CA1180297A - Thermal regenerative cracking (trc) apparatus and process - 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

The present invention relates to improvements in Thermal Regenera-tive Cracking (TRC) apparatus and process, as described in ~.S. Letters Patent Nos. 4,061,562 and 4,097,363 to McKinney et al.
Accordingly, the present invention relates to a TRC
process wherein the temperature in the reaction chamber is between 1300 and 2500F and wherein the hydrocarbon fluid feed or the hydrosulfurized residual oil along with the en-trained inert solids and the diluent gas are passed through the reaction chamber for a residence time of 0.05 to 2 seconds, the improvement wherein the step of cracking hydrocarbon feed to produce olefins comprises: delivering hydrocarbon feed to a first zone; thermally cracking the hydrocarbons in the firs-t zone at temperatures above 1,500DF.; discharging the cracked effluent from the first zone to a second zone; delivering a second hydrocarbon feed to the entry of the second zone; and mixing the cracked effluent from the first zone and the second hydrocarbon 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.
The invention will be more fully appreciated from reference to the accompanying drawings in which FIGURE 1 is a schematic diagram of a TRC system and process according to the prior art.
FIGURE 2 iq a schematic diagram of -the fuel gas generation system and process of the subject invention.
FIG~RE 3 is an alternative embodiment ~herein the fuel gas is burned to flue gas to provide additional heat for the particulate solids.
FIG~RE 4 is a cross-sectional elevational view of the solids feeding device and system as applied to tubular reac-tors and for use with gaseous feeds.
~' ~~ -2-FIGURE 5 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 ~he 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 ~ is a view through line 8-8 of FIGURE 5.
FIGURE 9 is an isometric view of thé plug which extenas into the mixing zone to reduce flow areaO
FIGURE 10 is an alternate 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 ]2 is an alternate embodiment of the preferred invention wherein a second feed gas is contemplated.

,... .

~ ~ ~`()29 7 (S&W) 1 FIGURE 13 is a view of the apparatus of FIGURE 12

2 thLough 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.
FIGURE 15 is a schematic flow diagram of the 6 separation system of the present invention as appended to a 7 typical tubular reactor.
8 FIGURE 16 is a cross sectional elevational view of g the prererred embodiment of the separator.
FIGURE 17 is a cutaway view through 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 chamber.
21 FIGURE 21 is a sketch of an alternate en~odiment of 22 the separati~n device wherein a stepped solids outlet is employed, 23 said outlet having a section collinear with the flow path as well 2~ as a gravity flow section.
FIGURE 22 is a variation of the embodiment of FIGURE
26 21 in which the solids outlet of FIGUPE 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 sec~ion of the solids outlet.

1 FIGURE 24 is a variation of the e~bodiment of 2 FIGURE ~3 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;
FIGURE 26 is a detailed cross sectional elevational 6 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 FI~URE 28 iS a detailed drawing of the reactor ou,tlet and fluid bed quench riser particle entry area.
11 FIGURE 29 is a schematic diagram of the system of 12 the inven-tion for vaporizing heavy oil.

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

22 i~ Referring to FIGURE 1, in the prior art TRC process 23 l~and system, thermal cracker feed oil or residual oil, with or 24 I~ithout blended distillate heavy gas, entering through line 10 ~an~ hydrogen entering through line 12 pass through hydrodesulfurized 26 ~Izone 14. Hydrosulfurization effluent passes through line 16 and 27 'enters flash chamber 19 from which hydrogen and contaminating gases, 28 l~ncluding hydrogen sulfide and ammonia are removed overhea~ through 29 ~line 20, while flash liquid is removed through line 22. The flash ¦!

S&W ;`
696 147 ~ .'2~7 1, '' i 1 ~ liquid passes th.rough 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 30O
4 1l A stream of hot regenerated solids is charged ' through line 32 and admixed with steam or other fluidizlng gas 6 l entering through line 34 prior to entering the bottom of riser 7 ,i 28. The oil, steam and hot solids pass in entrained flow up-war~ly through riser 28 and are discharged through a curved g ,, segment 36 at the top of the riser to induce centrifugal separ-ation of solids from the effluent stream. A stream containing 11 i~ most of the solids passes through riser discharge segment 38 and 12 l~ c~n be mixed, if desired, with make-up solids entering through 13 line 40 before or after entering solids separator-stripper 42.
14 , Another stream contai.ning most of the cracked product is dis-1! 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 jl Stripper steam is charged to solids separators 42 19 'i and 48 through lines 50 and 52, respectively. Product streams 11 are removed from solids separators 42 and 48 through lines 54 21 ll and 56, respectively, and then combined in line 58 for passage 22 '¦ to a secondary quench and product recovery ~rain, not shown.
23 !I Coke-laden solids are removed from solids separators 42 and 48 24 ~11 through lines 60 and 62, respectively~ and combined in line 64 ¦I for passage to coke burner 66. If required, torch oil can be 26 ~¦ added to burner 66 through line 68 while stripping stea~, may be 27 ~¦ added through line 70 to strip combustion gases from the heated 28 ll solids. Air is ~harged to the burner through line 69. Combustion 29 '¦ gases are removed from the burner through line 72 for passage I to heat and energy recovery systems, not shown, while regenerated ,l 6 ,, !;
S&W ~ 7 6!36-147 .' , I .
1 ~' hot solids which are relatively free of coke are removed from 2 the burner tnrough line 32 for recycle to riser 28. In order 3 I to produce a cracked product containing ethylene and molecular 4 1~ hydrogen, petroleum residual oil is passed through the catalytic ,, hydrodesulfurized zone in thepresence of hydrogen at a tem-6 perature between 650F and 9~0F, with the hydrogen ~eing 7 chemically combined with thRoil during the hydrocycling step.
8 The hydrosulfurization residual oil passes through the thermal g ' crac}-ing zone together with the entrained inert hot solids ,l functioning as the heat source and a diluent gas at a temperature 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 ,' ethylene and hydrogen. For the production of ethylene by 14 I thermally cracking a hydrogen feed at least 90 volume percent '' of which comprises lisht 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 18 ,, through the cracking zone at a temperature between 1300F and 19 ' 2500~F for a residence time of 0.05 to 2 seconds. The weight , ratio of oil gas to fuel oil is at least 0.3, while the cracking 21 I 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 ¦~ immediately upon leaving the cracked zone to a temperature 24 I below 1300F ensures that the ethylene yield is greater than ! the methane yield on a weight basis.

S&W
696-147 ~ 7 1 (a) Improved Fuel Gas Generation For Solids 2 Heating.

4 I FIGURE 2 illustrates the improved process and system of ~the invention as may be e~bodied in a prior art TRC system, in lieu 6 lof the coke burner 66 (FIGURE 1~. Palticulate solids and hy~arbon feed gas 7 i'enter a tubular reactor 13A through lines llA and 12A xespectively. The cracked 8 I effluent from the tubular reactor 13A is separated from the particulate solids 9 in a separator 14A and quenched in line 15A by quench material injected from line 17A. The solids separated from the effluent are dPlivered through line 16A to a soli~s 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 1 free product gas is taken overhead from the secondary separator ~,18A through line l9A.
16 _ _ 696-147 ~ 7 ., , .

1 l' The particulate solids in the solid stripper 22~, 2 'having delivered heat during the thermal cracking in the tubular 3 reactor 13A, must be reheated and returned to the tubular reactor 4 l 13A to continue the cracking process.
I The particulate solids prior to being reheated, are 6 ,stripped of gas in the solid strippex 27A by steam delivered to 7 I~the solid stripper 22A through line 23A.
8 ! After the particulate solids have been stripped of g yas impurities in the solid stripper 22A, the particulates solids;
, are at a temperature of about 1,450F.
11 ' The fuel gas generation apparatus of the invention 1~ consists of a combustion vessel 30Ar and pre-heat equipment for 13 l fuel, air (or 0~) and steam which are delivered to the oombustion 14 ! vessel 30A. Pre-heaters 32A~ 34A, and 36 are shown in fuel line li 38A, air line 40A, and steam line 42A respectively.
16 ,~ l~he system also includes a transfer line 44A into 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 ,l line 44A is sized to afford sufficient residence time for the 21 I,' steam emanating from the combustion vessel 30A to decompose by 22 ij the reaction with carbon in the presence of hydrogen and to remove 23 l, the net carbon from the solids-gas mixture. In the preferred 24 1~ embodiment the transfer line 44 will be about 100 feet long. A
2~ ¦l line 26A is provided for pneumatic transport gas if necessary.
26 il A separator, such as a cyclone separator 46A is 27 1 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 v~ssel 27A
', and the fuel gas is taken overhead through line ~OA.

,i _9_ S&W
696-147 ~ 7 , I .

1 In the process, fuel, air and steam are delivered 2 ~ through lines 38A, 40A and 42A respectively to the combustion 3 I vessel 30A and combusted therein to a temperature of about 4 2,300F. to produce a fuel gas having a hi~h ratio of CO to C02 and at least an equivalent molal ratio o 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 9 i carbon in the fuel gas-particulate solids will mix will be ; removed before reaching the separator 46A.
11 ' The fuel gas rom combustion vessel 30A at a 12 temperature of about 2,300F. is mixed in the tubular vessel 44A
13 l with stripped particula~e 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 tubular 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 !~ As a result, the particulate solid-fuel gas mixture is cooled by about 30F. i.e., from 1,780F. to 1,750F.
21 ' The particulate solid-fuel gas mixture is separated 22 in the separator 46A and the fuel gas is taken at 1,750F. through 23 ,~ line 50A. The particulate solids are delivered to the hot solids 24 il, ho~ vessel 27A at 1,750F. and then to the tubular reactor 13A.
'! In the alternative embodiment of the invention 26 !, 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 ;l .

s~w~ 'r~ 37 696-147 j' ii 1.
I.i, 1 l'the line 44A is burned to elevate the temperature of the particu-2 ,;late solids to about 1,750F. The resultant flue gas is separated 3 li from the hot solids in the separator 46A and discharged through 4 ¦,the line 52A. The hot particulate solids are returned to the ,system to provide reaction heat.
6 1l An example of ~he 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 MM SCPD
' of fuel ~as.
11 i The 15 MM SCFD of fuel gas at 2,300F. is ~ixed in 12 ,the transfer line 44A with 1 MM 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 com-l'posite fuel gas-particulate solids gas mixture reaches an ,lequilibrium temperature of 1,480F, at 5 psig in about 5 milli-17 ,Iseconds. Thereafter, 13 ~ SCFD of air is delivered to the 18 lltransfer line 44A and the 15.6 l~M SCFD of fuel gas is burned 19 ,Iwith the air to elevate the solids temperature to 1,750F. and iI burn the ',600 pounds per hour of carbon from the particulate 21 ,solids.
22 ¦¦ The ~ombusted ~as from the transfer line 44A is 23 i! separated from the solids in the separator 46A and discharged 24 ¦las flue ~asO

~1 , i"

S&~ )2~7 1 (b) Improved Solids Feeding Device and 2 System.

~Again referring to FIGURE ~ in lieu of the system of the

5,prior art (see FIGURE 1) wherein the stream of solids plus fluidizin 6gas contact the flash liquid-dilution steam mixture entering reactor 7'28, structurally the apparatus 32B of the subject invention comprise a solids reservoir vessel 33B and a housing 34B for the internal jelements described below. The housing 34B is conically shaped in ~the embodiment oS FIGURE 4 and serves as a transition spool piece 11 " between the reservoir 33B and ~he reactor 32B to which it is 12-,flageably connected via flanges 35B, 36B, 37B and 38B. The par-13 ,,ticular geometry of the housing is functional rather than critical.
14The housing is its~lf comprised of an outer metallic shell 39B, 15preferably of steel, and an inner core 40B of a castable ceramic 16 ,material. It is convenient that the material of the core 40B
17 I,'forms the base 41B of the reservoir 33B.

1 9 ~ ~
~1 \

\

s&w ~96-1~7 ?2~7 1 . 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.
~ 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 33~ to 11 ~ the reaction zone 44B. The conduits 46B are supported by the 12 material of the core 4 OB, and ~erminate coplanarly with a plate 13 45B, which has apertures 47B to receive the conduits 45B. The 14 region immediately below the plate 45~ is hereinafter referred to as a mixing zone 5~ which is also part of the reaction zone 16 44.
17 i As shown in FIGURE 5, an enlarged partial view of the 18 . intersection of the conduit 46~ and the plate 45B, the apertures 19 47B are larger than the outside dimension of conduits 46B, forming l, therebet~Jeen annular orifices 48B for the passage of gaseous feed 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 1 of the conduit wall 51B. In this way the gaseous stream from 24 I,i the chamber 42B is angularly injected into the mixing zone 53B
!1 and intercepts thesolids phase flo~ing from conduits 46B. A
26 1' projection of the gas flow would form a cone shown by dot~ed lines 27 ,j 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 ~I rapidly and uniformly, and form a homogeneous reaction phase.

The mixing of a solid phase with a gaseous phase is a function of Ii " 1 , , S&W
696-147 , i.
l the shear surface between the solids and gas phases, and the 2 flow area. A ratio of shear surface to flow area (S/A) of 3 infinity defines perfect mixing; 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, pene~ration of the

9 phases is obtained and even faster mixing results. By using a plurality of annular gas feed points and a plurality of solid ll feed conduits, even greater mixing is more rapidly promoted, 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 14 the mixinq zone. 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 , 45B, as shown in FIGURES 4 and 5, reduces the flow area, and 18 forms discrete mixing zones 53B. The combination of annular gas l9 ~ addition around each solids feed point and a confined discrete , mixing zone greatly enhances the conditions for mixing. Using 21 i this preferred embodiment, the time re~uired to obtain an 22 j essentially homogeneous reaction phase in the reaction zone 23 ! 44B is quite low. Thus, this preferred method of gas and solids 24 i' addition can ~e used in reaction systems having a residence time '' below l second, and even below lO0 milliseconds. In such r~-26 actions the mixing step must be performed in a fraction of the 27 1, total residence time, generally under 20% thereof. If this 28 criteria is not achieved, localized and uncontrolled reaction 29 occurs which deleteriously affec~s the product yield and dis-'I tribution. This is caused by the maldistribution of solids i! :

!! ~4~

696-i~7, ' 1 l 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 by placing the 4 1 apertures ~7B as close to the walls of the mixing zone 53B as l 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 FIGURE5. In order to 12 prevent movement of conduits 46B by vibration and to retain the 13 uni~orm width of the annular orifices 48B, spacers 56B, are 14 ; used as shown in FIGURE8. However, the conduits 46B are pri-j marily supported within the housing 34B by the material of the 16 , core 4 OB as stated above.
17 '' Referring to FIGURE9, the plug 54B serves to 1~ reduce the flow area and define discrete mixing zones 53B.
i9 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. ~lternatively, a plurality of plugs 54B can be used 23 1, to obtain a mixing zone 53B of the desired geometric con-24 I figuration.
~, Referring again to FIGURE 4, the housing 34B may 26 ',' preferably contain a neck portion 57B with corresponding lining 27 ', 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 '~ portion 57B. This neck portion 57B defines mixing zone 53~, Ij ;l -15-, il ~
s ~;w ~
696-I47 ,1 ' I .
,, and allows complete removal of the housing 34B without dis-!; assembly of the reactor 31B or the solids reservoir 33B. Thus, 3 , installation, removal and maintenance can be accomplished 4 ' easily. Ceramic linings 60B and 62B on the reservoir 33B
' and the reactor walls 61B 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 I~ gas to locally fluidize the solids 63B is supplied by nozzles ~ 64B sy~metrically placed around th~ conduits 46B. Gas to 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 o~ the 17 ' particulate solids through the conduits 46~.
18 ' By locally fluidizing the solids 63B, the solids 19 , 63B assume the characteristics of a fluid, and will flow through , the conduits 46B. The conduits 46B have a fixed cross sectional 21 ll area, and serve as orifices having a specific response to a 22 ~ change in orifice pressure drop. Generally, the flow of 23 ~I fluidized solids through an orifice is a function of the pressure 24 i, drop through the orifice. That orifice pressure drop, in turn, ! is a function of bed hei~ht, bed density, and system pressure.
26 l, However, in the process and apparatus of this 27 l, invention the bulk of the solids in reservoir 33B are not 28 ~I fluidized. Thus, static pressure changes caused by variations 29 'I in bed height are only slowly communicated to the inlet of the ¦' conduit 46B. Also the bed density remains approximately constant I ' , s~w 1 1, until the point of incipient fluidization is reached, that is, 2 point 'la" 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 5 !i FIGUR~ 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 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 bulX 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 essentially 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.
19 Together with the rapid mixing features described above, the present invention offers an integrated system for 21 , feedin~ particulate solids to a reactor or vessell especially 22 , to a TRC tubular reactor wherein very low reaction residence 23 ¦¦ times are encountered.
24 ;' FIGURES loandll depict an alternate preferred embodi-1l ment of the control features of the present invention. In this 26 ,i embodiment the reservoir 33B extends downwardly into the core 27 ~, material 40B to form a secondary or control reservoir 71B. The 28 ,I screen 66B is positioned over the entire control reservoir 71B.
29 ,', The aeration nozzles 64B project downwardly to fluidize essentially ~' these solids 63B beneath the screen 66~. The bottom 41B of the .1 .

, .

, . , ~ ;V~7 S&W
696-1~7 1 , reservoir 33B is again preferably formed of the same material 2 ,l 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 i entering the nozzles 72B. Headers 65B and 74B provide the gas 8 , supply to nozzles 64B and 72B respectively.
9 . The conduits 46B communicate with the reservoir 71B
through leading section 46'B, The leading sections 46'B are 11 ', formed in a block 75~ made of castable erosion resistent ceramic i! i 12 .~ material such as Carborundum Alfrax 201. The bloek 75B i5 13 ~ removable, and can be replaced if eroded. The entrance 75B to 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 without 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'B will still provide a known 1' orifice size and pressure drop response. The conduits 46B
21 ll are completed as before using erosion resistent metal tubes 22 !! 51B, said tubes being set into core material 40B and affixed 23 ¦ to the block 75B.
24 I FIGURE ~lis a plan view of FIGURE ~o along section ¦~ 9~9 showing an arrangement for the nozzles 64B and 72B, and the 26 1l headers 65B and 74B. Gas is supplied to the headers 65B and 74B
27 3, through feed lines 79B and 80B respectively, which extend out 28 ~ beyond the shell. 34B. It is not necessary that the headers be 29 ,j set into the material of the core 40B, although this is a 1~ convenience from the f~brication standpoint. Uniform flow ~&W ! ~ .J2~7 696-1~7 , ~

., , 1 1 distribution to each of the nozzles is ensured by the hydraulics 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 79~ and 80B is regulated via valve means not shown.
S FIGURES 12 and 13 show the pertinent parts of an 6 alternate embodiment of the invention wherein a second gas dis-7 1 tribution assembly for feed gas is contempla~ed. As in the other 8 ~ embodiments, a gas distribution chamber 42B termina~ing in annular 9 ~ orifice 48B surrounds each solids delivery conduit 46B. However, rather than a common ~tall between the chamber 47B and the conduit 11 ~ 46B, a second annulus 83B is formed between the chamber 42B
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 42B and the annular vpening 84B in the annulus 83B at an angle to the flow of solids from the conduits 46~.
16 ; The angular entry of the feed gas to the mixiny zone 53B is 17 I provided by beveled walls 49B and 85B, which define the openings 18 ~l 48B and beveled walls 50B and 89B which define the openings 19 I 84B. Gas is introduced to the annulus 83B through the header l~ 86B, the header being set into the core 40B if convenient.
21 l FIGURE i2 is a plan view of the apparatus of FIGURE
22 ll 13 through section 11-11 showing the conduit openings and the 23 ~l annular feed openings 48B and 84B. Gas is supplied through feed 24 5! lines 87B and 88B to the headers 43B and 86B and ultimately !I to the mixing zones through the annular openings~ Uniform flow 26 ~ from the chambers 42B and 83B is ensured by th~o annular orifices 27 j, 48B and 84B. Therefore, it is not essential that flow dis-2B 1I tribution means such as venturis or orifices be included in 29 !~ the header 43B. The plug 54B is shaped symmetrically to ll define discrete mixing zones 53B.

,, ~

S&W !

,, I ' "
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 coneO With the introduction 8 of the gas phase through the annular orifices 48B at velocities 9 , between 30 and 8no 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 i 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.
, The initial superficial velocity of the two phases in ;
16 ,'the mixing zone 53B is preferably about 20 to 80 ft./sec., 17 I although this velocity changes rapidly in many reaction systems, 18 ;jsuch as thermal cracking, as the gaseous reaction products are 19 ,~formed. The actual averag~ velocity through the mixinq zone 53B
20 11 and the reaction zone 44B is a process consideration, the velocity 21 l being a function of the allowed residence time therethrough.
22 !j By employing the solid feed devic~ and method of the 23 ¦I present inventions, the mixing length to diameter ratio necessary 24 ,11 to in~imately mix the two phases is greatly reduced. This ratio 2; !~ is used as an informal criteria which defines good mixing. Gen 26 l,erally, an L/D~length/dia.3 ratio of from 10 to 40 is required.
27 11 Using the device disclosed herein, this ratio is less than 5, with 28 ~ratios less than 1.0 being possible. Well designed mixing devices 29 ~of the present invention may even achieve essentially complete Imixing at L/D ratios less than 0.5.

, ~96~ 7 l8~)~97 1 (c~ Improved Sequential Thermal Cracking 2 Process.

I Turning now to the sequential cracking process 2C
6 l' f the subject invention, as illustrated in FIGVRE 14, 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 6CI a Il, secondary reactor 8C and downstream equipment. The downstream g 1.
': equi~ment is ~x~rised essentially ~f an indirect heat exchanger lOC,a ~8 S&W ` I
696-147 ~ 97 1 i fractionation tower 12C, and a recycle line 14C from the 2 fractionation tower 12C to the entry of the primary reac~or 3 6C.
4 The system also includes a first hydrocarbon feed ' line 16C, a second hydrocarbon feed-quench line 18C, a transfer 6 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 i from the solids heater 4C. The first or primary reactor 6C
in which the first feed is cracked is at high severity conditions.
11 The hydrocarbon feed, from line 16C, may be any hydrocarbon gas 12 or hydrocarbon liquid in the vapori~ed state which has been used 13 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 polnt gas oils having a ' boiling range between 650 to 950F and compatible combinations 21 ~ of same. These constituents may be introduced as fresh feed 22 j or as recycle streams through the line 14C from downstream 23 ,' purification facilities e.g., fractionation tower 12C. Dilution 24 , steam may also be delivered with the hydrocarbon through lines l 16C and 14C. The use of dilution s~eam 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 mo~ecular weight, hydrocarbon gases, !~
' -22- ' S&W ,l i 696-147 ~ 7 I .
ll i 1 li light hydrocarbon liquids, light gas oils boiling between 350 2 Il and 650F, and combinations of same. These feedstocks offer the I
3 ll greatest increase improvement in selectivity at high severity 4 !' and short residence times~
The hydrocarbon feed to the first reaction zone is 6 ll preferably pre-heated to a temperature of between 600 to 1200F
7 l before introduction thereto. The inlet pressure in the line 16C
8 l is 10 to 100 psig~ The feed should be a gas or gasified liquid. I
9 il The feed increases rapidly in temperature reaching thermal equi-ll librium with the solids in about 5 milliseconds. As mixing of 11 ; the hydrocirbon 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. Tl~e residence time maintained within the primary reactor 1 is about 50 milliseconds, preferably between 20 and 150 milli-16 I seconds, to ensure a high conversion at high selectivity. Typi-17 ll 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 il butadiene, hydrogen, methane, C4 hydrocarbons, distillates such 21 ' as gasoline and gas oils, 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 I weight of feed, and the yield of ethylene for liquid feedstocks ~i is about 25 to 45% by weight of the ~eed, with selectivities of 26 ~ about 2.5 to 4 pounds of ethylene per pound of methane.
27 l~ A second feed is introduced ~hrough the line 18C
28 1 and combines with the cracked ~as from the primary reactor 6C
29 I between the primary reactor 6C and the secondary reactor 8C. The l combined stream comprising the second unreacted feed, and the i, -~3-1/ , 696-147 ~ 97 1 , 1 ~I fir~t 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 1l 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 ~as oils, high ~oiling point gas oils, 8 ¦ and combinations of same.
9 I Supplemental dilution steam may be added with the ~ secondary hydrocarbon stream entering through stream 18C. However, 11 ,l 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 BC. It should 14 ll 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 ll is preferably virgin gas oil 400-650F. The second feed is pre-18 ~, 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 1I pounds of hydrocarbon delivered through the line 18C will quench 22 1l 60 pounds of effluent from the primary reactor 6C. At this tem-23 !i perature level, the cracking reactions of the first feed are 24 ~'l essentially terminated. However, coincident with the quenching lj f the effluent from the primary reactor, the secondary feed 26 il entering through line 18C is thermally cracked at this temperature 27 ¦' ~1500 to 1200F) and pressures of 10 to 100 psig at low severity 28 i, by providing a residence time in the secondary reactor between 29 ,j 150 and 2000 ~illiseconds, preferably between 250 to 500 milli-'' seconds. Typically, the KSF cracking severity in the secondary S&h l 69~-147 Z~7 1 l reactor is about 0.5 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 ' Reaction products from the low severity reaction 20ne comprise 1 ethylene with lesser amounts o propylene and butadiene, hydro-6 , gen, methane, C4 hydrocarbons, petroleum distillates and gas 7 I oils, heavy fuel oils, coke and an acid gas. Minor amounts of 8 ~I other products may also be produced. Feed conversion in this g , second reaction zone is about 30 to 80% by weight of feed, 'l and the yield of ethylene i5 about 8 to 20% by weight of feed, 11 I with selectivities of 2.5 to 4.0 pounds of ethyl~ne per pound 12 ~l of methane.
13 Although the products from the high severity reaction 14 1 are combined with the second feed, and pass through the second jl reaction zone, the low severity conditions in the second reaction 16 l, zone are insufficient to appreciably alter the product dis 17 I tribution of the primary products from the high severity reaction 18 ' zone. Some chemical changes will occur, however these reaction 19 products are substantially stabilized by ~he direc~ quench ,I provided by the second feed.
21 l~ The virgin gas oils normally contain aromatic 22 1l molecules with paraffinic hydrocarbon side chains. For some 23 ' gas oils the number of carbon atoms associated with such 24 1, paraffinic side chains will be a large fraction of the total 1 number of carbon atoms in the molecule, or the gas oil will 26 have a low "aromaticity".
27 'i In the secondary reactor, these molecules will 28 11 undergo dealkylation - splitting of the paraffin molecules, 29 11 leaving a reactive residual methyl aromatic, which will tend ~ to react to form high boilers. The paraffins in the boiliny !
~ -25-696-1~7 ~ 3'7 1, 1 ll range 400 to 650F are separated from the higher boiling 2 ll 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 ,, The total effluent leaves the secondary reactor 7 ' and is passed through the indirect quench means 10C to generate 8 ll steam for use within and outside the system. The effluent is 9 j 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~ Hydro-13 ~l gen and methane are taken overhead through the line 36C. C4 and 14 ll lighter olefins, C5 - 400F and 400-650F fractions are removed " 'rom the fractionator 12C through lines 26C, 28C and 30C re-16 spectively. Other light paraffinic gases of ethane and propane 17 li are recycled through the line 14C to the high severity primary 18 l~ reactor. The product taken through line 28C consists of liquid 19 l~ hydrocarbons boiling between C~ and 400F, and is preferably l exported although such material may be recycled to the primary 21 ~, reactor 6C if desired. The light gas oil boiling between 400 22 I to 650F is the preferred recycle feed, but may be removed through 23 ¦ line 30C. The heavy gas oil which boils between 650-950F is 24 l~ exported through stream 32C, while excess residuim, boiling above 950F is removed from the battery limits via stream 34C.
26 i~ The heavy gas oil and residuim may also be used as fuel within 27 ; the system.
28 '~ In the preferred embodiment of the process, the 29 ~ second feed would be one which is not recommended for high severity operation. Such a feed would be a gas oil boiling Il -26-S&W- , 696-147 ~ 7 1 above 400~F which contain~ a signifi~ant amount of high molecular 2 i' weight aromatic components. Generally, these components have 3 1~ paraffinic side chains which will orm olefins under proper 4 ' conditions. ~owever, even at moderate severity, the dealkylated ;
; aromatic rings will polymerixe to form coke depositsO By 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 ,I 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 eeds 15 ll thereto. The aromatic components of the effluent are removed 16 I from the purification facilities 12C as part of the heavy gas 17 ~l oil product, and either recycled for use as fuel within the 18 ,, system, or exported for further purification or storage 19 I An illustration of the benefits of the process of ~i the invention is set forth below wherein feed cracked and the 21 I resultant product obtained under conventional high severity 22 !i cracking and quenching conditions is compared with the same feed 23 'I sequentially cracked in accordance with this invention.
24 ~ -26 '~ \

28 , \
29 ,' ~ 7-S&~ tf!~ ~ 7 695`-1~7 1 (d) Improved Resldence Time Solld-Gas Separation 2 Device and System.

4 Referring to FIGURE 15 in the subject invention, in lieu i separation zone or curved segment region 36 and the quench area 6 44 of the prior art TRC system (see FIGURE 1), solids and gas en-ter 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 i5 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 mater al from line 17D n Usually, the product gas contains residual solids and is sent to a secondary separator 16 l18D, preferably a conventional cyclone. Quench material should 17 ,Ibe introduced in line 15D in a way that precludes back flow of 18 I quench material to the separator. The residual solids are removed 9 l 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 i 26 need be, sent to down-stream purification units. Stripped solids s~ 7 1 jl removed from the vessel 22D through line 25D are sent to re-2 generation vessel 27D using pneumatic transport gas from line 3 ,1 26D. Off gases are removed from the regenerator through line 28D, 4 11 After regeneration the solids are then recycled to reactor 13D
S ,I via line llD.
6 `l The separator 14D should disengage solids rapidly 7 iI from the-reactor effluent in order to prevent produc~ degradation 8 1 and ensure optimal yield and selectivity of the desired products.
9 il Further, the separator 14D operates in a manner that eliminates 1l or at least significantly reduces the amount of gas entering the i 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 il has been provided between the separator 14D and the stripping '' vessel 22D. Finally, the separator 14D operates so that 16 1 erosion is minimized despite high temperature and high velocity 17 Ij conditions that are inherent in many of these processes. The 18 ll' separator system of the present invention is designed tc meet 19 ' each one of these criteria as is descrihed below.
l~ FIGURE16 is a cross sectional elevational view 21 lj showing the preferred embodiment of solids-gas separation device 22 l~ 14D of the present invention. The separator 14D is provided 23 , with a separator sheli 3'7D and is comprised of a solids-gas 24 'I 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 ,~ ~he inlet 32D and the solids outlet 34D are preferably located 27 '' at opposite ends of the chamber 31D. While the gas outlet 33D
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.
li The separator shell 37D and manways 35D and 36D preferably are ., ' ,, 29 -,, , S&~ 7 6~6-1~7 !l 1 ~l lined wJ~II erosion resi~tent linings 3~D, 39D and 41D re-2 ,, spectively which may be required if solids at high velocities 3 l are encountered. Typical commercially available materials 4 1! for erosion resistent lining include Carborundum Precast Carbofrax D, Carborundum Precast Alfrax 2Ql or their equivalent 6 ,l A thermal insulation lining 40D may be placed between shell 37D
7 ll and lining 38D and between the manways and their respective 8 ,l erosion resistent linings when the separator is to be used g i in high temperature serviceO Thus, process temperatures above ' 1500F. (870C.) are not inconsistent with the utili~ation of ll ,, this device.
12 I FIGURE ~7shows a cutaway view of the separator 13 alor.g section 4-4. For greater strength and ease of construction 14 ~! the separator 14D shell is preferably fabricated from cylindrical~
sections such as pipe 50D, although other materials may, of l6 course, be used. It is essential that longitudinal side walls 17 ll 51D and 52D should be rectilinear, or slightly arcuate as in-18 l~ dicated by the dotted lines 51D and 52D. Thus, flow path 31D
l9 , through the separator is essentially rectangular in cross ,j 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 l~ 52D. Alternatively, baffles, inserts, weirs or other means 24 ,¦ may be used. In like fashion the configuration of walls 53D
'I and 54D transverse ~o the flow path may be similarly shaped~
26 ~ although this is not essential. FIGURE 18 is a cutaway view 27 ~¦ along Section 4-4 of FIGURE16 wherein the separation shell 37D
2~ 1, is fabricated from a rectangular conduit. Because the shell 37D
29 i¦ has rectilinear walls 51D and 52D it is not necessary to adjust I the width of the flow path with a thickness of liningO Linings ~~ -3~-,:

S&W ,i 696-147 ~ 7 1 ! 38D and 40D could be added for erosion and thermal resistence 2 1 respectively.
3 ,l Again referring to FIGURE 1~ inlet 32D and outlets 4 , 33D are disposed normal to flow path 31D (shown in FIGURE 17 so 'I that the incoming mixed phase stream from inlet 32D is required 6 I to undergo a 90 change in dlrection upon entering the chamber.
7 1 As a further requirement, however, the gas phase outlet 33D is 8 , also oriented so that the gas phase upon leaving the separator 9 ,- has completed a 180 change in direction.
I Centrifugal force propels th~ solid particles to 11 l~ the wall 54D opposite inlet 32D of the chamber 31D, while the gas 12 I portion, llaving less momentum, flows through tlle vapor space of 13 ' the chamber 31D. Initially, solids impinge on the wall 54D, 14 1~ but subsequently accumulate to form a static hed of solids 42D, l, which ultimately form in a surface configuration having a curvi-16 1 linear arc 43D of approximately 90. Solids impinging upon 17 l the bed are moved along the curvilinear arc 43D to the solids 18 , outlet 34D. which is preferably oriented for downflow of 19 i solids by gravity. The exact shape of the arc 43D is determined ~ by the geometry of the particular separator and the inlet stream 21 li parameters such as velocity, mass flowrate, bulk density, and 22 il particle size. Because the force imparted to the incoming solids 23 1 is directed against the static bed 42D rather than the separator 24 l' 14D itself, erosion is minimal. Separator effieiency, defined as 1i the removal of solids from the gas phase leaving through outlet 26 ¦~ 33D, is, therefore, not affected adversely by high inlet velocities, 27 l up to 150 ft./sec., and the separator 14D is operable over a 28 l wide rang~ o~ dilute phase densities, preferably between 0.1 29 1 and 10.0 lbs./ft3. The separator 14D of the present invention i achieves efficiencies of about 80%, although the preferred 31 embodiment, di5cussed below, can obtain over 90% removal of solids.

S&~ 7 696-1~7 1 It has been found that separator efficiency is 2 , dependent upon separator geometry inasmuch as the flow path must 3 , be essentially rectangular and the relationship hetween height 4 , H, and the sharpness of the U-bend in the gas flows.
l Referring to FIGURES 16 and17 we have found that 6 for a given height H of chamber 31D, efficiency increases as 7 the 180 U-bend between inlet 32D and outlet 33D becomes 8 I progressively sharper; that is, as outlet 33D is bxouyht pro~
9 l gressively closer to inlet 32D. Thus, for a given H the efficlency of the separa~or increases as the flow path and, hence, residence 11 time decreases. Assuming an inside diameter Di of inlet 32D, 12 the preferred distance CL between the centerlines of inle~ 32D
13 and outlet 33D is less than 4.0 Di, while the most preferred 14 distance between said centerlines is between l.5 and 2.5 Dl.
~ Below 1.5 Di better separation is obtained but difficulty in 16 fabrication makes this embodiment less attractive in most in-17 ' stances. Should this latter embodiment be desired, the separator 18 , 14D would probably require a unitary casting design because 19 ' inlet 32D and outlet 33D would be too close to one another to l, allow welded fabrication.
21 ll It has been found that the height of flow path H
22 , should be at least equal to the value of Di or 4 inches in height, 23 ~' whichever is greater. PRactice teaches that if H is less than 24 ~ Di or 4 inches the incoming stream is apt to disturb the bed , solids 42D, thereby re-entraining solids in the gas product 26 ; leaving through outlet 33D. Preferably H is on the order of 27 twice Di to obtain even greater separation efficiency. While 28 I not otherwise limited, it is apparent that too large an H
29 eventaully merely increases residence time without substantive i' increases in efficiency The width W of the flo~ path is .

S&W
696-147 !l ' 1preferably between 0.75 and 1.25 times Di, most preferably between 2 l 0.9 and 1.10Di.
3 i Outlet 33D may be of any inside diameter. However, 4 , velocities greater than 75 ft./sec. can cause erosion because 'I of r~sidual solids entrained in the gas. The inside diameter 6 l~ of outlet 34D should be sized so that a pressure diferential 7 ~ between the stripping vessel 22D shown in FIGUREls and the 8 I 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 " entering the stripping vessel 22Do The magnitude of the 12 ,I pressure di~ferential between the stripping vessel 22D and the 13 l~ separator 14D is determined by the force required to move the 14 ~I solids in bulk flow to the solids outlet 34D as well as the , ehight of solids in line 16D~ As the differential increases 16 ll the net flow of gas to the stripping vessel 22D decreasesO
17 , Solids, having gravitational momentum, overcome the differential, 18 while gas preferentially leaves through the gas outlet 33D~
19 lj ~y regulating the pressure in the stripping vessel 20 ~ 22D it is possible to control the amount of sas going to the 21 1, stripper. The pressure regulating means may include a check 22 1¦ or "flapper" valve 29D at the outlet of line 16D~ or a pressure 23 ~l control 29D device on vessel 22D~ Alternatively, as suggested 24 ~, above, the pressure may be regulated by selecting the size of l' the outlet 34D and conduit 16D to obtain hydraulic forces 2 6 l¦ acting on the system that set the flow of gas to the stripper 27 ! 32D. ~hile such gas is degraded, we have found that an increase 28 il in separation efficiency occurs with a bleed of ~as to the 'i i 29 ~I stripper of less than 10%, preferably between 2 and 7~. Economic ~, and process considerations would dictate whether this mode of ., .
1' -33- , ! !

696-1~7 .1 .

1 ,, operation should be used. It is also possible to design the 2 l~ system to obtain a net backflow of gas from the stripping 3 1l vessel. This ~as flow should ~e less than 10~ of the total 4 ,I feed gas rateO
By establishing a minimal flow path, consistent 6 l with the ahove recommendations, residences times as low as 7 ,l 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 ll in diameter is possible in many systems where residence times , approaching 0.5 seconds are allowable.
11 i In the preferred embodiment of FIGURE 16, a weir 44D
12 '~ is placed across the flow path at a point at or just beyond the .
13 1¦ gas outlet to establish a positive height of solids prior to 14 l~ solids outlet 34D. By installing a weir (or an equivalent l' restriction) at this point a more stable bed is established 16 ,; thereby reducing turbulence and erosion. Moreover, the weir 17 ,1 44D establishes a bed which has a crescent shaped curvilinear 18 l, arc 43D of slightly more than 90~. An arc of this shape 19 l, diverts ~as towards the gas outlet and creates the U-shaped l gas glow pattern illustrated diagrammatically by line 45D in 21 ~I FIGURE 16. Without the weir 44D an arc sor,lewhat less than or 22 jl equal to 90 would be formed, and which would extend asymptoti-23 l 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 li adversely, the flow pattern of line 61D increases residence time, 27 ll and more importantly, creates greater potential for erosion at 28 11 areas 62D, 63D and 64D.

29 The separator of FIGURE 20 is a schematic diagram of 301 another embodiment of the separator 14D, said separator 14D
'!
: -34-., .
,1 .

S &W ~ 8~ 7 696-147 Ij l!
Il l !I haviny an extended separation chamber in the longitudinal 2 dimension. Here, the horizontal distance L between the gas ,l outlet 34D and the weir 44D is extended to establish a solids 4 bed of greater length. L is preferably less than or equal ! to 5 Di. Although the gas flow pattern 61D does not develope 6 l' the preferred U-shape, a cresceilt, shaped arc i5 obtained 7 j~ which limits erosion potential to area 64D. Embodiments 8 1l shown by FIGURES 19 and 20are useful when the solids loading 9 i 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 14 I the flow path as well as a gravity flow section 67D. Wall 68D
~ replaces weir 44D, and arc 43D and flow pattern 45 are similar 16 1 to the preferred embodiment of FIGURE 16. Because solids accumu-17 'I late in the restrlcted collinear section 66DI pressure losses 18 I,i are greater. This embodiment, then, is not preferred where the 19 , incoming stream is at low velocity and cannot supply sufficient l force to expel the solids through outlet 65D. However, because 21 ~ of the restricted solids flow path, better deaeration is obtained;
22 i,1 and gas losses are minimal.
23 1~ FIGURE 22 illustrates another embodiment of the 24 l separator 14D of FIGURE 21wherein the solids outlet is stepped.
1' Although a weir is not used, the outlet restricts solids flow 26 'j which helps from the bed 42D. As in FIGURE 20, an extended L
27 i distance between the gas outlet and solids outlet may be used.
28 ~ The separator of FIGURE21 or22 may be used in 29 , conjunction with a venturi, an orifice, or an equivalent flow ; restriction device as shown in FIGURE 23. The venturi 69D having i, , l --35-- 1 s&W ` ` 7 69~-147 l'l . ' ' 1 li dimensions Dv (diameter at venturi inlet), DVt (diameter of 2 ,I venturi throat), and ~ (angle of cone formed by projection 3 1 of convergent venturi walls) is placed in the collinear section 4 li 66D of the outlet 65D to greatly improve deaeration of solids.
I The embodimen-t of FIGURE 24 is a variation of the separator 6 ,I 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 l follows the U-shaped pattern of line 45D However, an asymptotic 11 I bed will be formed unless there is a restriction in the solids i! I
12 l 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~ Most preferably, ,I the venturi 69D is placed in collinear section 66D as shown in 16 ,~ FIGURE24 to improve the deaeration of the solids. Of course, 17 ll each of these alternate embodiment may have one or more of the 18 ~ optional design features of the basic separator discussed in 19 i relation to FIGURES 16, 17 and 18.
~'he 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 2~ I figures and listed in the Nomenclature below:
27 i'l CL Distance between inlet and gas outlet centerlines I D Inside diameter of inlet 29 1l Dog Inside diameter of gas outlet D Inside diameter of solids outlet v Diameter of venturi inlet DVt Diameter of venturi throat -36~

S&W
696-1~7 ., , 1 H Height o flow path 2 i' ~ Height of weir or step "
3 ; L Length fxom gas outlet to weir or step as 4 1 indicated in Figure 6 5 , W Width of flow path 6 ~ Angle of cone formed by projection 7 of convergent venturi walls, degrees 8 ,,' Table I
9 Dimensions of Se~ara ors in Examples 1 to 10, inches*
Example 11 '', .
12 ' CL 3.875 3.875 3~875 3.875 3.875 3.875 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 , v ~ ~ ~ _ _ 2 17 l, 9vt ~ ~ - - - _ _ _ _ 1 18 ~, ~ 4 4 4 4 4 4 12 12 7.5 6.7 19 1 Hw 0 75 0.75 0.75 0O75 0.75 0.75 2.25 2.25 0 4~7 20 ,I L 0 2 2 0 0 0 0 0 10 0 21 ,I W 2 2 2 2 2 2 6 6 2 2 22 , 3, degrees - - - - - - - - 28.]
23 * Except as noted 24 ,, Example 1 25 ,l In this example a separator of the preferred 2~ , embodiment of FIGURE 16was tested on a feed mixture of air and 27 silica alumina. The dimensions of the apparatus are specified 28 . in Table I. Note that the distance L from the gas outlet to 29 i, the weir was zero.
'l , 1, -37 1, , , a~
S&W 'i I
696-1471l !
ii 1 l, The inlet stream was comprised of 85 ft.3/min.
2 ¦1 of air and 52 lbs./min. of sllica alumina havlng a bulk densit~
3 1 of 70 lbs./ft3 and an average particle size of 100 microns.
4 ! The stream density was 0.612 lbs./ft.3 and the operation was 'I perf~rmed at ambient temperature and atmospheric pressure.
6 , 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 l.75 inch diameter outlet. A positive seal of solids 9 ll in the solids outlet prevented gas from being entrained in the ', solids leaving the separatorO Bed solids were stabilized by ll '' placing a 0.75 inch weir across the flow path.
12 lll The observed separation efficiency was 89.1~, 13 ~' and was accomplished in a gas phase residence time of approximately 14 ¦¦ 0.008 seconds. Efficiency is defined as the percent removal of 15 ll solids from the inlet stream.
16 ll Example 2 17 l~ The gas-solids mixture of Example 1 was processed 18 , in a separator having a configuration illustrated by FIGURE 2Q.
19 ,l In the example the L dimension is 2 inches; all other dimensions l~ are the same as Example 1. By extending the separation chamber 21 1 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. Separation 24 !1 efficiency for this example was 70.8%.
1l Example 3 26 1,l The separator of Example 2 was tested with an inlet i 27 'j stream comprised of 85 ft.3/min. of air and 102 lbs./min. of 28 1 silica alumina which gave a stream density of 1.18 lbs./ft.3, 29 1' or approximately twice that of Example 2. Separation efficiency 30 1l improved to 83.8%.
."

i ' .

696-147 1 ~ ~8~)29~

1 1 Example 4 2 ,, The preferred separator of Example 1 was tested 3 i' at ~he inlet flow rate of Example 3. Efficiency increased 4 ~i slightly to 91.3~.
S , Example 5 6 ,, The separator of FIGURE 16was tested at the con-7 1, ditions of Example 1. Although the separation dimensions are 8 i 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 1I Example 6 15 '~ Same conditions apply as for Example 5 except 16 , that the solids loading was increased to 102 lbs./min. to give 17 l, a stream density of 1.1~ lbs./ft.3. As observed previously 18 " in Examples 3 and 4, the separator efficiency lncreased with 19 ,' higher solids loading tc 90.6%.
,' Example 7 21 ,' The preferred separator configur~tion of FIGURE ~16 i 22 iI was tested in this Example. Howeverl in this example the appara-23 l~ tus was increased in size over the previous examples by a 24 1l factor of nine based on flow area. A 6 inch inlet and 4 inch i' outlet were used to process 472 ft.3/ min. of air and 661 lbs./min.
26 ,l of silica alumina at 180F. and 12 psig. The respective velocities 27 ! were 40 and 90 ft./sec. The solids had a bulk density of 70 28 ~i lbs./ft3 and the stream density wasl.37 lbs./ ft.3 Distance 29 ll CL between inlet and gas outlet centerlines was 11 inches, or l 1.83 times the inlet diameter; distance L was zero. The bed was .

S&W i~ 3 7 696~ 7 1 I stabilized by a ~.25 inch weir, and gas loss was prevented 2 by a posi~ive seal of solids. However, the solids were 3 ,I collected in a ciosed vessel, and the pressure differential 4 ll was such that a positive flow of displaced gas from the ~l collection vessel to the separatox was observed. This volume 6 , was approximately 9.4 ft.3/min. Observed separator ~fficiency 7 , was 90.0~, and the gas phase residence time approximately 8 , 0.02 seconds.
9 , Example 8 !I The separator used in Example 7 was tested with 11 ll an identical feed of gas and solids. Howeverr 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 ft~3/
14 I min. exited through the solids cutlet at a velocity of 3.6 ;j 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 l 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 the qas outlet (dimension L). ~ weir was not used~ The 22 1, feed was comprised of 85 ft.3/min. of air and 105 lbs./min. of 23 ' spent fluid catalytic cracker catalyst having a bulk density 24 ~, of 45 lbs./ft. and an average particle size vf 50 microns. This 2' gave a strea~ 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 ' in Example 7 there was a positive counter-current flow of 28 l~ 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%.
, -40-l!2~:?7 s&w ., ~i 696-147 jl i, I
Example lO
' The separator of FIGURE23 was tested on a feed 3 comprised of 85 fto3/ 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 l inch in diameter which resulted in an outlet 7 l velocity of 262 ft ./sec. This separator had a stepped 8 l, solids outlet with a venturi in the collinear section of the 9 1~ outlet. The venturi mouth was 2 inches in diameter~ while ~ the throat was l inch. A cone of 281.1 was formed by pro-11 jection of the convergént walls of the venturi. An observed 12 efficiency o~ 92.6~ was measured, and the solids leaving the 13 ' separator were completely deaerated except for interstitial gas 14 , remaining in the solids' voids.

17 ~, \

1 9 ~i \

24 ,, 25 1 \

27 'I \

28 , \
29 li ' ;&W
i96-1~7 1 (e) Im~roved Solids Quench Boiler and 2 Process.

1 As see~ in FIGURE 25,in lieu of guench zone 44, 46 (see 6 FIGURE 1) of the prior art, the composite 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 separator 8E

9 " with a solids return line lOE to the fluid bed-riser 6F and a line 1~ ~ 36E for the delivery of gas to the fluid bed-quench riser~
11 The quench exchang~r 4E as best seen in FIGURES 26 and ~7, 12 .l is 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&~
6~6-1~7 ~ ?~
i , ., , 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 i in effect, the outer wall of the quench exchanger 4E. The 4 ,, inner circles of -tubes 18E and 20E are spaced apart and allow l for the passage of effluent gas and particulate solids there~
6 ~ around. The arrays of tubes 16E, 18E and 20E are manifolded 7 l 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 insure a pressure tight vesselO The quench exchanger inlet hood 12 26E extends from the quench riser 6E to the lower torus 24E.
13 l~ The quench exchanger outlet hood 28E extends from the upper 14 ,I torus 22E and is connected to the downstream piping equipment l by piping such as an elbow 30E which is arranged to cl~liver the 16 !! cooled effluent and particulate solids to the cyclone separa~or 17 ,, 8E.
18 ~, The fluid bed quench riser 6E is essentially a sealed 19 vessel attached in sealed relationship to the quench exchanger l' 4E. The fluid bed-quench riser 6E is arranged to receive the 21 1 reactor outlet tube 36E which is preferably centrally disposed at 22 ' the bottom of the fluid quench riser 6E. A slightly enlarged 23 ~ centrally disposed tube 38E is aligned with the reactor outlek 24 '' 36E and extends from the fluid bed-quench riser 6E into the i' quench exchanger 4E. In the quench exchanger 4E, the centrally 26 1 disposed fluid bed-quench riser tube 38E terminates in a conical 27 opening 40E. The conical opening 40E is provided to facilitate 28 'I nonturbulent transition from the quench riser tube 38E to the 29 enlarged opening of the quench exchanger 4E. It has been found that the angle of the cone ~, best seen in FIGURE26, should 31 be not greater ~han 10 degrees.
, -~3-S fiW~ 7 1 ¦ The fluid bed 42E contained in the ~luid bed quench 2 ll 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 1, solids from the bed 42E. Although virtually any particulate ,i solids can be used to provide the quench bed 42E, it has been 6 ,I found in practice that the same solids used in the reactor are 7 ' preferably used in the fluidized bed 42E. IllustrationS of 8 il the suitable particulate solids are FCC alumina solids.
9 ,' As best seen in FIGURE 23,the opening 48E through ¦I which the fluidized particles from the bed 42E are drawn into 11 ` the quench riser tube 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 ,¦ cone 46E located on the outer surace of the reactor outlet 14 ~.ll tube 36E. In practice, it has been found that the refractory Ij 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
17 ,~ and the refractory cone 46E, is preferably 3-4 square feet for la '' a unit of 50 ~MBTu/HR capacity. The opening is sized to insure 19 il penetration of the cracked gas solid mass velocity of 100 to ,' 800 pounds per second per square foot is required. The amount 21 ll of solids from bed 42E delivered to the tube 38FJ is a function 22 1¦ of the velocity of the gas and solids entering the tube 38E
23 il from the reactor outlet 36E and the size of the opening 48E.
24 ,j In practice, it has been found that the Thermal 1I Regenerative Cracking (TRC) reactox effluent will contain 26 i! approximately 2 pounds of solids per pound of gas at a tem-27 ¦I perature of about 1,400F to 1,600F.
28 'i The process of the solids quench boiler 2E of 29 ¦E'IGURES 25-2~ is illustrated by the following example. Effluent l` from a TRC outlet 36E at about 1,500F is delivered to the quench i -44-il , ., !

6~6-147 1.
1 I riser tube 38E at a velocity of approximately 40 to 100 feet 2 1, per second. The ratio of particulate solids to cracked effluent 3 l er.tering or leaving the tube 36E is approximately two pounds of 4 'I solid per pound of gas at a temperature of about 1,500F. At ,' 70 to 100 ~eet per second the particulate solids entrained into 6 l, the effluent stream by the eductor effect is between twenty five 7 ,l and fifty pounds solid per pound of gas. In S 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 ll composite effluent and solids to 1,030F. The gas-solids mixture ~ ' is passed from the quench riser tube 38E to the quench exchanger 12 , 4E wherein the temperature is reduced from 1,030F to l,OOGF
13 !i by indirect heat exchange with the boiler feed water in the tubes 14 1 16E, l~E, and 20E. With 120,000 pounds of effluent per hour, '~ 50 MMBTUs per hour of steam at 1,500 PSIG and 600F will be 16 1, generated for system service. The pressure drop of the gas 17 , solid mlxture passing through quench exchanger 4E is 1.5 PSI. The 18 ¦I cooled gas-solids mixture is delivered through line 30E to the 19 I cyclone separator 8E wherein the bulk of the solids is removed 1, from the quenched-cracked gas and returned through line lOE
21 ,I to the quench riser 6E.

23 1 ~ ' 24 ,i \
, "
26 ~ \

29 'j ~ ;
30 ll \

il -45-S&W ~ 7 1 (f) Improved Preheat Va orization System.
2 - P _ _ 3 ~ Again referring ~ FIGURE29,in lieu of preheat z~ne 24 (FIGURE 1) 4 of the system 2F of the subject invention is embGdied in a TRC system and is , comprised of essentially a liquid feed heater 4F~ a mlxer 8F for flashing 6 steam and the heated feedstock, a separator lOF to separate 7 jlthe 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 11 ,,hydrocarbon feedstock such as desulfuriæed Kuwait HGO to 12 initially elevate the te~perature of the feedstock~
13 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 delivered from the liquid feed heater 4F by a line 18F.
17 The system separator lOF is to separate the liquid and lB '~vapor produced by flashing in the mixer 8F. Separated gas is 19 ' ~ -S& ~ I' 696-1~7 ~

Il l 1 lldischarged through a line 22F from the separator overhead and 2 Ijthe remaining liquid is discharged through a line 26F~
3 ¦¦ 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 ! The second mixer 14F is provlded to Elash the vaporized 7 ll gaseous discharge from the vapor feed superheater 12F and the a ,, liquid bot~oms 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 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 30Fo 14 I,l In the subject process, the heavy oil liquid hydro-Iicarbon 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 delivered through the line 18 l~18F to the mixer 8F. Superheated steam from the line 6F if 19 l,mixed with the heated heavy oil hydrocarbon feedstock in the ilmixer 8F and the steam-heavy oil mixture is flashed to about 21 ,,700 to 800F. For lighter feedstock the flashing temperature 22 lwill be about 500 to 600F., and for heavier feedstock the 23 ,~ 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 ., 1, 26 " 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 of about 700~ to 29 ,800F. The temperature level and percent of hydrocarbon I vaporized are determined within the limits of equipment fouling ,1 ,i f, s 69~-147 ll I
iL8~ 7 1 ,Icriteria. 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 30g and 7Q% 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 th~ system vapor feed superheater 12F wherein it is heated to 9 i' about 1,030F. The heated vapor is taken from the vapor feed Isuperheater 12F through the line 24F and s~nt to the second mixer 11 'j 14E'. Liquid bottoms ~rom the separator 10F is also delivered 12 ' to the second mixer 14F and the vapor-liquid mix is flashed in 13 l~ the mixer 14F to a temperature of about 1,000F.
14 ' The flashed vapor is then sent downstream through the ,. :
1, line 28F to the knockout drum 16F for removal of any liquid 16 ~l from the vapor. Finally, the vaporized hydrocarbon feed is 17 'I sent through the line 30F to a reactor.
18 1 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 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, SUS 210 F 44.2 26 ;' Sulfur .10 Pour Point, F +90.0 i Nitrogen .047 Carbon Residue, Ramsbottom .09 27 ! Nickel .10 Aniline Point, C 87.0 ! Vanadium .10 28 , 29 ' ,, .

''I
~, ~
1 Distillation i, i 21I Vol. 96 3 IlIBP
l 10 669.2 4 30 755.6 `50 820.4 5',70 874.4 9~4.6 6 EP 1,005.8 8 t3,108 pounds per hour of ~he 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 11 ! superheated steam at 1,100F. is simultaneously delivered to the 12 Imixer 8F. The pressure in the mixer is 50 psia.
13 ' The superheated steam and Heavy Gas Oil are flashed in 14 I the mixer 8F to a temperature of 760F. wherein 60 of the H~avy ,,Gas Oil is vaporized.
16 ,I The vapor and liquid from the mixer 8F are separated 17 l'in the separator 10F. 622 pounds per hour of steam and 1,864.8 18 I pounds per hour of hydrocarbon are taken in line 22F as overhead 19 , vapor. 1,243.2 pounds per hour of hydrocarbon are discharged lthrough the line 26F 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 ,l24F to ~he mixer 14F. The mixer 14F is main~ained at 45 psia.
l The 1,243.2 pounds per hour of liquid at 760F. and 26 the vaporous mixture of 622 pounds per hour o~ steam and 27 '~1,864.8 pound per hour of h~drocarbon are flashed in the ~ixer 28 l14F to 990F.
29 -' The vaporization of the hydrocarbon is effected with a ,!steam to hydrocarbon ratio of 0.2. The heat necessary to vaporize Il , _~ g_ s~w ~ q~
6 96-l~ 7 1!

1 1I the hydrocarbon and generate the necessary steam is 2.~24 MM
2 IBTU/hr.
3 ll The same 3,108 pounds per hour of Nigerian Heavy Gas 4 Il Oil feedstock vaporized by a conventional flashing operation llrequires steam in a 1:1 ratio to maintain a steam temperature 6 of 1,434F. The composite heat to vaporize the hydrocarbon and 7 I generate the necessary steam is 60541 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.
,lis required, which temperature is effectively beyond design limitations.
12 i 1 13 ¦1 SUMMARY
14 , With reference to the new and improved separation 15 , (see FIGURES 15-24), it is noted that short residence time ~, 16 l~favors selectivity in C2~I4 production. This means that the 17 il reaction must he quenched rapidly. When solids are used, they 18 i! must be separated from the gas rapidly or quenched with the gas.
19 ¦ If the gases and solids are not separated rapidly (but l separated) as in a cyclone, and then quenched, product 21 degradation will occur. If the total mix is quenched, to avoid 22 ilrapid separation, a high thermal inefficiency will result since 23 i! all the heat of the solids will be rejected to some lower 24 11 level heat recovery. Thus, a rapid high efficiency separator, li according to the subject invention, is optimal for a TRC process.
26 ~ Similarly, in connection with the subject solidc 27 ifeed device (see FIGURES 4-13), it is noted that in order to 28 ~ifeed solids to an ethylene reactor, the flow must be controlled 29 jl to within ~2 percent or cracking severity oscillations will be 1I greater than that presently experienced in coil cracking. The ~l -50-., , 1 subject feed device (local fluidization) minimizes 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 a thus uniquely suited to short residence time reactions. Further, for short residence time reactions, the rapid and intimate Ç mixing are critical in obtaining good selectivity (minimize 7 mixing time as a percentage of total reacting ti~e). Both of 8 the features permit the TRC to move to shorter residence times 9 which increase selectivity. Conventional fluid hed 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 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 18 600~ 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 However, instead of an oil quench, a circulating solids stream 24 could be used to quench the effluent. ~s 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 ~Ihich 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 696-147 ~ i 8~!2 9 7 1 an oil that is too viscous). The use of solids can be done for 2 both TRC or a coil, but it is especially suîted to a TRC since 3 it already uses solids. ~uring quenching, the coke accumulates ~ 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 TRC it could use the regenerator that already exists.
7 With reference to the preheat vaporization system of 8 the subject invention (see FI~URE 29), it is noted that 9 the TRC has ~axi~um economic advanta~es when cracking heavy 1~ feedstocks (650E~+ boiling point) at low steam dilutions.
11 Selectivity is favored by rapid and intimate mixing. Rapid 12 and intimate mixing is best accomplished with a vapor feed 13 rather than a liquid feed.
14 Finally, with reference to the sequential crasking system of the invention (see FI~URE 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 addition to any yield advantages. It can 19 be applied to both TRC 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 While there has been described what is considered to be preferred embodiments of the invention, varia-tions and modif-26 ications therein will occur to those skilled in the axt once 27 they become acquainted with the basic concepts of the invelltion.
28 Therefore, it is intended -that the appended 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 (11)

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 reaction chamber is between 1300° and 2500°F and wherein the hydrocarbon fluid feed or the hydrosulfurized 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 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 hydrocarbon 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.

2. A 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.

3. A 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.

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

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

6. A process as in Claim 3 wherein the fraction returned to the first zone is light paraffinic gases of ethane and propane.

7. 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.

8. 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.

9. A process as in Claim 4 wherein the kinetic severity function in the first zone is about 3.5.

10. A process as in Claim 4 wherein the kinetic severity factor is about 0.5 at about 300 to 400 milliseconds.

11. 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.