CA1194504A - Fuel gas generation - 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

This application is a division of co-pending application Serial No. 361,734 filed September 30, 1980.

The present invention relates to improvements in Thermal Regenerative Cracking (TRC) apparatus and process, as described in U.S. Letters Patents Nos. 4,061,562 and 4,097,363 to McKinney et al.

According to one broad aspect, the present invention relates to a TRC process wherein the temperature in the cracking zone is between 1300° and 2500° F and wherein the hydrocarbon feed or the hydrodesulfurization residual oil along with the entrained inert solids and the diluent gas are passed through the cracking zone for a residence time of 0.05 to 2 seconds, the improvement comprising the process for generating fuel oil and removing coke deposits on said solids comprising generating fuel gas having a high CO to CO2 ratio and a high molal ratio of H2O to H2 from fuel, air and steam; delivering the fuel gas to a tubular transfer line; delivering particulate solids having coke deposits thereon to the tubular transfer line; mixing the fuel gas and the particulate solids to reach an equilibrium temperature; and removing coke from the particulate solids with heat and steam from the particulate solids-fuel gas mixture.
According to another broad aspect, the present invention relates to a TRC process wherein the temperature in the cracking zone is between 1300° and 2500°F. and wherein the hydrocarbon feed or the hydrodesulferization residual oil along with the entrained inert solids and the diluent gas are passed

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through the cracking zone for a res:idence time of 0.05 to 2 seconds, the improvement comprising the process for generating fuel oil and removing coke deposits on said solids comprising generating a fuel gas from fuel and air; delivering the fuel gas to a transfer line; mixing the particulate solids with the fuel gas in the transfer line to elevate the temperature oE the solids; and combusting the fuel ~as in the transfer line to elevate the temperature oE the solids and remove the coke from the solids.

According to a still further broad aspect, the present invention relates to a TRC system wherein the temperature in the cracking ~one is between 1300~ and 2500F~ and wherein the hydrocarbon feed or the hydrodesulfurization residual oil along with the entrained inert solids and -the diluen-t gas are passed throuyh a cracking zone Eor a residence time of 0.05 to 2 seconds, the improvement comprising a system for heating and removing coke from the par-ticulate solids comprising means Eor generating fuel gas having a high molal ratio of H20 to H2 from fuel, air and steam; a transfer line; and means to mi~ the Euel gas and particula-te solids in the transEer line, whereby the fuel gas elevates the temperature of the particulate solids b~
intimate contact therewith and the steam in the fuel gas removes the coke Erom the solids during travel through said transEer line.

According -to yet another broad aspect, the present i.nvent..ion re].a-tes to a process for generating Euel gas and removing coke deposits on solids comprising genera-ting fuel gas having a high C0 l:o C02 ra-tio and a high molal ratio of H20 to -2a-, H2 Erom Euel, air and steam; delivering the fuel gas to a tubular transfer line, delivering particulate solids having coke deposits thereon to the tubular transfer line, mixing -the Euel gas and the particulate solids to reach an equilibrium temperature, and removing coke ~rom the particulate solids with heat and steam from the particulate solids-fuel gas mixture.

According to another broad aspect, the present invention relates -to a process for heating particulate solids and removing coke from the solids beore delivering the solids to a thermal reaction chamher comprising generating a fuel gas froln fuel and air; delivering the fuel gas to a transfer line;
mixing the particulate sGlids with the fuel gas in the transfer line to elevate the temperature of the solids; and combus-ting the fuel gas in the transfer line to elevate the tempera-ture of the solids and remove the ~oke from the solidsO

According to yet another broad aspect, the present invention relates to a system for heating and removing coke from particulate solids compr1sing means for generating fuel gas having a high molal ratio of H2O to H2 from fuel, air and steam; a transfer line; and means to mix the fuel gas and particulate solid~ in the transfer line whereby the fuel gas eleva-tes the temperature of the solids by intimate contact therew:ith and the stearn in the fuel gas removes the coke ~rom the solids during travel t:hrough the transfer line.

This invention will be more fully understood from reference to the accompanying drawings in which:

-2b-FIGURE 1 is a schema-tic diagram of a TRC system and process according -to the prlor art.
FIGURE 2 l.s a schematic diagram of the fuel gas generation system and process of the subject inventionO
FIGURE 3 i5 an alternative embodiment wherein the fuel.
gas is burned to fl.ue gas to provide additi.onal heat for the particulate solids.
FIGURE ~ is a cross-sectional elevational v.iew of -the solids feeding device and system as applied to tubular reactors and for use with gaseous feeds.
FIGURE 5 is an enlarged view o* -the intersection of the solid and gas phases within the mixing zone of the reaction chamber.
FIGURE 6 is a top view of the preferxed plate geometry, said pla-te serving as the base of the gas distribu-tion chamber.
FIGURE 7 is a graph of the relationship be-tween 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 isornetric view of the plug which extends into the mixing zone to reduce flow area.
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 12 is an alternate embodimen-t of the preferred invention wherei.n a second :Eeed gas is contemplated.

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696~147 1 FIGURE 13 is a view of the apparatus of FIGURE 12 2 through line 13-13 of FIGURE 12.
3 FIGURE 14 is a schematic diagram of the sequential

4 thermal cracking process and system of the present invention.
FIGURE 15 is a schematic flow diagram of the separation svstem of the present invention as ap?ended to a 7 typical tubular reactor.
8 FIGURE 16 is a cross sectional elevational view of g the preferred e~bodiment of the separator.
FIGURE 17 is a cutaway view throu~h sec~ion 17-17 11 of FIGURE 16.
12 YIGURE 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 i6 the present invention indicating gas and solids phase flow 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 FIGI]RE 21 is a sketch of an alternate er,~odiment of 22 the separation device wherein a stepped solids outlet is emplo~ed, 23 said outlet having a section collinear with the flow path as well 24 as a gravit~ flow section.
FIGU~E 22 is a variation of the embodiment of FIGUR~
26 21 in which the solids outlet of FIGU~E 20 is used, but is not 27 stepped.
28 FIGURE ?3 is a sketch of a variation of the 29 separation device of FIGURE 8 wherein a venturi restriction is incorporated in the collinear section of the solids outlet.

_4_ (S&W) 1 FIGURE 24 is a variation of the e~bodiment of 2 FIGURE 23 oriented for use with a riser type reactor.

3 FIGUR~ 25 is a sectional elevational view of the 4 solids quench boiler using the quench ri.ser;

FIGURE 26 ls a detailed cross sectional elevational 6 view of the quench exchanger of the system;

7 FIGU~E 27 is a cross sectional plan view taken 8 through line 27-27 of FIG~RE 26;

9 FIGURE 28 is a detailed drawing of the reactor outlet and fluid bed quench riser particle en-try 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.
~1 22 1l Referring to FI~URE 1, in the prior art TRC process~

23 ¦and system, thermal cracker feecl oil or residual oil, with or 24 l~ithout blended distillate heavy gas, entering through line 10 `25 'lancl hyclrogen entering through line 12 pass through hydrodesulfurized 26 ¦'zone 14. Hyd.rosulfurization effluent passes through line 16 anc 27 I,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 1!

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., i 1 1 liquid passes through preheater 2~, is admixed with dilution 2 I steam entering through line 26 and then flows to the bo~ttom 3 1 of thermal cracking reactor 28 through line 30~
4 1l A stream of hot regenerated solids is charged ' through line 32 and admixed with steam or other Eluidizing gas 6 1 entexing through line 34 prior to entering the bottom of riser 7 ,l 28. The oil, stearn and hot solids pass in entrained flow up-8 I wardly through riser 2~ and are discharged through a curved g ,. segr,ent 36 at the top of the riser to induce centrifugal separ ation of solids from the effluent stream. A stream containing 11 il most of the solids passes through riser discharge segment 38 and 12 ~ can be mixed, if desiredj with make-up solids entering through 13 i line 40 before or after entering solids separator~stripper 42.
14 Another stream containing 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 i7 ~. 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 1l are removed from solids separators 42 and 48 through lines 54 21 `I and 56, respectively, and then combined in line 58 for passage 22 1I to a secondary quench and product recovery train, not shownO
2 3 1I Coke-laden solids are removed from solids separators 42 and 48 24 I through lines 60 and 62, respectively, and combined in line 64 1I for passa(3e to coke burner 66~ ~f required, torch oil can be 26 1, added to burner 66 through line 68 wilile stripping steam may be 27 1l added through line 70 to strip cornbust.ion gases from the heated 2B 11' solids. Air is charged to the burner through line 6g. Combustion 29 '¦ gases are removed from the burner through line 72 for passage to heat and energy recovery systems~ not shown, while regenerated Il .
~ ~6-,, S&W
69~147 ., ~

1 ~ hot solids which are relatively free of coke are xemoved from 2 the burner tnrough line 32 for recycle to riser 28~ In order 3 l; to produce a cracked product containing ethylene and molecular 4 I hydrogen, petroleum residual oil is passed through the catalytic I hydrodesulfurized zone in thepresence of hydrogen at a tem-6 perature between 650F and 900F, with the hydrogen being 7 chemically combined with theoil during the hydrocycling step.
3 The hydrosulfurization residual oil passes through the thermal cracking zone together ~ith the entrained inert hot solids functioning as the heat source and a diluent gas at a temperature 11 l, between about 1300F and 2500F for a residual time between 12 1 about 0.05 to 2 seconds to produce the cracked product and 13 1 ethylene and hydrogen For the production of ethylene by 14 ` thermally cracking a hydrogen feed 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 1 with diluent gas and entrained inert hot gases are passed 18 , through the cracking zone at a temperature between 1300F and 19 2500F 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 , severity corresponds to a methane yield of at least 12 weigh~
22 l percent based on said feed oil. Quench cooling of the product 23 !~ immediately upon leaving the cracked zone to a temperature 24 ~I below 1300~F ensures that the ethylene yield is greater than ~25 l, the methane yield on a weight basis.
26 `'== -~
27 \

_ _ . ... .. . . . . . _ _ _ _ .

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696-1~7 1 (a) Improved ~uel Gas Generation For Solids .. . ..
2 Heati.ng.

-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 (FIG~ 1). Particulate solids and hydrocarbon feed gas 7 l'enter a t~ular reactor 13A throuah lines llA and 12A respectively. The cracked 8 ' effluent frornthe tubular reactor 13A is separated fr~nthe partic~ate solids 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 separatorO The residual 12 solids are removed from the quenched product gas in a secondary 13 separator lBA and delivered to the solid stripper 22A. The soiids-14 ' free product gas is taken overhead from the secondary separator ~l18A through line 19A.

16 ~ . _ __ s ~

., .
, 1 , The partlculate solids in the solid stripper 22A, 2 having delivered heat during the thermal cracking in the tubular 3 ~ reactor 13A9 must be reheated and returned to the tubular reactor 4 1 13A to continue the cracking process~
The particulate solids prior to being reheated/ are 6 ,stripped of gas in the solid stripper 27A by steam delivered to 7 I the solid stripper 22A through line 23A.
8 After the particulate solids have been stripped of g gas impurities in the solid stripper 22A, the particulates solids are at a temperature of about 1,450F.
ll ' 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 combustion 14 , vessel 30A. Pre-heaters 32A, 34A, and 36 are shown in fuel line lS j~ 38A, air line 40A~ and steam line 42A respectively.
16 rl~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 transfe~
,~ line 44A is sized to afford sufficient residence time for the 21 I, steam emanating from the combustion vessel 30A to decompose ~y 22 l, the reaction with carbon ln 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
¦1 line 26A is provided for pneumatic transport gas if necessary.
26 ,l A separator, such as a c~lclone separator 46A is 27 1 provided to ~separate the heated decoked particulate solids from 2~ ¦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.

, I
_ 9 _ , S&W
6g6~1~7 ll l 1 In the process, fuel, air and steam are delivered throuyh lines 38A, 40A and 42A 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 CO~
and at least an equivalent molal ratio of ~120 to H2. The H20 6 to H2 ratio of the fuel gas leaving the combustioll vessel 30A
7 is above the ratio required to decompose steam by reaction with 8 , carbon in the presence of hydrogen and ~.o insure that the net 9 ;; carbon in the fuel gas particulate solids will mix will he removed before reaching the separator 46A.
11 ' The fuel gas from combustion vessel 30A at a 12 temperature of about 2,300F. is mixed in the tuhular vessel 44A
13 ~ with stripped particulate solids having a temperature of about 14 1,450F. The particulate solids and fuel gas rapidly reach ar.
' 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 stear.-l, 19 ll As a result, the paxticulate solid-fuel gas mixture is cooled by about 30F. l.e., from 1,780~F~ to 1,750F~
21 , The particulate solid-fuel gas mixture is separated 22 in the separator 46A and the fuel g.as is taken at 1,750F, through 23 ' line SOA. The ~articulate solids are dellvered to the hot soli~s 24 1l hold vessel 27A at 1,750F. and then to the tubular reactor 13A. .
.25 In the alternative embodiment of the invention 26 1, illustrated in FIGURE 3, onl.y fuel and air are delivered to the 27 combustor 30 and burned to a temperature of about 2,300F~ to 28 1I 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. Thereaf~er air is delivered i to the transfer line 44~ through a line 54A. The fuel gas in !

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6 ~ 6- 1 4 7 1 I the line 44A is burned to elevate the ~em~erature of ~he partlcu-2 Il;late solids to about 1,750F. The resultant Elue gas is separated 3 l,from the hot solids in the separator 46A and discharged -through 4 I the line 52A. The hot yarticulate solids are returned to the I,system to provide reaction heat.
6 ,l An example of the system and process of FIG~RE 3 7 , follows: 7,000 pounds per houx of fuel pre-heated to 600F. in 8 the preheater 32A and 13 MM SCFD of air heated to 11000F. are 9 burned in the combustor 30A to 2,300F. to produce 15.6 MM SCFD
' of fuel ~as.
11 1; .The 15 MM SCFD of fuel ~as at 2,300F. is ~xed ln 12 Ithe transfer line 44A with 1 ~ pounds per hour of ~tripped particu-13 ,late solids from the solids stripper 22A. The particulate sol1ds 14 have 1,600 pounds per hour of carbon deposited thereon. The com-.posite fuel gas-partlculate solids gas mixture reaches an 16 equilibrium temperature of 1,480F~ at 5 psig in about 5 milll-17 ~Iseconds~ Thereafter, 13 ~l SCFD of air is delivered to the 18 l,transfer line 44A and the 15.6 ~M SCFD of fuel gas is burned 19 ~Iwith ~he air to elevate the solids temperature to 1,750F. and il burn the 1,600 pounds per hour of carbon from the particulate 21 ,,solids~
22 ¦¦ The combusted ~as from the transfer line 44A is 23 llseparat~d from the solids in the separator 46A and discharged 24 ¦1 as flue ~as.
~25 j!~

,, 27 11 \
2a ~

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(b) Improved Sollds Feeding Device and 2 System.

:
4 Again xeferring to FIGIJRE 4 in lieu of the system of the 5prior art ~see FIGURE 1) wherein the stream of solids plus flu1dizir 6;: gas contact the flash liquid dilution steam mixture entering reactor 728, structurally the apparatus 32B of the subject invention compris~
a solids reservoir vessel 33B and a housing 34B for the internal elements d~escribed below~ The housing 34B is conically shaped in the embodiment o. FIGURE 4 and serves as a transition spool piece 11 1 between the reservoir 33B and the 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~
14 The housing is itself comprised of an outer metallic shell 3913, . preferably of steel, and an inner core 40B of a castable ceramic 16 material. It is convenient that the material of the core 40B
17 forms the base 41B of the reservoir 33B.

19 .

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2g .,, ,, , _ _ . _ _ . , . _ _ _ _ . _ _ _ _ .

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~g6-1~7 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 l of unitary construction, it is preferred that the base separating the chamber 42B from reaction zone 44B be a removable plate 45B~
6 One or more conduits 46B extend downwardly from the xeservoir 7 33B to the reaction zone 44B, passing through the base 41B, 8 and the cham~er 42B. The conduits 46s 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 ro ll ! the reaction zone 44B. ~he conduits 46B are supported by the 12 material of the core 40B, and terminate coplanarly with a pla~e 13 45B, which has apertures 47B to receive ~he conduits 46B . The 14 region immediately below the plate 45B is hereinaf.er referred to as a mixing zone 5~ which is also part of the reaction 70ne 16 44.
17 As shown in FIGURE 5, an enlarged partial view 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 therebetween annular orifices 48B for the passage of gaseous reec 21 from the chamber 42B. Edges 49B of the apertures 47B are pre--22 ferably convergently beveled, as are the edges 50B, at ~che tip 23 1 of the conduit wall 51B. In this way the gaseous stream from 24 1I the chamber 42B is angularly injected into the mixing zone 53B
25 1l and intercepts thesolids phase flowing from conduits 46B~ A
26 l projection of the gas flow would form a cone shown by dotted lines 27 ,l 52B -the vert:ex of which is beneath the flow path of the solids~
28 " By introduci.ng the gas phase angularly, ~he two phases are mixed 29 1l rapidly and uniforml.y, and form a homogeneous reaction phase~
The mi.xing of a solid phase with a gaseous phase is a function of ., . 1 ~
. .
., S&W
~96-1~7 1 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 9 I tne solids are introduced at the wall of the reaction zone. In the system of the present invention, the gas stream is lntro-6 duced annularly to the solids which ensures high shear surfaceO
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. By using a plurality of annular gas feed points and a plurality of solld 11 feed conduits, even greater mixing is more rapidly promoted, 12 since the surface to area ratio for a constant solids ~low area 13 is increased. Mixing is also a known function of the L/D of 14 the mixing 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 19 addition around each solids feed point and a confined discrete 1l mixing zone greatly enhances thelconditions for mixing. Using 21 ~ this preferred embodiment, the time required to obtain an 22 !: essentially homogeneous reaction phase in the reaction zone 23 1 44B is quite low. Thus, this preferred method of yas and solids 2~ ~l addition carl be used in reaction systems having a residence tlme 25 1I below 1 second, and even below 100 milliseconds. In such re-2~ actions the mixing step must be perEormed in a fraction of the 27 l, total residence time, generally under 20% thereof. If this 28 criteria is not achieved, localized and uncontrolled reaction 29 occurs which deleteriously affects the produc~ yield and dis-,I tribution. This is caused by the maldistribution of solids ,, 1. .

l ~
696-i41 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 47B as close to the walls of the mixing zone 53s as ~ possible FIGURE 6 shows the top view of plate 45Bhavingin 6 complete circular apertures 47B symmetrically spaced along ~he 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 ahove. In this embodiment, the apertures 47B are completed by the side walls 55B of gas 11 ` distribution chamber 42B as shown in FIGURE 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 18 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 ~ zone 44B. Alternatively, a plurality of plugs 54B can be used 23 I to ohtain a mixing zone53B of the desired geometric con-24 1 figuration.
ll Referring again to EIGURE4, the housing 34B may 26 l; preferably contain a neck portion 57B with corresponding lining 27 l 58B of the castable ceramic material and a flange 37B to cooperace 2~ with a flange 38B on the reaction chamber 31Bto mount the neck 29 'I portion 57B. This neck portion 57B defines mixing zone 53B, Ij Il -15-. . .

s~w 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 1 installation, removal and maintenance can be accomplished 4 ~l easily. Ceramic linings 60B and 62B on the reservoir 33B
I and the reactor walls 61B respectively are provided to prevent ~ erosion.
7 I 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 63Bis supplied by nozzles ;64B symmetrically placed around the conduits 46B~ Gas to 11 ,I nozzles 64Bis supplied by a header 65B~ Preferably, the header 12 65B is set within the castable material of the core 40B~ but 13 I 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 64Bto prevent debris and large particles from 16 1 entering ~che reaction zone 44B or bloc)cing the passage of the 17 part:iculate solids through the conduits 46Bo 18 , By locally fluidizing the solids 63B~ the solids 19 63B assume the characteristics of a fluid, and will flow throlgn 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 I change in orifice pressure dropO C;enerally, the flow of 23 ~I fluidized solids through an orifice is a function of the pressure 24 ll drop through the orifice. That orifice pressure drop, ln turn~
2s !~ is a function of bed height, bed density, and system pressureO
26 1! However, in the process and apparatus of thls :27 !' invention t:he bulk of the solids in reservoir 33B are not 28 I fluid.ized. Thus, static pressure changes caused by variations 29 l in bed height are only slowly communicated to the inlet of the 30 1! condLlit 46E~. Also the bed density remains approximately constant I, --1 6--1.

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696-1~7 1 l~ until the point of incipient fluidization is reached, that is, point "a" of FIGURE 7. In the present invention, however, it 3 l is essential that the amount of aeration gas be below that 4 , amount. Any aeration gas flow above that at point ~'a" on S 1 FIGVRE 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 fluldized bed can be varied. Accordingly, the pressure drop ; across the orifice is regulated and the flow of solids thereby ~ regulated as shown in FIGURE 7. As gas flow rates below 11 incipient fluidizatlon, significant pressure increases 12 above the orifice can be obtained without fluidizing the bul~ Gf 13 ~ the solids. Any effect which the bed height and the bed denslty 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 j 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 2i , feedin~ particulate solids to a reactor or vessel, especially 22 to a TRC tubular reactor wherein very low reaction residence 23 ll times are encountered.
24 ' FIGU~ES loandll depict an alternate preferred embodi jl ment of the control features of the present invention~ In thls 26 l embodiment the reservoir 33B extends downwardly into the core 27 1, material 40B to form a secondary or control reservoir 71s. The 28 ~I screen 66B :Ls positioned over the entire control reservoir 71s.
29 I The aeration nozzles 64B project downwardly to fluidize essentiall~
1' these solids 63B beneath the screen 66~. ~he bottom 41~3 of the .1 ~ 3~

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696-1~7 Ij 1 reservoir 33B is again preferably formed of the same rnaterial 2 ,l as the core 4OB.
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 ~ay have accumulated 6 ~ on the screen 66B. Porous stone filters 73B prevent solids frorn 7 entering the nozzles 72Bo Headers 65B and 7~1B provide the gas 8 , supply to nozzles 69B and 72B respectively.
9 The conduits 46B communicate with the reservoir 7 lB
I through leading section 46'B~ The leading sections 46iB are 11 formed in a block 75B made of castable erosion resistent ceramlc 12 material such as Carborundum Alfrax 201~ The block 75B is 13 ~ removable, and can be replaced if eroded. The entrance 75B tG
14 each section 46 IB can be sloped to allow solids to enter more , easily. In addition to being erosion resistent, the block 16 75B provides greater longevity because erosion may occur ~7ithout 17 i loss of the pxeset response function. Thus, even if the conduit @~ 18 leading sections 46~B erode as depicted by dotted lines 77B~
19 the remaining leading section 46lB will still provide a knowr.
~ orifice size and pressure drop response. The conduits 46B
21 1' are completed as before using erosion resistent metal tubes 22 !~ 51Br said tubes being set into core material 40B and affixed 2 3 1I to the block 75~
24 I FIGURE llis a plan view of ~IGURE ~o along section .25 1, 9-9 showing an arrangement for the nozzles 64~ and 72B~ and the 26 ~j headers 65B and 74B. Gas is supplied to the headers 65B and 74B
27 I through feed lines 79B and 80B respectively, which extend out 28 ~ beyond the shell 34B~ It is not necessary that the headers be 29 'I set into the material of the core 40B, although this is a 1~ convenience from the fabrication standpoint. Uni~orm flow ., 6g6-147 1 1 distribution to each of the nozzles is ensured by the hydraulics 2 of the no~zles themselves) and does not require other devices 3 j such as an orifice or venturi. The gas supplied to feed lines 4 1 79B and 80B is regulated via valve means not shown.
S ; FIGURES 12 and 13 show the pertinent parts of an 6 1 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 distrihution chamber 4 2B terminating in annular 9 orifice 48B surrounds each solids delivery conduit 46B~ Howeverl ' xather than a common ~lall between the charnber 47B and the cor,dult 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 opening 84~3 in ~he annulus 83B at an angle to the flow of solids from the conduits 46B.
16 The angular entry of the feed gas to the mixin~ zone 53Bls 17 , provided by beveled walls 493 and 85B, which define the openings 18 ~48B and beveled walls 5013 and89B which define the openlngs lg '!84B. Gas is introduced to the annulus 83B through the header ,i86B, the heade~ being set into the core 40B if convenient.
21 ,, FIGURE i2 is a plan view of the apparatus of FIGURE
22 1l 13 through sectlon 11-11 showing the conduit openings and the 23 l' annular feed openings 48B and 84B. Gas is supplied through feed 24 '¦ lines 87B and 88Bto the headers 43B and 86B and ultimately 2S 1I to the mixing zones through the annular openingsO Uniform flow 26 ¦ from the charn~ers 42B and 83Bis ensured by the annular orifices 27 j, 48B and 89E~. Therefore~ it is not essential that flow dis 28 l tribution means such as venturis or orifices be included in 2g !I the header 43B~ The plug 54Bis shaped symmetrically to 30 1I define discrete mixing zones 53B.
I

s~ s~
696-147j ,. Il il .
1 1 Mixing efficiency is also dependent upon the velocitles 2 of the gas and solid phases. The solids flow through the condui~s 3 I 46B in dense phase flow at mass velocities from preferably 200 4 ~ to 500 pounds/sq. ftO~sec, although mass velocities between 50 and 1 1000 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 f-t./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.
The initial superficial velocity of the two phases in 16 Ithe ~ixing zone 53B is preferably about 20 to 80 ft./sec., 17 although this velocity changes rapidly in many reaction systemsJ
18 ;Isuch as thermal cracking, as the gaseous reaction products are 19 formed. The actual average velocity through the mixing zone 53B
~o l and the reaction zone 44B is a process consideration, the veloclty 21 l being a function of the allowed residence time therethrough.
22 Ij By employing the solid feed device and method of the 23 1I present inventions, the mixing length to diameter ratio necessary 24 ¦¦to intimately mix the two phases is greatly reduced. Thls ratlo I'is used as an informal criteria which defilles good mixing. Gel~-26 l erally, an L/D(length/dia.) ratio of from 10 to 40 is required.
27 ¦IUsing the device disclosed herein, this ratio is less than 5~ with 28 Iratios less than 1.0 being possible. Well designed mixing devi~es 29 11 of the present invention may even achieve essentially complete Imixing at L/D ratios less than 0.5.

I I .
! --20--, ~96-]:~7 1 (c) Improved Sequential Ther~al Crac~ing ___ 2 Process.

4 !l i Turning now to the seguential cracking process 2C

j, of the subiect invention, as illustrated in FIGURE 1~, in lieu of 6 l reactor 28 (see FIGUR~ 1) of the prior art, the system of the , invention includes a solids heater 4C, a primary reactor 6C, a I secondary reactor 8C and dow:nstream equipment, The downstream 9 ,:
' equipment is comprised essentially of an indirect heat exchanger lCCJa 1~ \

ZO

2'7 S&W

,, .
., 1 1 fractionation tower 12C, and a recycle line 14C from the 2 ~ fractionation tower 12C to the entry of the primary reactor 3 , 6C.
4 1 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 from the solids heater 4C. The first or prirnary reactor 6C
in which the first feed is cracked is at high severity conditions 11 Tne hydrocarbon feed, from line 16C, may be any hydrocarhon 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.
i4 Thusr 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 650~F, high boiling point gas oils having a 20 I boiling range between 650 to 950F and compatible combinations 21 of same. These constituents may be introduced as fresh feed 22 or as recycle streams through the line 14C from downstream 23 purification facilities e.g., frac-tionation tower 12C. Dilution steam may also be delivered with the hydrocarbon through lines l 16C an~ l~C. The use of dilution steam reduces the partial 26 l' pressure, improves cracking selectivity and also lessens the 27 ' tendency of high boiling arornatic components to form coke.
28 il The preferred primary feedstock for the high 29 ~ severity reaction is a light hydrocarbon material selected from 3() 1l the group consisting of low molecular weight, hydrocarbon gases, ~ -22-s~w 69~ 7 l I light hydrocarbon liquids, light gas oils boiling between 350 2 and 650F, and combinations of same~ These feedstocks offer the 3 ~ greatest increase improvement in selectivity at high severity 4 ~ and short residence times.
The hydrocarbon feed to the first reac-tion zone is 6 preferably pre-heated to a temperature of between 600 to 1200F
7 before introduction thereto. The inlet pressure in the line 16C
8 is 10 to 100 psig. The feed should be a gas or gasified liquid~
9 The feed increases rapidly in temperature reaching thermal equi-librium with the solids in about 5 milliseconds. As mixing of ll the hydrocarbon with the heated solid occursl 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. T~e residence time maintained within tne primary reactor is about 50 milliseconds~ preferably between 20 and 150 milli-16 seconds~ to ensure a high conversio~ at high selectivity. Typ 17 cally~ the ~SF (Kinetic Severity Function) is about 3.5 ~97~
18 conversion of n-pentane~. Reaction products of this reaction are l9 olefins, primarily ethylene with lesser amounts of propylene and ~ butadiene, hydrogen, methane, C4 hydrocarbons, distillates sucn 21 ; as gasoline and gas oils, heavy fuel oils, coke and an acid qas.
22 j Other products may be present in lesser quentities. Feed con--23 ! version in this first reaction zone is about between 95 to 100r-~, by 24 weight of feed, and the yield of ethylene for liquid feedstocks is about 25 to 45~ ~y weight of the feed, with selectivities o~
26 about 2.5 to 4 pounds of ethylene per pound of rnethane.
27 A second feed is introduced through the line 18C
28 i and combines with the cr~cked ~as from the primary reactor 6C
29 between the primary reactor 6C and the secondary reactor 8C. The ' combined stream comprising the second unreacted feed, and the ~ ,.

__ .

S&W
6~6 147 1 ~ first reaeted feed passes through the secondary reaetor 8C under 2 ~ low severity reaction conditions. The second feed introduced 3 !I through the line 18C is preferably virgin feed stock but may 4 ll also be comprised of the hydrocarbons previously mentioned, including recycle streams containing low molecular weight 6 ; hydrocarbor. gases, light hydrocarbon liquids, low boiling 7 point, light com~atible gas oils, high boiling 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 in most instances the amount of steam initially delivered to the 12 primary reactor 16C will be sufficient to achieve the requisite 13 , partial pressure reduction in the reactors 6C and 8C. It should 14 be understood that the recycle stream 14C is illustrative, and not ' speeific to a particular recycle constituent.
16 ~ ~he hydrocarbon feed delivered through the line 18C
17 ' is preferably virgln gas oil 400-650F. The second feed is pre~
18 !, heated to between 600 to 1200F. and upon entry into the seconlary 19 ' 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 throu~3h the line 18C wiil querlch 22 l~ 60 pounds of effluent from the primary reactor 6C. At this tem~
23 11 perature level, the cracking reactions of the first feed are 24 i essentially terminated. However, coincident with the quenching 'I of the effluent from the primary reactor, the secondary feed 26 1l entering throuc3h line 18C is thermally cracked at this temperature 27 ! (1500 to 1200F) and pressures of 10 to 100 psicJ at low severlt:y 2a I by providing a residence time in the secondary reactor between 29 1l 150 ar,d 2000 milliseconds, preferably between 250 to 500 milli-seeonds. I'ypically, the KSF cracking severity in the secondary , I :
-2~-,; , 69~-147 I'.

1 ' reactor is about 0.5 at 300 to 400 milliseconds.
2 I The inlet pressure of ~he second feed in line 18C lS
3 between 10 and 100 psig, as is the pressure o~ the first feed~
4 Reaction products from the low severity reactlon zone comprise , ethylene with lesse~ 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/
i and the yield of ethylene is about 8 to 20% by weight of feed, 11 ' with selectivities of 2.5 to 4.0 pounds of ethylene per pound-12 of methane.
13 Although the ~roducts from the high severity reaction 14 ; are combined with the second feed, and pass through the second ~5 ! reaction zone, the low severity conditions in the second reac-tion 16 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 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 a-toms associated with such 24 ~' paraffinic side chains will be a large frac~ion of the total , number of carbon atoms in the molecule, or the gas oll will '26 have a low "aromaticity"O
27 ; In the secondary reactor, these molecules will 28 '~ undergo dealkylation - splitting of the paraffin molecules, 29 1 leaving a reactive residual methyl aromatic, which will tend to react -to form hi~h boilers. The paraffins in the boiling I
~25-S&W
696-1~7 , 1 ~ range 400 to 650F are separated from the hiqher boiling 2 1 aromatics in column 12 and constitute the preferred recycle 3 ; to the primary reactor.
4 ' Other recycle feed stocks can include propylener ; butadiene, butenes and the C5 - 900F pyrolysis gasoline.
6 ! The total effluent leaves the secondary reactor 7 and is passed throu~h the indirect quench-means l0C to generate 8 ~ steam for use wi-thin and outside the system. The effluent is 9 then sent to downstream separation facilities 12C via line 24C~
The purification facilities 12C employ conventio-al 11 separation methods used currently in thermal cracking processes~
12 FIGURE 2 illustrates schematically the products obtained. Hydro~
13 ; gen and methane are taken overhead throu~h the line 36C. C4 and 14 lighter olefins, C5 - 400F and 400-650F fractions are removed from the fractionator 12C through lines 26C, 2~C and 30C re-16 spectively. Other light paraffinic gases of ethane and propane 17 are recycled through the line 14C to the high severity primary 18 reactor. The product taken through line 28C consists of liquld 19 ' hydrocarbons boiling between C5 and 400F, and is preferably exported although such material may be recycled to the primary 21 ' reactor 6C if desired. The li~Jht gas oil boiling between 400 22 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 ' exported through stream 32C, while excess residuim~ boiling above 950F is removed from the hattery limits via stream 34C.
26 The heavy gas oil and residuim may also be used as fuel within 27 ` the system.
2~ !' In the preferred embodiment of the process, the 29 ~ second feed would be one which is not recommended for high 1 severity operation. Such a feed would be a gas oil boiling S&W
696-147 Ij 1 ~l above 400F which contains a significant amount of high molecular 2 l weight aromatic components. Generally, these components have 3 paraffinic ,ide chains which will form olefins under proper conditions. However, even at moderate severity, the dealkylated aromatic rings will polymerize to form coke deposits. By pro-6 cessing the aromatic gas oil feed at low severity, it is 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 selectivlty as previously defined is high. Hence~ low severity ~a .
~-' 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 lq I severity section, such compounds being the preferred feeds 15 1I thereto. The aromatic components of the effluent are removed 16 ' from the purification facilities 12C as part of the heavy gas 17 ! oil product, and either recycled for use as fuel within the 18 ; system, or exported for further purification or storage.
19 ~n illustration of ~he benefits of the process of , the invention is set forth below wherein feed cracked and the 21 l 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~
2 4 ~_ _ _ r _ _ 25 1 \
~26 ~ \
27 l S&~7 696~147 1 (d) Improved ~esidence Time Solid-Gas Separation 2 Device and System.

Referring to FIGURE 15 in the subject invention, in lieu ; separation zone or curved segment region 36 and the quer.ch area 6 44 of the prior art TRC system (see FIGURE 1~, solids and gas enter 7 the tubular reactor 13D through lines llD and 12D respectively~
8 The reactor effluent flows directly to separator l~D where a 9 separation into a gas phase and a solids phase stream is effected.
The gas phase is removed via line 15D, while the solid phase is 11 sent to the 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 ma~erial should 17 Ibe 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 aas is removed overhead through line l9D. Solids 21 from lines 16D and 21~ 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 4 the stripping vessel through line 24D and, if economical or if ~ need be, sent: to down-stream purification units. Stripped solids -'~ Ij S&W
6~6 147 1 j I I . ;
!
jl removed from the vessel 22D through line 25D are sen-t lto re 2 I generation vessel 27D usi.ng pneumatic transport gas from line 3 ll26D. Off gases are removed from the regenerator throuyh line 28D.
4 1 After regeneration the solids are then recycled to xeactor 13D
via line llD~
6 The separator 14D should disengage solids rapidly 7 from the reactor effluent in order to prevent product degradation 8 i and ensure optimal yield ancl 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 I stripping vessel 22D inasmuch as this portion of the gas product ;
12 would be severely degraded by remaining in intimate contact wlth 13 the solid phase. This is accomplished with a positive seal whlch 14 has been provided between the separator 14D and the stripping lS vessel 22D. Finally, the separator 14D operates so that 16 erosion is minimized despite high temperature and high velocity 17 conditions that are inherent in many of these processesO The 18 I separator system of the present invention is designed to meet 19 each one of these criteria as is described below.
FIGUR~16 is a cross sectional elevational view 21 '! showing the preferred embodiment of solids-gas separation de~ice 22 14Dof the present lnvention. The separator 14D is provided 23 with a seyarator shell 37D and is comprised of a solids-gas 24 , disengaging chamber 311) having an inlet 32D for the mixed phase;
25 I stream, a gas phase outlet 33D, and a solids phase outlet 34D.
~26 , The inle~ 32D and the solids outlet 34D are preferably locatecl 27 ' at opposite ends of the chamber 31D. While the gas outlet 3:1D
28 lies at a point therebetween. Clean-out and maintenance manw~ys 29 35D and 36D may be provided at either end of the chamber 31D.

The separator shell 37D and manways 35D and 36D preferably are Il , S &W
6~6-1~7 11' .

~ ned W~ Lh erosion resistent linings 38D, 39D and 41D re-2 spectively which may be required if solids at high velocities 3 j are encountered. Typical commercially available materials 4 l for erosion resistent lining include Carborundum Precast ~ Carbofrax D, Carborundu~ Precast Alfrax 201 or their equivalentO
6 A thermal insulation lining 40D may be placed between shell 37D
7 ~ and lining 38D and between the manways and their respective 8 1 erosion resistent linings when the separator is to be used g 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 construction 14 the separator 14D shell i5 preferably fabricated from cylindrlcal sections such as pipe 50D, although other materials may, of 16 course, be used. It is essential that longitudinal side walls 17 51D and 52D should be rectilinear, or slightly arcuate as in-18 I dicated by the dotted lines 51D and 52D. Thus, flow path 31D
19 through the separator is essentially rectangular in cross section having a heisht H and width W as shown in FIGURE 17.
21 The embodiment shown in FIGUXE 17 defines the geometry of the 22 flow path by adjustment of the lining width for walls 51D and 23 j 52D. Alternatively, baffles, inserts, weirs or other means 24 ,I may be used. Xn like fashion the configuration of walls 53D
' and S4D transverse to the flow path may be similarly shaped, 26 although this is not essential. FIGURE 18is a cutaway view 27 l along Section 9~4 of FIGURE16 wherein the separation shell 37D
2R !~ is fabricated from a rectangular conduit. Because the shell 37D
l!
29 has rectilinear walls 51D and 52D it ls not necessary to adjust 1 the width of the flow path with a thickness of lining. Linings ! -30-~ , 4~

38D and 40D could be added for erosion and thermal resistence respectively.
Again reEerring to Figure 1~ inlet 32D and ou-tlets 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 charnber. As a further requirement, however, the gas phase outlet 33D is also oriented so that the gas phase upon leavlng the separator has completed a 180 change in direction.
Centrifugal force propels the sol.id par-ticles to the wall 54D opposite inlet 32D oE the chamber 31D, while the gas portion, having less momentum, flows through the vapor space of the charnber 31D. Initially, solids impinye on the wall 54D, but subsequen-tly accumulate to form a static bed of solids 42D, which ultima~ely 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 i5 preferably oriented for downflow of solids by gravi-ty. The exact shape of the arc 43D is determined by the geometry of the par-ticular separator and the inlet stream parameters such as velocity, mass flowrate, bulk densi-ty, and particle size. Because the force imparted to the incoming solids is directed against the static bed 42D rather than the separator 14D itself, erosion is minima].. Separator efficiency, defined as the removal of solids from the gas phase leaving -through outlet 33D, is, therefore, no-t affec-ted adversely by high inlet velocities up to 150 ft~/sec., and -the separator 14D
ls operable over a wide range of dilute phase densi-ties, preferably between 0.1 and 10.0 lbs./f-t3. The separa-tor 14D of -the present invention achieves efficiencies of about 80~, al-though the preferred embodi.ment, discussed below, can obtain over 90~ rernoval of solids.

It has been found -~ha-t 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 U~bend be-tween inlet 32D and outlet 33D becomes progressively sharper; tha-t is, as outlet 33D is brought progressively closer to inlet 32D. Thus, for a given H the efficiency of the separator increases as the flow path and~ hence, residence time decreases. Assuming an inside diameter Di of in]e-t 32D, the preferred distance CL between the centerlines of inlet 32D and outlet 33D is less than 4.0Di, while the most preferred distance between said centerlines is between 1.5 and 2.5 Di. Below 1.5 Di better separation is ob-tained 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 a-t 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 ~ inches the incoming stream is apt to dis-turb the bed solids ~2D, 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 o-therwise limited, it is apparen-t that too large an H
even-tually merely increases resi~ence time without substantive increases in ef~iciency~ The wid-th W of -the flow path is ! - 32 -'I !
s~w , 696-147 ~l ' 1preferably between 0.75 and 1.25 times Di 9 most preferably between 2 ~ o.9 and 1.10Di a 3Outlet 33D may be of any inside diameter. However, 4 , velocitles 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 l; between the stripping vessel 22D shown in FIGURE1s 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 fcrms a positive seal which prevents gases from 11 1 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 ehight of solids in line 16D. As the differential increases 16 , the net flow of gas to the stripping vessel 22D decreases.
17 Solids, having ~xavitational momentu~, overcome the differentlal 18 while gas preferentially leaves through the gas outlet 33D.
19 ~' By regulating the pressure in the stripping vessel 20 ' 22D it is possible to control the amount of ~as going to the 21 , stripper. The pressure regulating means may include a check 22 l, or "flapper" valve 29D at the outlet of line 16D, or a pressure 23 I control 29D device on vessel 22D o Alternatively, as suggested 24 l~ above, the pressure may be re~ulated by selecting the size of 2~1 ! the outlet 34D and conduit 16D to obtain hydraulic forces ~6 ~! acting on the system that set the flow of gas to the stripper 27 I' 32D. ~Jhile such gas is degraded, we have found that an increase 2~3 '' in separation efficiency occurs with a bleed of gas to the 29 ' stripper of less than 10~, preferably between 2 and 7~. Economic , and process considerations would dictate whether this mode of 1 ~

, . ", ,. -S&W / i ~;g6-147 .1 . I

1 l operation should be used. It is also possible to design the 2 ~ system to obtain a net backflow of gas from the stripping 3 ~I vessel. This ~as flow should be less than 10~ of the total 4 I feed gas rate.
By establishing a minimal flow path, consistent 6 with the above recornmendations, residences times as low as 7 O.l seconds or less may be obtained, even in separators 8 ~I having inlets over 3 feet in diameter~ Scale-up to 6 fee-t 9 , in diameter is possible in many systems where residence times approaching 0.5 seconds are allowable.
11 In the preferred embodiment of FIGURE 16, a weir 44D
12 ' is placed across the flow path at a point at or just beyond the 13 gas outlet to establish a positive hei~ht of solids prior to 14 1; 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 ~Jas towards the gas outlet and creates the U~shaped I gas glow pattern illustrated diagrammatically by line 45D ir 21 ' FIGURE 16. Without the weir 44D an arc sornewhat less than or 22 equal to 90 would be formed, and which would extend asymptoti-23 , cally toward outlet 34D as shown by dotted line 60D in the 24 'I schematic diagrarn of the separator of FIGURE 19~ While neither efficiency nor gas loss (to the stripping vessel) is affected 26 11 adversely, the flow pattern of line 61D increases residence t1me, 27 and more importantly, creates greater potential for erosion at 28 ~l areas 52D, 63D and 64D.
29 'rhe separator of FIGURE 20 is a schematic diasram of 30 1 another ernbodiment of the separator 14D, said separator 14D
l l S &W ! ~ L r3~
69~-147 ll ~

Il 1 1 having an extended separation chamber in the lon~itudinal 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 I bed of greater length. L is preferably less than or equal !i to 5 Dio Although the gas flow pattern 61D does not develope 6 ,! the preferred U-shape, a crescent, shaped arc is obtained 7 ! which limits erosion potential to area 64D. Embodiments 8 ,I shown by FIGURES 19 and 20are useful when the solids loading 9 , of the incoming stream is low. The er~odiment 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 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 to the preferred embodiment of FIGUR~ 16. Because solids accumu-17 1, 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 obtained 22 ' and gas losses are rninimal.
23 , FIGURE 22 illustrates another embodiment of the 24 separator 14D of FIGURE 21wherein the solids outlet is steppea.
l Althou~h a weir is not used, the outlet restric~s solids flow 26 , which helps from the bed 42D. As in FIGURE 20, an extended L
27 , distance between the gas outlet and solids outlet may be used~
28 The separator of FIGURE21 or22 may be used in 29 I conjunction with a venturi, an orifice, or an equivalent flow restriction device as shown in FIGURE 23. The venturi 69D hav~ng ' ' .

.. . . . . _ _ _ S&W
696-147 j, l 'I 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 ,, 66D of the outlet 65D to greatly improve deaeratlon of solids~
The embodiment of FIGURE ~4 is a variation of the separator 6 shown inFIG~RE 23. Here, inlet 32 and outlet 33D are oriented for use in a riser type react:or. Solids are propelled to the 8 wall 71D and the bed thus formed is kept in place by the force g of the incoming stream. As before the gas portion of the feed follows the U-shaped pattern of line 45D. However, an asymptotic ll bed will be formed unless there is a restriction in the solids 12 outlet. A weir would be ineffective in establishing bed heign~7 13 and would deflect solids into the gas outlet. For this reason 14 I the solids outlet of FIG~E 23 is preferred. ~lost preferably, ; the venturi 69D is placed in collinear section 66D as shown in l6 FIGU~E24 to improve the deaeration of the solidsO Of course, 17 each of these alternate embodiment may have one or more of the l8 optional design features of the basic separator discussed in 19 , relation to FIGURES 16, 17 and 18.
rihe separator of the present invention is more 21 clearly illustrated and explained by the examples which Eollow.
22 ' In these examples, which are basedon data obtained during 23 experimental testing of the separator design, the separator 24 j has critical dimensions specified in I~able I. These dimenslons ~ (in inches except as noted) are indicated in the various drawlny ~26 figures and listed in the Nomenclature belowo Distance between inlet and gas outlet centerlines ~ D Insiae diameter of inlet 29 'I Do~ Inside diameter of gas outlet 30 1 Dos Inside diameter of solids outlet v Diameter of venturi inlet DVt Diameter of venturi t~roa~

H Height of flow path HW Height of weir or step L Length from gas outlet to weir or step as indicated W Width of flow path O Angle of cone formed by projection of convergent venturi walls, degrees TABLE I

Dimensions of Separators in Examples 1 to 10, inches*

Example Dimension 1 2 3 4 5 6 7 8 9 10 CL 3.875 3.875 3.875 3.875 3.875 3.875 11 11 3.5 3.5 Di 2 2 2 2 2 2 6 6 2 2 Dog 1.75 1.75 1.75 1.75 1.75 1.75 4 4 1 1 Dos 2 2 2 2 2 2 6 6 2 2 Dv - - - - - - - - - 2 Dvt - - - - - - - - - 1 H 4 4 4 4 4 4 12 12 7.5 6 Hw 0.75 0.75 0.75 0.75 0.75 0.75 2.25 2.25 0 4 0, degrees - - - - - - - - - 28 *Except as noted Example 1 In this example a separator of the preferred embodiment of FIGURE 16 was tested on a feed mixture of air and silica alumina. The dimensions of the apparatus are specified in Table I. Note that the distance L from the gas outlet to the weir was zero.

S&~
696 1~7 i 1 ~ The inlet stream was comprised of 85 ft.3/min.
2 'l of air and 52 lbs./min~ of silica alumina having a bulk density 3 ~, of 70 lbs./ft3 and an average particle size of 100 microns.
4 I The stream density was 0~612 lbsO/ft~3 and the operation was ~ performed at ambient temperature and atmosphexic pressureO
6 1 The velocity o the incoming stream through the 2 inch inlet 7 ' was 65.5 ft./sec., while the outlet gas velocity was 85~6 ft./secO
8 I~ through a 1.75 inch diameter outlet. A positive seal of solids 9 ~' in the solids outlet prevented gas from being entrained in the ; solids leaving the separator. Bed solids were stabilized by 11 placing a 0.75 inch weir across the flow path~
12 I The observed separation efLiciency was B9.1~, 13 ' and was accomplished in a gas phase residence time of approximately 1~ ll 0.008 seconds. Efficiency is defined as the percent removal of ll solids from the inlet stream.
16 'I Example 2 17 The gas-solids mixture of Example 1 was processed 18 , in a separator having a configuration illustra~ed by FIG~RE 2Q.
19 l In the example the L dimenslon is 2 inches; all other dimensions ' are the same as Example 1. By extendiny the separation chamber 21 along its longitudinal dimension, the flow pattern of the gas 22 ' began to deviate from the U-shaped discussed aboveO As a resu t 23 l~ residence time was longer and turbulence was increasedO Separ~tion 24 1l efficiency for this example was 70.8%.
~5 1l Example 3 26 I The separator of Example 2 was tested with an inlet 27 ~I stream comprised of 85 ft~3/min. of air and 102 lbs./min~ of 28 I silica ~lumina which gave a stream density of l.lB lbs./ft.3 29 1 or approximately twice that of Example 2. Separation efficiency , improved to 83.8%.

-3~-, . . . _ . ~ _ . . . . . .

Example 4 The preEerred separator of Example 1 was tested at the inlet flow rate of Example 3. Efficiency increased slightly to 91.3~.
Exarnple 5 The separator of FIGURE 16 was tested a-t the conditions oE Example 1. Although the separation dimensions are specified in Table I note that -the distance CL between inle-t and gas outlet centerlines was 5.875 inches, or about -three times -the diameter of the inlet. This dimension is outside the most preferred range for CL which is between 1.50 and 2.50 Di.
Residence time increased to 0.01 seconds, while efficiency was 73.0%.
Example 6 Same conditions apply as for Example 5 except that the solids loading was increased to 102 lbs./min. to give a stream density of 1.18 lbs./ft.3. ~s observed previously in Examples 3 and 4, the separator efficiency increased with higher solids loading to 90.6~.
Example 7 The preEerred separa-tor configuration of FIGU~E 16 was tested in this Example. However, in this example the apparatus was increased in size over the previous examples by a factor of nine based on flow area. A 6 inch inlet and 4 inch outlet were used to process 472 ft.3/min. o air and 661 lbs./min of silica alumina at 180~F. and 12 psig. The respective velocities were 40 and 90 ft./sec" The sollds had a bulk density of 70 lbs./ft3 and the stream density was 1.37 lbs./f-t.3 Distance CL between inle-t and gas out:Let cen-terlines was 11 inches, or 1.83 times the inlet diameter; distance L was ~ero. I'he bed was , S&W '!

1 stabili3ed by a 2.25 inch weir, and gas loss was prevented ? ,; by a positive seal of solidsO 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/rnin. Observed separator efficiency 7 was 90. 06, 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 I 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 outlet at a veloci~y of 3.6 , ftv/sec. Separator efficienc~ 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 I having a 2 inch inlet and a 1 inch gas outlet. The solids out--let was 2 inches in diameter and was located 10 inches away 21 from tne 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 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 f~./secO As 27 ln Example 7 there was a positive counter current flow of 28 ' displaced gas from the collection vessel to the separator.
29 I This flow was approximately 1.7 ft.3/ min~ at a velocity of 3~ 1.3 ft./sec. Operation was at amhient temperature and atmos-31 pheric pressure. Separator efficiency was 95O0%~

' S&W , 696-147 ll 1 IExample 10 The separator of FIGU~E23 was tested on a feed 3 comprised of 85 ftv3/ min. of air and 78 lbs./min, of spent 4 ,l Fluld 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 o 262 ft ./sec. This separator had a stepped 8 solids outlet with a venturi in the colllnear section of ~he 9 outlet. The venturi mouth was 2 inches in diameter, while 10 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 gas 14 remaining in the solids' voids~

17 ` \

18 ~ \

29 , \
,;

~IJ

1 (e) Improved Solids Quench Boiler and 2 Process.

S As see~ in FIGURE 25,in lieu of quench 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 se~arator 8E
9 with a solids return line lOE to the fluid bed-riser 6E and a line 1~ 36E for the delivery of gas to the fluid bed-quench riser.
11 The quench exchanger 4E as best seen in FIGURES 26 ~nd ~7 12 is for~ed 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 h~ .
6~-147 quench exchanger 4E. The tubes 16E are joined together, pre-2 I ferably by welding, and form a pressure-tight ~em~rane wall ~hich 3 is in effect, the outer wall o the quench exchanger 4E. The 4 inner circles of tubes 18E ancl 20E: are spaced apart and allow S for the passage of effluent gas and particulate solids there-6 around. The arrays of tubes 16E, 12E and 20E are manifolded 7 to an inlet torus 24Eto 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 ou-tlet hood 23E 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 delive~ the 16 cooled effluent and particulate solids to -the cyclone separatcr L7 ' 8E.
18 The fluid bed quench riser 6E is essentially a sealed 19 vessel attached in sealed relationship to the quench exchanger 4E. The fluid bed-quench riser 6E is arranged to receive the 21 reactor outlet tube 36E which is preferably centrally disposed at 2Z the bottom of the fluid quench riser 6E. A slightly enlarged 23 centrally disposed tube 38E is aligned with the reactor outlet.
24 ;36E and extends from the fluid bed-quench riser 6E into the quench exchanger 4E. In the quench exchanger 4Ef the centrally '26 disposed fluid bed-quench riser -tube 38E terminates in a conical 27 opening 40E`. The conical opening 40E is provided to facilitate 28 ; nont:urbulent transition from the quench riser tube 38E to the 29 enlarged opening of the quench exchanger 4E. I~ has been found that the angle of the cone ~, best seen in EIGURE26~ should 31 be not greater than 10 de~rees.
; -~3--.,._ S&W~

. I .

1 ¦I The fluid bed 42E contained in the fluid bed quench 2 1! riser 4~ is maintained at a level well above the bottom of the 3 l, quench riser tube 3~E. A bleed line 50E is provided to bleed 4 ~I solids from the bed 42E. Although virtually any particulate '~ solids can be used to provide the ~uench bed 42E, it has been 6 ~I found in practice that the same solids used in the reactor are 7 ' preferably used in the fluidlzed bed 42E. Illustra-tions of 8 `1 the suitable particulate solids are FCC alumina solidsO
9 , As best seen in FIGURE 28,the opening 48E through ~I which the fluidized particles from the bed 42E are dra~n into 11 . ~he 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 surface of the reactor outlet 14 ~ tube 36E. In practice~ it has been found that the refractory I cone 46E can be formed of any refractory materialO The openiny 16 ' 48, defined by the conical end g4E of the quench riser tube 38E
17 . and the refractory cone 46E, is preferably 3-4 square feet for 18 ~ a unit of 50 ~MsTu/HR capacity. The opening is sized to insure 13 ~ penetration of the cracked gas solid mass velocity of 100 to ~ 800 pounds per second per square foot is required~ The amount 21 l, of solids from bed 42~ delivered to the tube 38E is a functior 22 l' of the velocity of the gas and solids entering the tube 38E
2 3 1l from the reactor outlet 36E and the size of the opening 48E.
24 l In practice, it has been found that the Thermal 1 Regenerative Cracking tTRC) reactor effluent will contain 26 ,~ approximately 2 pounds of solids per pound of gas at a tem-27 1I perature of about 1,400F to 1,600F.
28 l The process of the solids quench boiler 2E of 29 FIGUR~S 25-2E is illustrated by the following example. Effluent 1 from a TRC outlet 36E at about 1,500DF is delivered to the quench , . .

., .

s~w ll 6~6-147 I, .
1 ¦I riser tube 38E at a velocity of approximately 40 to 100 feet 2 l, per second. The ratio of particulate solids to cracked effluent 3 1 entering or leavlng the tube 36E i5 approximately two pounds of 4 ', solid per pound of ~as at a temperature of about 1,500E. At , 70 to 100 feet per second the particulate solids entralned into 6 1, the effluent stream by the eductor effect is hetween twenty flve 7 and fifty pounds solid pex pound of gas. In 5 milliseconds the 8 ' addition of the particulate solids from the bed 42E which is 9 at a temperature of 1,000F reduces the temperature of the ~ composite effluent and solids to 1,030F. The gas-solids mixture 11 , is passed from the quench riser tube 38E to the quench exchanger 12 4E wherein the temperature is reduced from 1,030F to 1,000F
13 by indirect heat exchange with the boiler feed water in the tuDes 14 ' 16E, 18E, and 20E. With 120,000 pounds of effluent per hour, 50 MMsTus per hour of steam at 1,500 PSIG and 600F will be 16 . generated for system service. The.pressure drop of the gas 17 , solld mixture passing through quench exchanger 4E is 1.5 PSI. The 18 l, cooled gas-solids mixture is delivered through line 30E to the 19 , cyclone separator 8E wherein the bulk of the solids is re~oved from the quenched-cracked gas and returned through line lOE
21 ,I to the quench riser 6E.

24 , '~6 ~7 1, ~
~8 ,j \
29 ' \
l, , -45-.

S&W

1 lf) Improved Preheat V~orization System.

3 ~gain referring to FIo~29,in lieu of pre~eat zone 24 (FIGURE 1 4 of the system 2F'of the subject invention is em~odied in a TRC system and is co~prised of essentially a liquid Eeed 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 12FI 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 desulfurized Kuwait ~GO to 12 initially e]evate the temperature of the feedstock.
13 The initial mixer 8F is used in the system 2F to initially flash superheated steam from a steam line 6F and the heated feedstock delivered from the liquid feed heater 4F by 16 a line 18F.
17 The system separator lOF is to separate the liquid and 18 vapor produced by flashing in the mixer 8F. Separated gas is 2~ \

~46-S&~

l lldischarged through a line 22F from the separator overhead and 2 ~ the remaining 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 '. The second mixer 14F is provided to flash the vapori~ed 7 ; gaseous discharge from the vapor feed superheater 12F and the 8 , liquid bottoms from the separator 10F, thereby vaPOrizing the 9 . composite steam and feed initially delivered to the system 2F.
~ knockout drum 16F is employed to remove any liquid ll from the flashed vapor discharged from the second mixer 14F
12 through the line 28F. The liquid-free vapor is delivered to a 13 reactor through the line 30F.
14 , In the subject process, the heavy oil liquid hydro-,carbon feedstock is first heated in the liquid feed heater 4F
16 to a tem~erature of about 440 to 700F. The heated heavy 17 oil hydrocarbon feedstock is then delivered through the line 13 18F to the mixer 8F. Superheated 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 21 700 to 800F. For lighter feedstock the flashing temperature 22 will be about 500 to 600F., and for heavler feedstock the 23 flashing tempera-ture ~ill be about 700 to 900F.
24 1l The flashed mixture of the steam and hydrocarbon is 25 ~ sent to the system separator lOF wherein the vapor or gas is 26 ,taken overhead throush the line 22F and the liquid is 27 I discharged t.hrough the line 26F. Both the overhead vapor and 28 liquid bottoms are in the temperature range o about 700 to 29 ~ 800F. The temperature level and percent of hydrocarbon I vaporized are determined within the limits of equipment fouliny ., ~, -47-"

1 criteria= The vapor stream in the line 22F is comprised of 2 essentially all of the steam delivered to the system 2F and 3 a large portion o~ the heavy oil hydrocarbon feedstock.
4 Between 30% and 70% of Lhe heavy oil hydrocarbon feedstock supplied to the system will be contained in the overhead 6 leavin~ the separator 10F through the line 22F.
7 The steam-hydrocarbon vapor ln the line 22F is delivered 8 to the system vapor feed superheater 12F wherein it is heated ~o 9 about 1,030F. The heated vapor is taken from the vapor Eeed ~superheater 12F through the line 24F and sent to the second mixer 11 14F~ Liquid botto~s from the separator 10F is also delivered 12 to the second mixer 14F and the vapor-liquid mix is flashed in 13 the mixer 14F to a temperature of about 1,000F.
14 The flashed vapor is then sent downstream through the ' line 28F to the knockou~ drum 16F for removal of any liquid 16 from the vapor. Finally, the vaporized hydrocarbon feed is 17 ,, sent through the line 30F to a reactor.
18 An illustration of the system preheat process is 19 ; seen in the following exampleO
20 A Nigerian Heavy Gas Oil is preheated and vaporized n 21 the system 2F prior to delivery to a reactor. The Niserian Hea~y 22 Gas Oil has the following composition and propertjes.
23 ~l 24 Flemental Analysis, Wt.~ _operties Carbon 86.69 Flash Point, F. 230.0 ~lydrogen 12.69 Viscosityl SUS 210 F 44.2 26 Sulfur .10 Pour Point, F ~90.0 Nitrogen .047 Carbon Residue, ~bot~Qm .09 27 I Nickel .10 Aniline Point, C87.0 Vanadium ol0 29 I ;

.

696-147 ~ 5~
"

" .
1 Distillation i 2 I:Vol. %
!
3 .IBP
1 10669.2 4 30755.6 50820.4 ' 70874.4 90g44.6 6 EP1,005.8 8 3,108 pounds per hour of the Nigerian ~leavy Gas Oil is 9 heated to 750E'. in the liquid feed heater 4F and delivered at a pressure of lS0 psia to the mixer 8F. 622 pounds per hour of 11 superheated steam at 1,100F. is simultaneously delivered to the 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 Heav~
, Gas Oil is vaporized.
16 . The vapor and liquid from the mixer 8F are separated 17 ~ in the separator 10F. 622 pounds per hour of steam and 1,864.8 18 ~ pounds per hour of hydrocarbon are taken in line 22~ as overhead 19 ; vapor. 1j243.2 pounds per hour of hydrocarbon are discharsed through the line 26F as liquid and sent to the mixer 14E~.
21 The mixture of 622 pounds per hour of steam and 22 1,864.8 pounds per hour of hydrooarbon 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 main-tained at 45 psia.
, 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 1,86fi.8 pound per hour of hydrocarbon are flashed in the mixer 23 . 14F to 990~F.
29 The vaporization of the hydrocarbon is effected with a ,Is-team to hydrocarbon ratio of 0.2. The heat necessary to vaporize ~&W
696 1~7 i "
1 ~ the hydrocarbon and generate the necessary steam is 2.924 MM
2 ,BTU/hr.
3 ~I The same 3,108 pouncls 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 BTU/hr. In order to 8 . reduce the input energy in the conventional process to the same 9 level as the present invention, a steam temperature of 3,208F~
,;is required, which temperature is effectively beyond design limitations.

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 l reaction must be. quenched rapidly. When solids are used, they 18 l 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 soli.ds will be rejected to some lower 24 ,llevel heat recovery. Thus, a rapid high efficiency separator, !laccording 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), lt is noted that in order to 28 feed solids to an ethylene reactor, the flow must be controlled 29 Il to within +2 percent or cracking severity oscillations will be 1I greater than that presently experlenced in coil cracking. The S&~

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 contri~ute 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 timej. 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 reacti~n 12 occurs if the solids are not contacting the gas (poor mixing).
13 In connection with the solids quench boilex 14 (see FIGURES 25-28), in the current TRC concept, a 90 percent separation occurs in the primary separator. This is followed 16 by an oil quench to 1300, and a cyclone to remove the remainder 17 of the solids. The mix is then quenched again with liquid to 18 600F. Thus~ all the available heat from the reaction outlet 19 temperature to 600F is rejected to a circulating oll 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 2~ 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 psLg~
28 ~lhich increased the overall thermal efficiency of the process.
2~ The oil loop can not operate at these temperatures due to ins-tabilities (too many light fractions are boiled off, yieldin ~!. ' 69~-147 1 an oil that is too viscous). The use of 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. ~uring ~uenchinq, 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 T~C it could use the regenexator ~hat alread~ e~ists.
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 maxirum economic advanta~es when crackin~ heavy feedstocks (650F+ boiling point) a-t low steam dilutions.
11 Selectivity is favored by rapid and intimate mi~ing. Rapid 12 and intirr.ate mixincJ is best accorrplished with a vapor feed 13 rather than a lir~uid feed.
1~ Finally, with reference to the sequential cracking 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 a be applied to both TRC and a coil. Its synergism with T~C
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 perrlits the use of loweL
23 temperatures compared to the coil case.
2~ ~Jhile there has been described what is considered to ~e preferred embodiments of the invention, variations and modif-26 ications therein will occur to those skilled in the art once -they become acquainted with the b~sic concepts of tne invention.
28 Therefore, it is intended that the appended claims shall be 29 constxued to include not only the disclosecl embodiments but all such variations and rnodifications that f~ll within the true 31 spirit and scope of the invention.

Claims (22)

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

1. In a TRC process wherein the temperature in the cracking zone is between 1300° and 2500°F and wherein the hydrocarbon feed or the hydrodesulfurization residual oil along with the entrained inert solids and the diluent gas are passed through the cracking zone for a residence time of 0.05 to 2 seconds, the improvement comprising the process for generating fuel oil and removing coke deposits on said solids comprising:

a. generating fuel gas having a high CO to CO2 ratio and a high molal ratio of H2O to H2 from fuel, air and steam;
b. delivering the fuel gas to a tubular transfer line;
c. delivering particulate solids having coke deposits thereon to the tubular transfer line;
d. mixing the fuel gas and the particulate solids to reach an equilibrium temperature; and e. removing coke from the particulate solids with heat and steam from the particulate solids-fuel gas mixture.

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

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

4. In a TRC process wherein the temperature in the cracking zone is between 1300° and 2500°F and wherein the hydrocarbon feed or the hydrodesulfurization residual oil along with the entrained inert solids and the diluent gas are passed through the cracking zone for a residence time of 0.05 to 2 seconds, the improvement comprising the process for generating fuel oil and removing coke deposits on said solids comprising:

a. generating a fuel gas from fuel and air;
b. delivering the fuel gas to a transfer line;
c. mixing the particulate solids with the fuel gas in the transfer line to elevate the temperature of the solids; and d. combusting the fuel gas in the transfer line to elevate the temperature of the solids and remove to coke from the solids.

5. A process as in Claim 4 further comprising the steps of pre-heating fuel to 600°F and air to 1,000°F; combusting 7,000 pounds per hour of the pre-heated fuel and 13 MM SCFD
air to generate the fuel gas for delivery to the transfer line and wherein 13 MM SCFD of air is delivered to the transfer line to facilitate combustion of the fuel gas in the transfer line.

6. In a TRC system wherein the temperature in the cracking zone is between 1300° and 2500°F. and wherein the hydrocarbon feed or the hydrodesulfurization 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 comprising a system for heating and removing coke from the particulate solids comprising:

a. means for generating fuel gas having a high molal ratio of H20 to H2 from fuel, air and steam;
b. a transfer line; and c. means to mix the fuel gas and particulate solids in the transfer line, whereby the fuel gas elevates the temperature of the particulate solids by intimate contact therewith and the steam in the fuel gas removes the coke from the solids during travel through said transfer line.

7. A system as in Claim 6 further comprising means to separate the fuel gas from the heated particulate solids after the solids have been heated and cleaned of coke.

8. A process for generating fuel gas and removing coke deposits on solids comprising:

a. generating fuel gas having a high CO to CO2 ratio and a high molal ratio of H2O to H2 from fuel, air and steam;
b. delivering the fuel gas to a tubular transfer line;
c. delivering particulate solids having coke deposits thereon to the tubular transfer line;
d. mixing the fuel gas and the particulate solids to reach an equilibrium temperature; and e. removing coke from the particulate solids with heat and steam from the particulate solids-fuel gas mixture.

9. A process as in Claim 8 wherein the fuel gas is generated by combusting to 2,300°F 6,500 pounds per hour of fuel pre-heated to 600°F., 114 MM SCFD of air pre-heated to 1,000°F. and 72,400 pounds per hour of steam pre-heated to 1,000°F.

10. A process as in Claim 8 wherein the particulate solids mixed with the fuel gas are at a temperature of about 1,450°F.
and the equilibrium temperature of the particulate solids-fuel gas mixture is 1,780°F.

11. A process as in Claim 10 wherein the particulate solids are decoked by the passage with the fuel gas in the vessel at about 100 feet per second and the steam decoking reaction reduces the particulate solids-fuel gas temperature from 1,780°F. to 1,750°F.

12. A process as in Claim 9 wherein the resultant fuel gas has a BTU value of about 100 BTU per SCF.

13. A process as in Claim 10 further comprising a residence time of about 5.0 millisecond for the particulate solid-fuel gas mixture to reach the equilibrium temperature and a residence time of about 1 second to steam decoke the particulate solids.

14. A process as in Claim 9 wherein the fuel gas has a composition of 8.9 mol per cent H2, 18.2 mol per CO, 9.5 mol per cent CO2, 11.7 mol per cent H2O and 51.7 mol per cent N2.

15. A process as in Claim 8 further comprising the step of combusting the fuel gas in the transfer line to further heat the solids and remove the coke from the solids.

16. A process for heating particulate solids and removing coke from the solids before delivering the solids to a thermal reaction chamber comprising:

a. generating a fuel gas from fuel and air;
b. delivering the fuel gas to a transfer line;
c. mixing the particulate solids with the fuel gas in the transfer line to elevate the temperature of the solids; and d. combusting the fuel gas in the transfer line to elevate the temperature of the solids and remove the coke from the solids.

17. A process as in Claim 16 wherein the fuel gas delivered to the transfer line is at a temperature of 2,300°F., the particulate solids mixed with the fuel gas are at a temperature of 1,450°F., the equilibrium temperature of the fuel gas-solids mixture is about 1,486°F. and the combustion of fuel gas in the transfer line elevates the temperature of the solids to 1,750°F.

18. A process as in Claim 16 further comprising the steps of pre-heating fuel to 600°F. and air to 1,000°F; combusting 7,000 pounds per hour of the pre-heated fuel and 13 MM SCFD air to generate the fuel gas for delivery to the transfer line and wherein 13 MM SCFD of air is delivered to the transfer line to facilitate combustion of the fuel gas in the transfer line.

19. A system for heating and removing coke from particulate solids comprising:
a. means for generating fuel gas having a high molal ratio of H2O to H2 from fuel, air and steam;
b. a transfer line; and c. means to mix the fuel gas and particulate solids in the transfer line whereby the fuel gas elevates the temperature of the solids by intimate contact therewith and the steam in the fuel gas removes the coke from the solids during travel through the transfer line.

20. A system as in Claim 19 wherein the transfer line is about 100 feet long.

21. A system as in Claim 19 further comprising means to separate the fuel gas from the heated particulate solids after the solids have been heated and cleaned of coke.

22. A system as in Claim 20 further comprising means to introduce air into the transfer line.