ~7Z6 i . , , .

l first reacted feed passes through the secondary reactor 8C under 2 ~ low severity reaction conditions. The second feed introduced 3 ll through the line 18C is preferably virgin feed stock but may 4 li also be comprised of the hydrocar~ons previously mentioned, including recycle streams containing low molecular weight hydrocarbon gasesl light hydrocarbon liquids, low boiling 7 point, light compatible ~as oils, high boiling point gas oils, 8 and combinations of same.
9 ! Supplemental dilution steam may be added with the secondary hydrocarbon stream entering through stream 18C. However ll in most instances the amount of steam initially delivered to the i2 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 l' specific to a particular recycle constituent.
16 ~ The hydrocarbon feed delivered throu~h the line 18C
17 is preferably virgin gas oll 400-650F. The second feed is pre-18 !I heated to between 600 to 1200F. and upon entry into the secondary 19 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 quen!h 22 1 60 pounds of effluent from the primary reactor 6C. At this te~
23 ll perature level, the cracking reactions of the first feed are 24 li essentially terminated. However, coincident with the quenching ' of the effluent from the primary reactor, the secondary feed 26 ~ entering through line 18C is thermally cracked at this temperature 27 1 (1500 to 1200F) and pressures of 10 to 100 psi~ at low severity 28 l by providing a residence time in the secondary reactor between 29 l 150 and 2000 milliseconds, preferably between 250 to 500 milli-seconds. Typically, the KSF cracking severity in the secondary

5&~

6 9 fi - 1 4 / l I
~2~`7~6 .
1 reactor is about 0O5 at 300 to 400 milliseconds.
2 , 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 ~eaction 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. Feed conversion in this g second reaction zone is about 30 to 80% by weight of feed, and the yield of ethylene is about 8 to 20% by weight of feed, ll ! 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 I reaction zone, the low severity conditions in the second reaction 16 zone are insufficient to appreciably alter the product dis-17 tribution of the primary products from the high severity reaction l~ zone. Some chemical changes will occur, however these reaction l3 products are substantially stabilized by the direct quench provided by the second feed.
21 The virgin gas oils normally contain aromatic 22 molecules with paraffinic hydrocarbon side chains. For some 23 gas oils the number 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 '26 have a low "aromaticity".
27 ; In the secondary reactor~ these molecules will 28 'I undergo dealkylation - splitting of the paraffin molecules, 2 9 1l leaving a reactive residual methyl aromatic, which will tend to react to form high boilers. The paraffins in the boiling .

s~w 696-1~7 1 ~Zl~7Z6 1 '~ rarlge 400 to 650F are separated from the higher boiling 2 aromatics in column 12 and constitute the preferred recycle 3 to the primary reactor.
4 Other recycle feed stocks can include propylene, butadiene, butenes and the C5 ~ 400F pyrolysis gasoline.
6 1 The total effluent leaves the secondary reactor 7 and is passed through the indirect ~uench 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
The purification facilities 12C employ conventional 11 separation methods used currently in thermal cracking processes~
12 FIGURE 2 illustrates schematically the products obtained. ~ydro-13 gen and methane are ta~en 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 hi.gh severity primary 18 reactor. The product taken through line 28C consists of liquid 19 ' hydrocarbons boiling between C5 and 400F, and is preferably exported although such material may be recycled to the primary 21 i reactor 6C if desired. The light gas oil boiling between 400 22 , to 650F is the preferred recycle feed, but may be removed through 28 ¦ line 30C. The heavy gas oil which boils between 650-950F is 24 l exported through strea~l 32C, while excess residuim, boiling ~ above 950F is removed from the battery limits via stream 34C.
26 The hea~y gas oil and residuim may also be used as fuel within 27 ; the ciystem 28 ~ In the preferred embodiment of the process, the 2 9 ! second feed would he one which is not recommended for high severity operation. Such a feed would ~e a gas oil boiling i, ' .

696-147 1 ~Z~7~

1 ,l above 400F which contains a significant amount of high molecular 2 , weight 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. ~y pro-6 cessing the aromatic gas oil feed at low severity, it is possible 7 to dealkylate the rings, but also to prevent subsequent poly-8 ; merization and coke formation. 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 ~' thereto. The aromatic components of the effluent are removed 16 I from the purification facilities 12C as part of the heavy gas 17 1 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 ! 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.

27 i , ~27-s~ LZ~`726 1 (d) Improved Residence Time Solid-Gas Separation 2 Device and System.

- Referring to FIGURE 15 in the sub~ect invention, in lieu ; separatlon zone or curved segment region 36 and the ~uench area 6 44 of the prior art TRC system (see FIG~RE 1), solids and gas enter 7 the tubular reactor 13D through lines llD and 12D respectively.
S The reactor effluent flows directly to separator 14D where a 9 separation into a gas phase and a solids phase stream is erfected.
The gas phase is removed via line 15D, while the solid phase is 11 sent to the stripping 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 gas 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 19D. Solids 21 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 admitted via line 23D. Vapors are removed from 24 the stripping vessel through line 24D and, if economical or if need be, sent to down-stream purification units. Stripped solids ~9 ~

.

S&W l!
6~6-14~ 12~726 1 ll removed from the vessel 22D through line 25D are sent to re-2 ,I generation vessel 27D using pneumatic transport gas from line 3 ~, 26D. Off gases are removed from the regenerator through line 28D.
4 I 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 degradatlon 8 I and ensure optimal yield and selectivity of the desired products.
9 ' Further, the separator 14D operates in a manner that eliminates or at least significantly reduces the amount of gas entering the 11 ; stripping vessel 22D inasmuch as this portion of the gas product i 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 veloci~y 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 described below.
FIG~R~ 16 is a cross sectional elevational view 21 ! showing the preferred embodiment of solids-gas separation device 22 1 14D of the present invention. The separator 14D is provided 23 ; with a separator shell 37D and is comprised of a solids-gas 24 disengaging chamber 31D having an inlet 32D for the mixed phase i stream, a gas phase outlet 33D, and a solids phase outlet 34D~
26 ~ The inlet 32D and the solids outlet 34D are preferably locatec.
27 at opposite ends of the chamber 31D. While the gas outlet 33D
~8 i lies at a point therebetween. Clean-out and maintenance manways 29 35D and 36D may be provided at either end of the chamber 31D.
I The separator shell 37D and manways 35D and 36D preferably are .1 696-la7 I, 121~7~6 !! .

1 l lined w~h erosion resistent linings 3~D, 39D and 41D re-2 , spec~ively which may be re~uired if solids at high velocities 3 are encountered. Typical commercially available materials 4 for erosion resistent lining include Carborundum Precast Carbofrax D, Carborundum Precast ~lfrax 201 or their equivalent.
6 A thermal insulation lining 40D may be placed between shell 37D
7 and lining 38D and between the manways and their respective 8 ! erosion resistent linings when the separator is to be used ~ in high temperature service. Thus, process temperatures above 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 constructlon 14 the separator 14D shell is preferably fabricated from cylindrical sections such as pipe 50D, although other materials may, of 16 course, be used. It is essential that longitudinal side walls 17 l 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 FIGU~E 17.
21 The em~odi~ent shown in FIGV~ 17 defines the geometry of the 22 flow path by adjustment of the lining width for walls 51D and 23 . 52D. Alternatively, baffles, inserts, weirs or other means 24 ,¦ may be used. In like fashion the configuration of walls 53D
~ and 5~D transverse to the flow path may be similarly shaped, 26 although this is not essential. FIGURE 18is a cutaway view 27 I along Section 4-4 of FIGURE 16 wherein the separation shell 37D
28 l~ is fabricated from a rectangular conduit. Because the shell 37D
29 has rectilinear walls 51D and 52D it is not necessary to adjust I the width of the flow path with a thickness o~ lining. Linings 1~`~

~Z1~726 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 required 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 subsequently 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 size. ~ecause 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/secD, 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 80%, although the preferred embodiment, discussed below, can obtain over 90% removal of solids.

,; ,.
.. .

12~(~7Z6 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 Figures 16 and 17 we have found that for a given height H of chamber 31D, efficiency increases as the 180 ~-bend between inlet 32D and outlet 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 10w path and, hence, residence time decreases. Assuming an inside diameter 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~ 5 Di. Below 1.5 Di 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 33D would be too close to one another to allow welded fabrication.
It has been found that the height of flow path H
should be at least equal to the value of Di or 4 inches in height, whichever is greater. Practice teaches that if H is less than Di or 4 inches ~he incoming stream is apt to disturb the bed solids 42D, thereby re-entraining solids in the gas product leaving through outlet 33D. Preferably H is on the order of twice Di to obtain even greater separation efficiency. While not otherwise limited, it is apparent that too large an H eventually merely increases residence time without substantive increases in efficiency. The width W of the flow path is X

~Z~LI','7Z6 S&W
696-147 l l preferably between 0.75 and l. 25 times Di, most preferably between 2 ~ 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 of outlet 34D should be sized so that a pressure differential 7 1 between the stripping vessel 22D shown in FIGUREls 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 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 ~5 ehight of solids in line 16D~ As the differential incrPases 16 , ~he net flow of gas to the stripping vessel 22D decreases.
17 Solids, having ~ravitational momentum, overcome the differentlal, 18 while gas preferentially leaves through the gas outlet 33D~
l9 ; By reaulatir.g the pressure in the stripping vessel 20 22D it is possible to control the amount of gas 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 ~I cont.rol 29D device on vessel 22Do Alternatively, as suggested 24 1 above, the pressure may be regulated by selecting the size of l' the outlet 34D and conduit 16D to obtain hydraulic forces 26 ~¦ acting on the system that set the flow of gas to the stripper 27 Il 32D~ ~hile such gas is degraded, we have found that an increa~e 28 in separation efficiency occurs with a bleed of sas to the 29 ' stripper of less than 10%, preferably be~ween 2 and 7%. Economic ~ and process considerations would dictate whether this mode of i i ' -33 . .

-696-1~7 ~ ~21~726 !
, 1 operation should be used. It is also possible to design the system to obtain a net backflow of gas from the stripping 3 ~I vessel. This gas flow should be less than 10~ of the total 4 I feed gas rate.
By establishing a minimal flow path, consistent 6 , with the above recommendations, residences times as low as 7 0.1 seconds or less may ~e obtained, even in separators 8 , having inlets over 3 feet in diameter. Scale-up to 6 feet 9 ,i in dlameter is possible in many systems where residence times approaching 0.5 seconds are allowable.
11 ~ In the preferred embodiment of FIG~RE 16, a weir 44D
12 is placed across the Elow path at a point at or just beyond the 13 gas o~tlet to establish a positive height of solids prior to 14 l, solids outlet 34D. By installing a weir (or an equivalent ~ 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. An arc of this shape 19 ~ diverts ~as towards the gas outlet and creates the U-shaped I gas glow pattern illustrated diagrammatically ~y line 45D in 21 ' FIGURE 16. Without the weir 44D an arc sor,lewhat less than or 22 equal to 90 would he 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 nei~he_ 25 ' efficiency nor gas loss (to the stripping vessel~ is affected 26 ll ad~ersely, the flow pattern of line 61D increases residence time~
27 ~ and more importantly, creates greater potential for erosion at 28 ! areas 62D, 63D and 64D.
29 The separator of FIG~RE 20 is a schematic diagram of 30 ' another embodiment of the separator 14D, said separator 14D

' -34-:

6S9&6W1~7 1 I having an extended separation chamber in the longitudinal 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 length. L is preferably less than or equal 5 1l to 5 Di. Although the gas flow pattern 61D does not develope 6 I the preferred U-shape, a crescent, shaped arc is obtained 7 which limits erosion potential to area 64Do E~bodiments 8 shown by FIGURES 1~ and 20are useful when the solids loading 9 1 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 FIGURE21 it is equally possible to use a 13 1 stepped solids outlet 65D having a section 66D collinear with lq the flow path as well as a gravity flow sectlon 67D. Wall 68D
replaces weir 44D~ and arc 43D and flow pattern 45 are similar 16 to the preferred embodiment of FIGURE 16. Because solids accu~u-17 l 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 sufficient , force to expel the solids through outlet 65D. However, because 21 of the restricted solids flow path, better deaeration is obta~ne~
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 steppec.
Ij 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 usedO
28 The separator of FIGURE21 or2~ may be used in 29 I conjunction with a venturi, an orifice~ or an equivalent flow restriction device as shown in FIGURE 23. The venturi 69D having ;! -35-~;
l, ~
S&~ 1 696-147 ll ~ ~07Z6 1 ~i dimensions Dv (diameter at venturi inlet), DVt tdiameter of 2 I venturi throat), and 0 (angle of cone formed by projection 3 1 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 inFIG~RE 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 b~ the force g of the incoming stream. As before the gas portion of the feed 1 follows the U-shaped pattern of line 45D. 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 this reason 14 i the solids outlet of FIGURE 23 is preferred. ~lost preferably, the venturi 69D is placed in collinear section 66D as shown lr 16 FIGURE 24 to improve the deaeration of the solids. Of course, 17 each of these alternate embodiment may have one or more of the 1~ ~ optional design features of the basic separator discussed in 19 relation to FIGURES 16, 17 and 1~.
The separator of the present invention is more 21 clearly illustrated and explained by the examples which follow.
22 ; In these examples, which are basedon data obtained during 23 experimental testing 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 drawing 26 ~igures and listed in the Nomenclature below:
27 I CL Distance between inlet and gas outlet centerlines 28 I D Insiae diameter of inlet 29 ~ Dog Inside diameter of gas outlet 30 ! 05 Inside diameter of solids outlet Dv Diameter of venturi inlet DVt Diameter of venturi t.hroat -S&W ~ q~7 26 . ' , H Height of flow path 2 ~ Height of weir or step 3 L Length from gas outlet to weir or step as 4 indicated in Figure 6 W Width of flow path 6 3 Angle of cone formed by projection 7 of convergent venturi walls, degrees 8 Table_I
9 : _m nsions of Separators in Examples 1 to 10, inches*
Example 11 Dimension _1 2 3 4 _ S 6 7 8 9~ 10 12 CL 3.875 3.875 3.8753.875 3.8753.87S 11 11 3.5 3.5 13 Di 2 2 2 2 2 2 6 6 2 2 14 Dog 1 75 1.75 1.75 1.75 1.75 1.75 4 4 Dos 2 2 2 2 2 2 6 6 2 2 16 Dv ~ ~ ~ 2 17 Dvt 18 i H 4 4 4 4 4 4 12 12 7.5 c 19 ' Hw 0-75 0.75 0.75 0.75 0.75 0.75 2.25 2.25 0 4 L 0 2 2 0 0 0 0 0 :L0 0 21 W 2 2 2 2 2 2 6 6 ~ ~
22 ~, degrees - - - - ~ 28 23 ~ Except as noted 24 ;, Example 1 25 ! In this example a separator of the preferred 26 embodiment of FIGURE 16was tested on a feed mixture of air and 27 silica alumina. The dimensions of the apparatus are specifier 28 in Table I. No~e tilat the distance L from the gas outlet to 29 , the weir was zero.

~ -37-lZlC~'726 The inlet stream was comprlsed of 85 Et.3/min. of air and 52 lbs./min. of silica alumina having a bulk density of 70 lbs./ft3 and an average particle size of 100 microns. The stream density was 0.612 lbs./ft.3 and the operation was performed at ambient temperature and atmospheric pressure.
The velocity of the incoming stream through the 2 inch inlet was 65.5 ft./sec., while the outlet gas velocity was 85.6 ft./sec. through a 1.75 inch diameter outlet. A positive seal of solids in the solids outlet prevented gas from being entrained in the solids leaving the separator. Bed solids were stabilized by placing a 0.75 inch weir across the flow path.
The observed separation efficiency was 89.1%, and was accomplished in a gas phase residence time of approximately 0.008 seconds. Efficiency is defined as the percent removal of solids from the inlet stream.
Example 2 The gas--solids mixture of Example 1 was processed in a separator having a confiyuration illustrated by Figure 20.
In the example the L dimension is 2 inches; all other dimensions are the same as Example 1. By extending the separation chamber along its longitudinal dimension, the flow pattern of the gas began to deviate from the U-shaped discussed above. As a result residence time was longer and turbulence was increased. Separation efficiency for this example was 70~8~.
Example 3 The separator of Example 2 was tested with an inlet stream comprised of 85 ft.3/min. of air and 102 lbs./min. of silica alumina which gave a stream density of 1.18 lbs./ft3, or approximately twice that of Example 2. Separation efficiency improved to 83.3%.

~2~7Z6 6~6-147 ,, 1 ' Fxample 4 2 , The preferred separator of Example 1 was tested 3 ~ at the inlet flow rate of Example 3. Efficiency increased 4 , slightly to 91.3%.
Example 5 6 , The separator of FIGURE 16was tested at the con-7 ditions of Example 1. Although the separation dimensions are 3 specified in Table I note that the distance CL between inlet 9 and gas outlet centerlines was 5.875 inches, or about three times the diameter of the inlet. This dimension is outside 11 the most preferred range for CL which is between 1.50 and 2.50 12 Di. Residence time increased to 0.01 seconds, while efficiency 13 was 73.0%.
14 ,, Example 6 Same conditions apply as for Example 5 except 16 that the solids loading was increased to 102 lbs./min. to ~ive 17 a stream density of 1.18 lbs./ft.3. As observed previously 18 ~ in Examples 3 and 4, the separator efficiency increased with 19 h~gher solids loading to 90.6%.
,:
Example 7 21 ~ The preferred separator configuration of FIGURE 16 22 1 was tested in this Example. However, in this example the appaxa-23 tus was increased in size over the previous examples by a 24 factor of nine based on flow area. A 6 inch inlet and 4 inch 25; outlet were used to process 472 ft.3/ min. of air and 661 lbs./min 26of silica alumina at 180~F. and 12 psig. Tile respective velocltie 27 I were 40 and 90 ft.~sec. The solids had a bulk density of 70 28 ~ lbs./ft3 and the stream density wasl.37 lbs./ ft.3 Distance 29 ,~ CL between inlet and gas outlet centerlines was 11 inches, or 1.83 times the inlet diameter; distance L was zero. The bed was : -as~

S&W ~' ~Z~726 696-1~7 ,~

1 stabilized 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 ~ was such that a positive flow of displaced gas from the I 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 ti~e approximately 8 0.02 seconds.
9 Example 8

10 , The separator used in Example 7 was tested with

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

12 collection vessel was vented tc the atmosphere and the pressure

13 differential adjusted such that 9g of the feed gas, or 42.5 ft.3/

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 FIGURE 22was tested in a unit 19 having a 2 inch inlet and a 1 inch gas outlet. The soiids out-let was 2 inches in diameter and was located 10 inches away 21 from the gas outlet (dimension L). A weir was not used~ The 22 feed was comprised of 85 ft.3~min. of air and 105 lbs./min. of 23 i spent fluid catalytic cracker catalyst having a bulk density 24 ; of 45 lbs./ft.3 and an average particle size of 50 microns. This gave 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 27 j in Example 7 there was a positive counter-current flow of 28 I displaced gas from the collection vessel to the separator.
29 I This flow was approximately 1.7 ft.3/ min. at a velocity of 1.3 ft./sec. Operation was at ambient temperature and atmos-31 pheric pressure. Separator efficiency was 95.0%.

~ ~iw 12~ 7Z~
696-147 l' 1 l Example 10 2 The separator of FIGURE23 was tested on a feed 3 comprised of 85 ft.3/ min. of air and 78 lbs./min. of spent 4 'I 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 outlet 7 i velocity of 262 ft ./sec. This separator had a stepped 8 solids outlet with a venturi in the collinear section of the 9 outlet. The venturi mouth was 2 inches in diameter, while the 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 efficiency of 92.6~ was measured, and the solids leaving the 13 separator were completely deaerated except for interstitial gdS
14 remaining in the solids' voids.

I \

23 ~ \

29 1 \

--Dl--96-147 lZ1~726 1 (e) Imoroved Solids Quench Boiler and 2 Process.

As see~n in FIGURE 25,in lieu of quench zone 44, 46 (see 6 FIGURE 1) of the prior art, the composite solids quench boiler ?E

7 i of the subject invention ls 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 lC 36E for the delivery of gas to the fluid bed-quench riser.

11 The quench exchanger 4E as best seen in FIGURES 26 an~ 27, 12 i5 formed with a plurality of 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 -~2-~`~

S&~1 6q~-l47 ~ 726 ., .

quench exchanser 4E. The tubes 16E are joined together, pre-2 ~ ferably by welding, and form a pressure-tight memL>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 allow for the passage of effluent gas and particulate solids there-6 around. The arrays of tubes 16E~1~3E 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 4E is 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 by piping such as an elbow 30E which is arranged to deliver the 16 cooled effluent and particulate solids to the cyclone separator 17 ; 8E.
18 The fluid bed quench riser 6E is essentially a sedled i9 vessel attached in sealed relationship to the quench exchange~
4E. The fluid bed-quench riser 6Ei'; arranged to receive the 21 reactor outlet tube 36E which is preferably centrally disposed at 22 the bottom of the fluid quench riser 6E. ~ slightly enlarged 23 , centrally disposed tube 38Eis aligned with the reactor outlet 24 36E and extends from the fluid bed-quench riser 6E into the 25 ; quench exchanger 4E~ In the quench exchanger 4E r 'che central`y 26 disposed fluid bed-(iuench riser tube 38E terminates in a conlc:al 27 - opening 40E. The conical opening 40E is provided to facilitate ~8 nonturbulent transition from the quench riser tube 38E to the 29 enlarged opening of the quench exchanger 4E.It has been ~ound ; that the angle of the cone ~, best seen in FIGURE2~, should 31 be not greater han 10 degrees~
, -~3~

696-14711 ~Z~726 ¦I The fluid bed 42E contained in the fluid bed quench 2 1, riser 4E: is maintained at a level well above the bottom of the 3 ! quench riser tube 3~E. A bleed line 50E is provided to bleed 4 ,, solids from the bed 42E. Although virtually any particulate ,~ solids can be used to provide the quench bed 42E, it has been 6 ~l 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 al~mina solids.
9 ,' As best seen in FIGURE 2~,the opening 48E through ~I which the fluidized particles from the bed 42E are drawn into 11the quench riser l:ube 38E is defined by the interior of a cone 12 ' 44E at the lower end of the quench riser tube 38E and a refractory 13 l~ cone 46E located on the outer surface of the reactor outlet 1~ ' tube 36E. In practice, it has been found that the refractory I cone 46E can be formed of any refractory material. The opening 16' 48, defined by the conical end 44E of the quench riser tube 38E
17and the refractory cone 46E, is preferably 3-4 square feet for 18 ~ a unit of 50 ;~MBTU/HR capacity. The opening is sized to insure 19 I penetration of tne cracked gas solid mass velocity of 100 to 20 ~ 800 pounds per second per square foot is required. The amount 21~~ of solid5 from bed 42E: delivered to the tube 38E is a function 22 l' of the velocity of tl-e gas and solids entering the tube 38E
23 jl from the reactor outlet 36E: and the size o the opening 48E. ., 24 1 In practice, it has been found that the Thermal I Regenerative Cracking (TRC~ reactor effluent will contain 26 ~' approximately 2 pounds of solids per pound of gas at a tem-27ll perature of about 1,400F to 1,600F.
2~I The process of the solids quench boiler 2E of 29 FIGURES 25-2~ is illustrated by the following example. Effluent 30, from a TRC outlet 36E at about 1,500F is delivered to the quench -~4-, . .

S&W
6~6-147 'l ~z~7z6 i, 1 ¦I riser tube 3~E at a velocity of approximately 40 to 100 feet 2 , per second. The ratio of particulate solids to cracked effluent 3 1 entering or leaving the tube 36E is approximately two pounds of 4 8 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 effluent stream by the eductor effect i5 between tw~nty five , 7 and fifty pounds solid per pound of gas. In 5 milliseconds the 8 addition of the particulate solids fro~ 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 ln ~he tubes 14 ' 16E, l~E, and 20E. ~ith 120,000 pounds of effluent per hour, 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 i9 ~ cyclone separator 8E wherein the bulk of the solids is re~ove~
; from the quenched-cracked gas and returned through line lOE
21 ,I to the quench riser 6E. , 23 ' '26 , 27 !1 \
i 28 ,~ \
29 ' \
30 l, \

.

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1 (f) Improved Preheat Vaporization System.

3 Again referring ~o FIGURE29,in lieu of pr~at zone 24 (FIGJRE 1) 4 of the system 2F of the subject invention is em~odied in a TRC system and is co~prised of essentially a liquud 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 mixer 14F for flashing. The system also preferentially includes a knockout drum 16F for the preheated vapor.
The liquid feed heater 4F is provided for heating the 2 hydrocarbon feedstock such as desulfurized Kuwait ~3GO to initially elevate the temperature of the feedstock.
The initial mixer 8F is used in the system 2F to 14 initially flash superheated steam from a steam line 6F and the heated feedstock deliv~red from the liquid feed heater 4F by 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 S&W '~
696-147 j, 1 ~Zl(~7~ 1 i ''discharged through a line 22F from the separator overhead and 2 the remalning liquid is discharged through a line 26F.
3 1, 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.
A 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 temperature of about 440 to 700F. The heated heavy 17 , oil hydrocarbon feedstock is then deli-~ered through the line 18 18F to the mixer 8E`. Superileated steam from the line 6F if 19 mixed with the heated heavy oil hydrocarbon feedstock in the ;mixer 8F and the steam heavy oil mixture is flashed to about 21 i,700 to 800F. For lighter feedstock the flashing temperature 22 will be about 500 to 600~F., and for heavier feedstock the 23 ; flashing temperature will be about 700 to 900F.
24 ,1 The flashed mi.Yture of the steam and hydrocarbon is ' sent to the system separator lOF wherein the vapor or gas is 2~ ,taken overhead through the line 22F and the liquid is 27 l~ discharged through the line 26F. Both the overhead vapor and 28 liquid bottoms are in the temperature range OL about 700 to 29 ~B00F. The temperature level and percent of hydrocarbon I;vaporized are determined wi~hin the limits cf equipment fouling , ~~ -47-Shl~
696-1~7 j ~ Zl~'72 ,,i 1 ~criteria. The vapor stream in the line 22F is comprlsed 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 through 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 ~lashed vapor is then sent downstream through the line 28F to the knockout drum 16F for removal of any liquid 16 from the vapor. Finally, the vaporized hydrocarbon feed is 17 sent through the lin~ 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 ~n 21 the system 2F prior to delivery to a reactor. The Nigerian Heavy 22 Gas Oil has the following composition and properties:
23 l 24 Elemental Analysis, Wt.% Properties Carbon 86.69 Flash Point, F. 230.0 Hydrogen 12.69 Viscosity, S~S 210 F 44.2 '26 Sulfur .10 Pour Point, F +90.0 Nitrogen .047 Carbon Residue, Ramsbottom .09 27 I Nickel .10 Aniline Point, C 87.0 Vanadium .10 29 l ;

, I -4~-_-- .

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6~6 1 Disti_lation 2 Vol. %

10669.2 4 30755.6 50820.4 S 70874.4 90944.6 6 EP1,005.8 8 3,108 pounds per hour of the Nigerian Heavy Gas Oil i5 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 11 superheated steam at 1,100F. 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 3il is vaporized.
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 19 vaporO 1,243. 2 pounds per hour of hydrocarbon are discharged through the line 2ÇF as liquid and sent to the mixer 14F .
21 The mixture of 622 pounds per hour of steam and 22 1,864.8 pounds per hour of hydrocarbon are superheated in the 23 Ivapor superheater 12F to 1,139F. and delivered through line 24 24F to the mixer 14F. The mixer 14F is maintained at 45 psia.
I 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 hydrocarbon are flashed in the mixer 28 14F to 990F.
29 The vaporization of the hydrocarbon is effected with a ''steam to hydrocarbon ratio of 0.2. The heat necessary to vaporize 696-147 ~ lZl(~7Z6 ., , 1 ~ the hydrocarbon and generate the necessary steam is 2.9~4 MM
2 'BTU/hr.
3 I The same 3,108 pounds per hour of Nigerian Heavy Gas 4 Oil feedstock vaporized 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 generate the necessary steam is 6.541 MM BTV/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 required, which temperature is effectively beyond design 11 ~1limitations.
12 l~
13 l~ SUM~RY
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 I 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. If the total mix is quenched, to avoid 22 'rapid separation, a high thermal inefficiency will result since 23 lall the heat of the solids will be rejected to some lower 24 1l level heat recovery. Thus, a rapid high efficiency separator, 25 !~ 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 solids to an ethylene reactor, the flow must be controlled 29 1l to within +2 percent or cracking severity oscillations will be ¦'greater than that presently experienced in coil cracking. The '. :

. _ S&~
696-i , 121(~726 1 subject feed device (local fluidization) mini~izes bed height 2 as a variable and dampens the effect of over bed pressure fluct 3 uations, both of which contribute to flow fluctuations. It is ~ 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 reacting time). Both of 8 the features permit the TRC to move to shorter residence times 9 which increase selectivity. Conventional fluid bed feeding devices are adequate for longer time and lower temperature 11 reactions (FCC) especially 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 FIGU~ES 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 600Fo 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 osig, 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.
28 However, instead of an oil quench, a circulating solids stream 24 could be used to quench the effluent. As in the reaction itself~
the coke 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 steam ~1500 psig~) 28 ~hich increased the overall thermal efficiency of the process.
29 The oil loop can not operate at these temperatures due to instabilities (too many light fractions are boiled off, yielding s~ ~
696-l l an oil that is too viscous). The use o~ solids can be done for 2 both T~C or a coil, but it is especially suited to a TRC since 3 it already uses solids. During quenchin~, the coke accumulates a on the solid. It 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 ~1ith reference to the preheat vapori~ation system of 8 the subject invention (see FI~URE 29), it is noted that 9 the T~C has maximum economic advanta~es ~hen crackin~ heavy feedstocks (650F+ boiling ~oint) at low steam dilutions.
ll Selectivity is favored by rapid and intimate mixing. Rapid 12 and intimate mixin~ is best accom21ished with a vapor feed 13 rather than a liquid feed.
la Finally, with reference to the sequer,tial crackin~
system of the invention (see FIGURE 14), it ls 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 addition to any yield advantages. It can l9 be applied to both T~C and a coil. Its synergism with TRC
is that it permits the use of longer 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 lower 23 temperatures compared to the coil case.
24 ~hile there has been described what is considered to be preferred embodiments of the invention, variations and modif-2~ ications therein will occur to those skilled in the art once 2i they become acquainted with the basic concepts of the invention~
28 Therefore, it is intended that the ap~ended claims shall be 29 construed to include not only the disclosed embodiments but all such variations and modifications that fall within the true 31 spirit and scope of the invention.

Claims (14)

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 hydrosulfurized residual oil along with the entrained inert solids and the diluent gas are passed through a cracking zone for a residence time of 0.05 to 2 seconds, the improvement in the process for preheating the heavy oil hydrocarbon 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.

2. A 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.

3. A process as in Claim 2 wherein the liquid heavy oil is preheated to 440° to 700°F.

4. A process as in Claim 2 further comprising the step of removing liquid from the flashed superheated vapor and liquid produced in the initial flashing step.

5. A process as in Claim 2 wherein the initial flashing of the steam and the liquid heavy oil hydrocarbon is at a temperature of 700° to 800°F.

6. In a TRC system wherein the temperature in the cracking zone is between 1300° and 2500°F and wherein the hydrosulfurized residual oil along with the entrained inert solids and the diluent gas are passed through a cracking zone for a residence time of 0.05 to 2 seconds, the 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. vapor feed superheater for heating the vapors from the first mixer to about 1,030°F.; and d. a second mixer for flashing the superheated vapor and the liquid from the first mixer.

7. A system as in Claim 6 further comprising means for removing liquid from the flashed superheated vapor and liquid from the first mixer.

8. A process for pre-heating heavy oil hydrocarbon feedstock comprising:

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.

9. A process as in Claim 8 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.

10. A process as in Claim 9 wherein the liquid heavy oil is preheated to 440° to 700°F.

11. A process as in Claim 8 further comprising the step of removing liquid from the flashed superheated vapor and liquid produced in the initial flashing step.

12. A process as in Claim 7 wherein the initial flashing of the steam and the liquid heavy oil hydrocarbon is at a temperature of 700° to 800°F.

13. A system for pre-heating 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 superheated vapor and the liquid from the first mixer.

14. A system as in Claim 13 further comprising means for removing liquid from the flashed superheated vapor and liquid from the first mixer.