Hydrocarbon gas processing

200~9~7 N~N G~S ~SI~* S~ FIC~IION E~CK~R~UND OF IHE INVENIIO~ This invention rela~e to a process for the separation of a gas containing hydrlcarbors.

Ethylene, ethane, prcpylene, propane and heavier hydr=c~rbcns can be recovered from a variety of gases, such as nstural gas, refincry gas, and synthetic gas streams obtained ~rom okher hyd~ro~rton materials such as rn~l~ crude oil, naphtha, oil shale, tar sands, and lignite. NatNral ~a~ ly has a ma~or prcportion of methane and ethane, i.e. methane and ethane together comprise at least 50 mole percent of the gas. The gas may also co~tain relatively lesser amounts of heavier hydroc~rbons ~uch as prcQane, bu~ares, pentanes, and the like as well as hydrogen, nitrcgen, carbon dioxide and okher gases.

The pre~ent invention i5 generally cnnoerned with the reoovery of ethylene, eth3ne, proQylene, prcpane and heavier hyor~c=rtons from such gas streams. A typical analysis of a ~a~ s*ream to be processed in accordance with this invention wculd be, in a~proximate m~le p#roent, 92.5% ~ethane, 4.2% ethane and cther C2 oooponects, 1.3% propane and other C3 components, ~ 2003 1 q7 0.4% ~re, 0.39~ no~al blt~r#~ 0.5% pentan~ plU5, W$t~1 t~be bal~noe ~hic~ ~n pr~vi~e ~re &~fficient r~ies of ~ pro~ . avail~ble ~nd refrigerati~ of the gz~c, oil ~ , ~ refrigera~d oil ~t~sly e~rdir~ ~ ~r~ct~ng l~t fr~ t~ gas be~ pn~.

I~iirq upon the pr~ of ff~ gas oc~s, the ri~s tethar~ u~ lS of these p~ ticn ~eof ~Day be ~q?lc~d.

e~Y~e n~x~ery b~use ~t prov~des ~aK~um 8inplicity wdth eas~ of db~t up, q~ting flQ~b~l~ty, gpod fficdency, 6afety, an~ good reliab~l~ty.

U.S.

Fa~ Nb6. 4,157,904, 4,171,964, and 4,278,457 ~nd U.S.

Patent Nos. 4,869,740, is6ued September 26, 1989, and 4,854, 955, issued August 8, 1989, describe relevant processes.

In ~ typical ~ ic Q~sion r~xN~ry pnxx~ ed gRs db~ ~m u~ ~r pr~w~e iB ccoled ~y h~at exdY~e with ~ ~r ~nY~s of the process an~/or external sKIDn~ss of refrigeration Euch a~ a prcpane 25, compression-refrigeration E~i3m. A~ the gas iB cooled, liquids m~y be ccrYh~-s ~ and collected in one or ~re separatcrs as h~gh~pressure liquids ccre~LLnlng 6c0e of the dbsin3d C2~ crl~D~ents. re4e#l1in~ on the richness of ths gas ~nd the ~mount of liguid formed, the high-FIis#~ur~ liguids nay be eX~bd to A lcwer pr~e ~na frackionated.

Ihe ~ rization ooo~ring during e~sion of the liq~ r~t~ in f~r coDling of ~he 6b~m.

Uhder ~om~ co~ions, ~n~oling the high pr~e liquid prior t~ the e~x~sion ~ay be des~n~e ~n on~r ~o fw~r lcwer the te~ature n~ting fnom tbe e~sion. ~he e ~ ~b~m, on~xis~ng ~ h~e of ,, . . . -. . .. , .. , . . . . ...... . . ~ ~ . -. ,- .. . . . ... ... .... . ..... . ....

2()031~7 liquid and vapor, is fractionated in a distillation (deme*handzer) column.

rn ~he oolumn, the expansion-oooled stream(s) is (are) di~tilled to separate residual methane, nitrogen, an~ okher volatile gases as overhead vzpor fr~m the ~esired C2 componRnts, C3 components, and heavier c~rpcnent~ as bcttom li ~ d prGduct.

If the feed gas is nok tot211y ccndbnsed (typi~ally it is not), the vapor remai ~ from the partial condensation can be split into two or more streams. One portion of the vapor is passed ~hrough a work expansion machine or engine, or expansion valve, to a lower pressure at which additionai liquids are condenssd as a result of further cooling of the stream. m e pressure after expansion is essentially the same as the pressure at whid the fractionation column is operated. m e oombined vaporliquid phases resulting fmm expansion are Q lied as feed to the column.

m e remaining portion of the vapor is cooled to substantial oondensation by heat exchange with other pr~cess ~treams, e.g. the cold fractionation tower overhead. Depend~na on the amount of high-pressure liquid available, eome or all of ~he high-pressure liquid may be oombined with this vapor portion prior to ooolin~. The resulting cooled stream is then expanded thrcugh an appropriate expansion de~ioe, such as an expansion valve, to the pressure at which the demethorlzer is o~erated. During expansion, a portion of the liquid will vaporize, resulting in oooling of the tokal stream. m e flash expanded ~tream is then Eupplied as t~p ~eed to the demethanizer. Iypically, the vapor portion of the expanded stream and the demethanizer overhead vapor ocmbine in an up~er, separator section in the fractionation tower as residual methane product gas.

Alternatively, the oooled and expanded stream may be supplied to a eeparator tD provide vapor and liquid streams. The vapor is oombined with the tower overhead and the liquid is suFplied to the oolumn as a tcp column feed.

In the ideal operation of such a separation prcoess, the residue ~c leaving the prooess will ccntain ~ubstanCially all of the methane in the feed gas with essentizlly none of the heavier hydrocarbcn ccmpcnents and the bot$oms fraction leav ~ the de=e*hanizer will oontain ~ubstartially all of the heavier c~nponents with essentially nD methane or ligh~er c~mponents.

--3- 2003~97 In practioe, hcwever, this ideal situation is nGt cbtained for th2 reæson that the conventional demethanizer is cperated largely as a stripping column.

The methane produ~t of the prccess, therefore, typically comprises vapors leaving the top fractionation ~tage of the column, togethQr with vapors nck subjected to any rectification step. C~nsiderable losses of C2 a~ cccur beca~ p li~d feed oontains s~bstantial q~antities of C2 caqpone~ts and heavier ccn~onents, resulting in correcp~ir~ equilibrium q~arltities of C2 ~ and heavier ~nents in the vapors leavir~ the tc~ fractianation stage of the demethanizer. me loss of these desir~le rp~ts could be significantly reduoed if the rising vapors cculd be brcught into corltact with a significant quantity of id (reflux), containir~ very little C2 o0~ents and heavier ~nents; that is, reflux capable of abeorbing the C2 canponents and heavier ~ fran the vapors. me present invention provides the mear~ for ac~ieving this objective and significantly in~aving the reoovery of ~ e desired products.

In accordance with the present inv~.tion, it has ~een found that C2 reooveries in excess of 99 percent can be obtained.

Similarly, in those 1nst3nces w~en reoovery of C2 components i~ not desired, C3 recoveries in exces~ of 99% can be achieved.

In ~d~ition~ the present invention makes possible essenkially 100 peroent separation of methane (or C2 C~l porents) and lighter cc~ponents from the C2 colponeot6 (or C3 ccrponent~) and heavier corpcnentJ at reduced energy regpire cnts.

The present invention, althcugh applicable at lower pressures and warmer temçeratures, is particularly ~dvart~gecus when prooessing feed gases in the range of 600 to 1000 psia or higher under ocnditions requiring column cverh~d temperatures of -110F or colder.

For a better urdLrEt3rl ing of the present invention, reference is made to the followin~ exa~ples and drawings.

Referring to the drawings:

FIG. 1 is a flow diagram of a cryogenic expansion natNral gas processing plant of the prior art ~ny~r~ing to U.S.

Patenk No. 4,157,904;

PIG. 2 is a flow disgram of a cryogenic expansion natural gas -4 Z003197 processing plant of an alternative prior art pr~c#ss aooDrding to U.5.

Patent No. 4,687,499:

FIG. 3 is a flow diagram of a natuxal gas prooessing plan~ Ln accordance with the present invention;

FIGS. 4 an~ 5 are fragmentary flow dia ~ illustrating alternative means of aFplication of ~he present Lnvention to a natNral gas s~n:

FIGS. 6, 7 and 8 are flow diagxame of additional natural gas prooessing plants in accordance with the p J ent invention: and FIG. 9 is a fragmentary flow diagram showing a natural gas process mg plant m aooordance with the present invention for a richer gas stre~m.

FIG. 10 is a flow diagram illustrating an alternative neans of application of the present Lnvention to a natural gas stream from which recovery of propane and heavier hy~mc~rtonL is desired.

In the followin~ explanation of the abDve figures, tables are provided summariz mg flow rates calculated for represen~ative process conditions.

In the tables appearing herein, the values for flow rates (in pound mDles per hcur) have been rounded`to the n2arest whole number for ccnvenience.

The tctal stream rates shown in the tables include all nonhydroc3lton ccnponents and hence are generally larger than the sum of the stream flow rates for the hydrccarbon cc~ponents.

Temperatures indicated Are approximate values, rownded to the nearest degree.

It should also be noted that the process design calculaticns perfcrmed for the purpose of comparing the pr~ocqr~- depicted in the figures are based on the assumption of no heat leak fram (or to) th~ surround1nqG to (or frcm) the prccess. m e ~lity of commercially available insulating materials makes this a very reasonable ass,umption and cne ~hat is typically made by tbose skilled in the art.

-5 ~QO~gt7 Refe~r~ n~w to E~CG. 1, in a sin~lation of the proo~s acoor~ir~ to U.S.

Paterlt No. 4,157,904, inlet gas enter~ ~e plant at 120F and 1040 psia as s~ 21.

If the inlet gas ~s a c~ation of sulfur compounds which would pn~nt the pm~ct stre~ fr~n ~eeting specifications, ~e sulfur ~mls are r~moved ~y appr~riate pretrea~t of the feed gas (r~t illustrat~l).

In additical, ~ e feed stre~n~ lly dehydrated to prevent hydrate (ioe) formation un~er cryogenic oonditioms.

Solid desiccant has typically been ~acP~ for this purpose.

The feed stream lo is divided into tw~ parallel streams, 22 and 23, and cooled to 66F ky heat exchange with cool residue gas at 50F in exchanger 10 and with dbmethanizer liquid at 43F in dea#*h~nizer reboiler 11.

From these e>dh~nqers, streams 22a and 23a reco~bin~ as stream 21a which enters exchanger 12 ~here it is cooled to 28F (stream 2Ib) by cool residue gas at -4F (stream 29a). m e feed gas oo~tinues to demethzr1zer side reboiler 13 and is further cooled by heat exchange with demothanizer liquid at -78F.

The further oooled 6tream 21c then enters separator 14 at 12F and 1025 psia where the vapor (stream 24) is separated from the ccndbn~ed liquid (6tream 28).

Ihe vapor (stream 24) fr~m ~eparator 14 ~ divided into two streams, 25 and 27.

Stream 25, containing abcut 38 peroent of the tokal vapor, is oombined with the ~eparator liquid (stream 2~).

Yhe oombined stream 26 then passes throu~h ~eat exchunqer 15 in 'heat ~ e relation with the demeehonizer overhead vapor stre~m 29 resultin~ in oooling and ~ubstartlal ccndensatian of the comb med stre~m. qhe sub6tantially COlldenied 6traam 26a at -138F is then flash expanded throu~h an apprepriate expansion devioe, 6uch as expansion valve 16, to the operating pressure (aFproxlmately 400 p6ia) of the fractionation tcwer 19.

During expansion a portion of the stream is vaporized, resultin~ in ooolLng of the total stre~m.

In the procYss illustrated in FIG. 1, the expanded 6tream 26b leaving expansion valve 16 reaches a temperature of -141F, and is supplied -6 2()03197 to separator section l9a in the upper region of fractionation tcwer 19. 1 liquids separated therein become the top fe3d ~o dbsrth~niiing section l9b.

The rema ~ 62 percent of the vapor from separator 14 (stream 27) enters a work expansion machine 17 in which mechanical energy is extracted from this portion of the hi~h pressure feed. m e machine 17 ~ s the vapor substantially i ~ ically frcm a pressure of about 1025 psia to a pressure of abcut 400 p6ia, with the work expansion c~oling t~ expanded stream 27a to a temperature sf approximately -75F. ~he typical conmercially available expanders are capable of recovering on the order of 80-85% of the w~rk theoretically available in an ideal isentrcpic expansion.

m e expanded an~ partially c~ndenoed stream 27a is supplied as feed to the distillation column at an intermsdiate point.

Ihe deme*h~nizer in fractionation tower 19 is a csnventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or som2 c~mbination of trays and packin~. AS is often the case in natural gas processing plan~s, the fractianation tcwer may oonsist of tw~ sections. m e upper section l9a is a separator wherein the partially vaporized tcp feed is divided into its respective vapor and li~uid portions, and wherein the vapor rising from the lower distillaticn or demethdrizing section is oombined with ths vapor portion of the tcp feed to form the cold residue gas distillation stream 29 which exits the top of the tower. m e lower, dc=ethanlzing section l9b contains ~he trays and~or packing and provide~ the nec~sory contact between the liquids falling downward and the vapors rising upward.

The demethanizing section also includes reboilers w.hidh heat and vaporize a portion of the liquids flowing dcwn the column to prcvide the stripping vapors whidh flow up the column.

Ihe bcktom product stream 30 exits the bottom of the tower at 69F, based on a typical specification of a nekhane to ethane ratio of 0.025:1 on a molar basis in the bottom product.

The resi~ ~ g~.c (stream 29) ceP~ c _ y to the inoommg feed gas in: (a) heat.c~ch3Dger 15 ~here it is heated to -4F (stream 29a), (b) heat cxnh3ng3r 12 whre it i_ heated to 50F (5tream 29b), and (c) heat exchonger 10 wh2re it 1_ heatsd to 93F ~stream 29c). m e - -7 20(~3~9~7 ca[~ressor 18 driven ~ the ~sial ma~i~e ~ e seoond stage is canpn~;sor 20 driven lty a suE~le~ pawer ~a~roe ~ic~h carpresses the residue gas to 1050 psia (str~ 29e), sufficient to meet line ~i~ 5 ~usually ~n ~e or~ of ~he inlet pressur~).

A su ~ ry of stre~n flow rates and energy oc ~ tion for the prcoess illustrated in FIG. 1 is set for~h in the following table:

~ELE I (FIG. 1) Stream Flow Summary - fLb.

Mbles/Hr) ..

Stream ~ e Ethane Prqpane ~an~* Tctal 21 25381 1161 362 332 27448 24 25323 1149 352 264 27298 28 58 12 10 68 150 9752 4~3 136 102 10513 ..

27 15571 706 216 162 16785 29 25355 118 7 1 25607 ~ .

26 1043 355 331 1841 Recoveries* 20Ethane 89.80% PmQane 98.16% Eutanes~ 99.53% Horsepower Residue Colpre6sian 13,061 *(Eased on un-rounded flow rates) me prior art illustrated in FIG. 1 is limited to the ethane recovery shown in Iable I by the equilibrium at the top of the c~lumn with the tcp feed to the deme*harizer. Lowering the feed gas temperature at ~eparator 14 below that shown in FIG. 1 will not inczease the reccvery appreciably, but will only reduce the power reocvered in the expansion madhine 17 and increase the residue ccmpression horsepower corresponding1y.

The cnly way tc significantly imprcve the ethane recovery of the prior art process of FIG. 1 is to lower the operating pressure of the dem2thanizer, but to do so will increase the residue oompression horsepcwer inordinately.

Even so, the ultimate ethane recovery po6sible will 6till. be dictat~d by the cc~positic,n of the top liquid pG~ to the demethanizer.

-8 2003~9~ One way t~ achieve higher ethane reoovery with the same cemethani~er Op~rd~Lllg ~L~S~me i~ a~e a l~an~r ~l~we~ c2+ w~e~) top feed. FIG. 2 repre~ents an alternative prior art process in aocordance with U.S.

Patent No. 4,687,499 that recycles a portion of the residue gas product to provide a leaner top feed to the des-ehanizer. The prooess of FIG. 2 is based on the same ~eed gas compo6iticn and ccnditions as dbscribed above for FIG. 1. In the simLlation of this prooess, as in the simulation for the process of Figure 1, operatLng conditions were selected to munImize energy oonsumpkion for a given reoavery level. m e feed stream 21 is divided into two parallel streams, 22 and 23, and cooled to 51F by heat ex~hange with cool residue gas at 46F in exchanger 10 and demethanizer liquid at 41F in demekhanizer reboiler 11.

From these exchangers, streams 22a and 23a reoombine and stra~m 21a enters cxrh~nqcr 12 where it is f ~ oooled t~ -31F (stream 2Ib) ~y ccol residue gas at -59F (stream 29a). m e feed gas co~t~nues to deme*h~nizer side re~oiler 13 and is oo~led by heat exchange with d-meth~nizer liquid at -104F.

The feed stream 21c then enters separator 14 at -55F and a pressure of 1025 psia where the vapor (stream 27) i8 separated fr~m the c~ndensod liquid (stream 28).

The vapor from separator 14 (stream 27) enters a wark expansion machine 17 in which nYdhanical energy is extract~d from this portion of the high pressure ~eed. qhe madhine 17 ~ s the vapor n~bEtanelally isentrcpi~lly from a pressure of abcut 1025 pcia to the cperating pressure of the demethanlzer of abcut 400 psia, with the wor~ exeansian coDling the e~Fanded st ~ to a temperature of apprcYim3tely -126F.

The expanded and partially con~en od stream 27a is sypplied as feed tD ~he distillation column at an int r3edlate pDint.

The ~eparator liguid (stream 28) is l~kewise expanded tD 40Q psia by expansicn valve 36, cooling stre~m 28 to -100CF (stream 28a) befcre it is su~plied to the demeeharizer in fr~cSlcn~tion tower 19 at a lower mid-column feed point.

A porticn of the high pressure residue y (~ream 34) is withdrawn from the nain rt~idue flow (stream 29f) to become the top distillation c~lumn feed.

Re~ycle gas stre~ 34 passes ~h h~t ~}}er 15 in heat e3~ar~e relatic~ with t~e cold oeæ~l3}1izer overhead 2[)03197 distillation vapor stream 29 result ~ in cooling and substantial ation of the recycle stream. Ihe cooled stream 34a at -143F is then expanded through an appropriate expansion device, such as expansion machine 16. The madhine 16 expands the stream substantially isentropically from a pressure of about 1040 psia to the demethanizer opera~ ~ pressure of about 400 psia, with the wark e ~ ion cooli~g the expanded stream to a temperature of approxImately -152F (stream 34b). The e~panded stream 3~b supplied to the tcwer as the top feed.

The bottom liquid product stream 30 exits the ~ottom of tower 19 a~ 68F.

The cold residue gas (stream 29) at a temperature of -145F passes coontcrw rrenely to the recycle gas stream Ln heat exchanger 15 where it is heated to -59F (stream 29a).

The residue gas tben passes c~untercurrently to the i~ feed gas in heat exchanqer 12 where it is heated to 46F (stre~n 29b) and in heat exchanger 10 where it is heated to 92F (stre~n 29c).

The residue gas ls then re~ressed in two stages.

Stream 29c is split into parallel stre~ns 31 and 32 which flaw to canpressors 35 aTX~ 18 (driven ~y the expansion machines 16 arx3 17 respectively) for the first stage of ca~;sian. stre~ 31a ar~ 32a reca~ine and stream 29d flows to the second stage of canpression (a~npressor 20 driven ~y a suE~lemental pw1er sa~me) whic~ canpresses the residue gas to the line pressure of 1050 psia (stream 29e). A~er stream 29e is cooled to 120F (stream 29f) ~y heat e~har~er 37, the re~cle stream 34 is wiWrawn an~l the residue gas product (stre~n 33) flows to the sales pipeline.

A sumnary of stream flow rates arx~ energy cans~atian for the pmc~s illustrated in FIG. 2 is set forth in the follawi~ table:

--10- Z003~!37 5AELE II (FIG. 2) stream Flow Sum~arv - (Lb.

Mbles~Hr~ 21253~1 1161 362 332 27448 2724685 1050 285 158 263~0 28 696 111 77 174 1068 34 7259 3 0 0 7300 2932611 ~5 0 0 3~795 o 30 29 1149 362 332 1953 3325352 12 G 0 25495 ~eoave, ries* Ethane 98.99% ~ 100. 00% Eutanes+ 100.00% Horsepower Residue Compression 17,380 *(Eased on unrrcunded flow rates) DESCRIPlloN OF THE INVENTION Exam~le 1 FIG. 3 illustrates a flow diagram of a prooess in accordance with the present invention.

The feed ~ c ~ ition and condltions considered in the process illustrat~ed in FIG. 3 are the same as thcse in FIGS. 1 and 2.

Accordlnaly, the FIG. 3 process can be c ~ with the FIGS. 1 and 2 processes to illustrate the advantages of the present invention.

In the simLlation of the FIG. 3 process, inlet gas enters at 120F and a pressure of 1040 psia as stream 21.

The feed stream is divided into tWD parallel ~ s, 22 and 23, and cool~ed to 69F by heat exchange with cool residue gas at 60F (stream 29b) in exch3rgcr 10 and with de=e*hcnizer liquid at 47DF in dk~e*hunlzer reboiler 11.

From th2se exchangers, streams 22a and 23a re~ombine and stream 21a enters exchonqer 12 where it is oooled to 26-F (stream 2Ib) by cool residue gas at -3F (stream 29a).

Ihe further oooled stream 2Ib o~ntinucs to dbmethanizer side reboiler 13 and is cooled by hea~ exchange with deme~hanizer liquid at -67F.

The feed s~r2am 21c . ,~.. : . . - . .. - . .. ,, . ~. , , . . ,, . . - - ... ., , . , - . .. ... - , , Z0031~ then enters high pressure separator 14 at 12DF and a pressure of 1025 psia where the vapor (~tream 24) is separated frcm ccndensod liquid ( ~ 28).

me vapor (stream 24) from separator 14 is divided int~ g2seous first and seoond streams, 25 and 27. Stream 25, contalning abcut 40 percent of the total vapor, is oombined with the sepaxator liquid (stream 28). The combined stream 26 then passes thrcugh heat exchanger 15 in heat exchange relaticn with the -144F cold residue gas stream 29 resultLng in cooling and substantial ocndbnsa~ion of the combined strelm. me s ~ ially condensed strelm 26a at -142F is then expanded t ~ an app~qpriate expansion devioe, such as expansio~ valve 16, to a pressure of appraximately 412 psia, i.e. 5 psi above the o~era~ing pressure of the fractionation ~ower 19. During expansion, the stream is ocoled to -144F (stream 2~b).

In the prccess illus~ratRd in FIG. 3, the expanded stre~m 2~b reaches a temperature of -144F and flows to the heat exchanger 31. The mixed phase s~re~m 26b is warmed in the exdhanger to -136F and partially vaporized as it provides cooling and sut~tant1al condRnsation of a compres6ed recycle portion (stre~m 40a) of distillation stre~m 39 leaving the top of fractionation tcwer 19. m e warmed stream 26c thRn enters the distillation column or demekhanizer at a lower mid-column feed pa6ition.

The distillation oolumn is in a lower region of fractionation tcwer 19.

m e ~ub~tarti~lly c~n~ersed stra~m 40b leaving exchanger 31 is then expanded thrcugh an appropriate e~Fansion device, such as expansion valve 33, to the demethanizer operating pressure. During expansion a portion of the stream is vaporized, resulting in cooling of the total stre~m. In the prooess illustrated in FIG. 3, the expanded stre~m 40c leaving expansion valve 33 reaches a temperature of -145F, and is supplied to the fractionation tower as the t4p tcwer feed. me vapor portion of strsam 40c co~bines wi~h ~he vapors rising from the tcp fractionation stage of ~he column to form distillation stream 39, which is withdrawn frcm an upper region of the tcwer. m is stream is then divided into two streame.

One portion, stream 29, is the cold volatile residue gas fraction. The okher portion, recycle stream 40, i oompressed to a pressure of about 490 psia, i.e. about 83 psi higher tban the demeehanizer, in oold recycle ., .... . ..... ,. . , . , .. . .. , ., , . . . ,.. ... ,- " .. . ~ . . - . .

~003197 caqpressor 32 . qhis a~ressed rec ycle 6tr~n 40a, naw at abaut -124 F, the31 flows to heat exbha~er 31 where i~ arJc~a~l~ ~y heat e~ge with mixed E~ase stre~n 26b.

P~rr~ to ~q~lc seooald stream 27, the remaining 60 ~t of the vapor fmn se~arator 14 ent~s an e~pansial devioe ~ as work ex~ansion ma~ine 17 ~n which ~ ical ener~3y is extrac~3d f m n ~is po~tion of the high pressure feed.

The machine 17 expands the vapor subetantially isentropi ally from a pressure of abcut 1025 psia to the pressure of the demethanizer (akout 407 psia), with the w~rk expansion caoling the e ~ stream to a temperature of approximately -74F (stream 27a).

The expanded and partially c~ndensei stream 27a is supplied as feed to the distillation colu~n at a seoond mid-column feed point.

The liquid product stream 30 exits the bct~om of tower 19 at 72F.

The cold residue gas stream 29 passes c~untercurrently to stream 26 in heat exchanger 15 whel~ it is heated to -3-F (stream 29a) as it provides cooling and sub6tantial condensation of stream 26. ~he partially warmed strsam 29a then flcws to heat exchanger 12 where it is further warmed to 60F (stream 29b) as it provides caoling of strEam 21a. m e further warmed residue gas stre~m 29b then flows to heat exchanger 10 where it is heated to 98F (stre~m 29c) as it pr wides cooling of inlet gas stream 22.

The residue gas is then ~ -cc pre~sed in tw~ stages.

The first stage of co0pressian is co~pressor 18 driven by the e~ ~ nsion machine 17.

The s~oo~ etage of cc0çression is oo~pressor 20 driven by a sypplemertal power scurce which comprssses the residue gas (stream 29d) to the line pressure of 1050 psia.

A summary of st~re~m flow rates and energy consumptlon for the process illustrated in FIG. 3 is set forth in the followLng table:

. .

~003~97 qABLE III (FIG. 3) Stream Flow Sun~(Lb. MDles/Hr~ 2125381 1161 362 332 27448 2425323 1149 352 264 27298 28 58 12 10 68 150 2510096 458 140 105 10883 2715227 691 212 159 16415 3931215 10 0 0 31414 40 5863 2 0 0 5900 2925352 8 0 0 25514 29 1153 362 332 1934 P~ri~s* EthAne 99.31% 100. 00% ~+ 100. 00% Horse~ower Residue Co~pression 12,968 Cold Recycle Compression244 Ictal Horsepswer 13,212 *(Eased on unrrounded flow rates) Using the refrigeration available in flash e~xux~3d stream 26b to pr wide cooling and ~b6tantial oondensation of stxeam 40a in e ~ er 31, it i8 po~sible to ninimize the bcrsepcwer required for conpression of the recycle stream.

It is n-cesuRry to ccnçress the recycde stream 40 only to a pressure at whi~h it ¢an be Ewbsezrtlally ocnden~ed at a temperature warmer than that of flash expanded stream 26b.

Typically this involves raising the pressure of recycle 6tream 40 by 70-100 psi.

This pravi~pc a tcp oolumn feed oontaining very little ethane and heavier hyc~czrbcns, which rectifies the vapors stripped from the feeds in the lower porticns of the oolumn, t ~ y reducing the equilikrium lc~ses of ethane and heavier hydrrla~tcns in the fractionation tower overhead stream 39 and Euboegp~ntly m the residue gas stream 29.

q Eon of the recovexy levels display2d in Tables I and III shows that the present invention improves ethane recovexy fm m 89.80% to 99.31%, prcpane recovery from 98.16% to 100.00~, and butan~s~ recovery from 99.53% to 100.00%.

CcmparisDn of TableL I and III further shows that the -14 ;~003~9~ improvement in yields was not s ~ ly the result of increasing the horsepower (utility) refuirements. Tb the ~ , when ~he present invention is employed, as in Example 1, not orly do ethane, prcpane, an~ ~utanes+ recoveries increase over those of the prior art process, but liquid recovery efficiency also increa~e~ by 9.3 percent (in terms of ethane recovered per u mt of hor~epower expen~ed).

Comparing the present m vention to the prior art process displayed in FIG. 2, Tables II and III show that the FIG. 2 prior art prooe ss essentially matches the reoovery levels of the present mvention for C2+ ~ onents. However, the FIG. 2 prcoess does so at the expense of greatly inorcused horsepower (utility) oonsumpkion. The present invention achieves the same recovery levels using only 76 percent of the extern21 power required by the FIG. 2 prior art process.

Example 2 FIG. 3 represents the preferred e~ixcirent of the present i~vention for the temperature and pressure conditions shown.

Anckher e~bodllent of the present invention is illustrated in the fragmentary proceæs flow dlagram shcwn in FIG. 4. rn the simulation of the process depicted in FIG. 4, the inlet gas oooling sdheme is identical to that used in FIG. 3.

The difference lieæ in the dispc6ition of the recycle stream 40 to be cGrpre3s d in the oompressor 32.

Rather than cowpressing the recycle stream directly into heat exchanger 31, the stream can first be warmed so that cryogenic metallurgy is not required in the oompressor.

One method of accomplishing this is as shown in FIG. 4, where the re~ycle vapor streæm 40 enters cross cx~h0rger 34 and is heated to 115CF by heat exchange with the wanm recycle co~Fre~sor dlsch~rgc stream 40c.

The warm stream 40a enters the warm recycle oarpreYsor 32 and is compressed to about 500 psia (stream 40b).

The oomlpressed stream is then oooled to 120~F in heat exchan~er 35 (stream 40c).

After cross exchange with the oool ~tream 40, stream 40d at -130~F enters excharger 31 where it is ccoled and substantially ooobeo~ed by heat exchange with ~ixed pbase stream 26b as diecus~ed previcusly.

The subst~nt1311y oondenoe~ stream 40e is th~n flash expanded in expansion valve -15 Zc)0~3~ 33.

The o~ld, flash expanded stre~m 40f, now at about -145~F, is supplied a~ the top fesd t~ fr~ _^r. t~ _a, A summary of stxeam flow rates and energy consumpkion for the process illustrated in FIG. 4 is eet forth in the following table:

la~LE IV (FIG. 4) Stream Flow ~ - (Lb. M~les~Hr) 21c25381 1161 362 332 27448 2425323 1~49 352 264 27298 28 58 12 10 68 150 2510096 458 .140 105 10883 2715227 691 212 159 16415 3931217 10 0 0 31406 405864 2 0 0 59~0 2925353 8 0 0 ~5506 30 28 1153 362 332 1942 ~eries* Ethane 99.31% Prcpane 100.00% BU~+ 100. 00% Honsepowex ~esidue Compression12,742 Wbrm Fe~ycle Compression 770 Tbkal Horsepower 13,512 *(Based on un-rounded flow rates) A comparison of Tables III and IV shcws that the FIG. 4 emtodiment of the present invention can maintain high recovery levels with a slight increase in the hcrEepcher (utility) requ1rcments- m e choioe between compressing stream 40 cold or warm d~pends on factors such as plant size and available equipment.

Exam~le 3 A third embodiment of the presenk invention is illustrated in the fragmentary flaw diagram shnwn in FIG. 5. In the simLlation of the process depir~ in FIG. 5, the inlet gas cool mg ~ is idbntical to that used in FIG. 3 and FIG. 4. The difference lies in the methcd used to cool and sr~gtanti311y condense the oonpresse~ recycle s~ream (40a). In the c bcdiment of FIG. 5, the compressed stream 40a ~s ccoled and srlYd~r~i~lly -16 Z003197 condensed in ex~ er 36 ~ich f~tions ~ a side rf~boil~ to the demet;hanizer. ~o~ is pr~ri~ea by a liquid ~ream withdra~n frc~n 'che dis~tillation column in fracti~ation ~r 19, str~n 41, at a te~erature of -143 ~F.

Ihi~s liq~d stream ~s heated a~ partially vaporized in e~char~er 36 and the mixed ~ str~ m 41a, naw at -137~F, is returned to the distillation column.

The sub6tantially ccndensed racycle stream 40b frcm exchanger 36 is at a temperature of -141F and is flash expanded to the operatLng pressure of the fractionation tawer in expansion valve 33. m e flash expanded stxeam 40c, now at about -145F, enters fractionation tower 19 at its top most f~o~ point.

Eecause stream 26b from expansion valve 16 is ~ct used to provide cooling of the oompressed recycle stream 40a, this mixed pbase stream flows - direc*ly to distillation column at a middle column feed position.

A summary of stream flow rates and energy consumption for the process illustrated in FIG. 5 is set forth in the following table:

(FIG. 5) streæm Flow summary - (Lb.

Mbles/Hr~ Stream ~banLe ~baaLe Prc~ane Eutanes~ 5~k21 21c25381 1161 362 332 27448 2425323 1149 352 264 27298 28 58 12 10 68 150 2510096 458 140 105 10883 2715227 691 212 159 16415 3931216 10 0 0 31415 405863 2 0 0 5900 2925353 8 0 0 25515 28 1153 362 332 1933 Reccveries* -~ Ethane 99.32% Propane 100.00% ~+ 100. 00% ~orsepower Residue Compression 13,008 Cold Recycle oompression247 Tbtal Horsepower 13,255 *(Based on un-rounded flow rates) ~ .

-17 ~003197 A c mDarlson of Tables III and V shcws that the ~bcdl~ent of FIG.

5 is only sliqhtly less efficient than that of FIG. 3.

It shculd also be noted that stream 28 in Figure 5 need not be co~bined with stream 25.

Alternatively, all or a pDrtion of stream 28 may be e ~ to the operating pressure of the column and then sNpplied to the - c~lumn at a mld-oolumn feed pnsition.

Example 4 A fourth embodiment of the present invention is ~ in FIG. 6.

The feed gas oomposition and conditions oonsidered in the process illustra W in FIG. 6 are the same as ~hose in FIGS. 1 throu~h 5.

In the simLlation of the prooess of FIG. 6, the inlet gas oooling and expansion scheme is identical to ~hat u6ed in FIG. 3.

The difference lies in where the gas stream to be compreosed, sub6tantially condb=sed and used as top tcwer feed to t!he demethurizer is withdrawn frcm the distillation streYm 39.

Referring to FIG. 6, the warmed distillation streæm 39a is divided into two streams, 29 and 40.

Stream 29 is the residue ~as fraction whidh is reroompressed in two stages as previously discussed.

Stream 40 is the recycle stream whidh is occpressed in warm recycle oo~reDor 32 to abalt 500 p~ia, i.e~ 90 p~i above ~he cperating pse of the oe~unizer.

The ca~ressed stream 40a is oo~led to 120-F (stream 40b) in ex~ihar~er 34. (oe~llr~ on the tenperature of the ~tream folla~ing calpre~;sion, e~ 34 may not be r~y.) q~e oooled stream 40b is then fur~r cooled in e~ar}y~r 35 to a tenperature of abalt -131F (stream 40c) ~y heat ex~ange with a partion (stream 36) of distillation stream 39.

~e further cooled stream 40c then er.ters e~r 31 wh~e it is ~bstantially oor~9c sei ~y heat ex~ange with the mixed p~ase stream 26b from expansion valve 16 as oisoussed previously for FIG. 3.

The I slbst4ntially oandenscj stream 40d is then flash expanded in expansion valve i 33 and the expan~ed stream 40e flows ~ tcp f~c~ at -145F to fractionation 0 tcwer 19.

-18 2003~9~ Distillation stream 39 leaves ~e uE~er r~gion of the tawer at warmed to abalt 96F (stream 37c) as it pr~vides cooling of the inlet feed in e~*~angers 15, 12 and 10.

Ihfe ath~r portion, stre~n 36, prt~ C cooling 5 of the calpressed rec~rcle str~n 40b in ex*lar~er 35 as di~ss~ previ~ ly. Ihe thus warmed str ~ns 36a and 37c recaribine as the warmed fractionation tcwer overhead stream 39a which is then divided as described above.

A summary of stream flow rateC and energy consumption for the prooess illustrated in FIG. 6 is set forth in the following table:

I~E VI (FIG. 6) Stxeam Flow SuDm~rY~LLk~ es~Hr) Stream ~hBn~ Ethane E~5~gLLe Eutanes~ l~tal 2125381 1161 362 332 27448 2425323 1149 352 264 27298 28 58 12 10 68 150 2510096 458 140 105 10883 2715227 691 212 159 16415 3931119 10 0 0 31302 40 5766 2 0 0 5800 2925353 8 0 0 25502 28 1153 362 332 1946 Recoveries* Ethane 99.33% Prcpane 100.00% Eutanes~ 100.00% ,F~ Residue CoIpression 12,754 Wanm Recycle v sion . _ 833 q~tal Ebrbepowcr 13,587 *(Based on un,rounded flow rates) A co~parison of qables III and VI shcws that the FIG. 6 emboddmcnt of the present invention oan ~lso maintain high recovery levels with only a slight increase in hcrse$cwer (utilit~) c~osu=ptlon. Ihe dhDi oe of where to withdraw recycle stream 40 in tbe ~ dbpends on factors which include plant size and awailable eqyiF ent.

-19 X0031~37 Alternatively, ~he prccYss of the invention can be operated to attaLn less than naxImum reoovexies.

If this is Ap~ired, the flow rate of recycle stream 40 is reduoed, thus reduc mg the quantity of top oolumn feed to fractionation bower 19 in FIG. 3.

As this flow rate is reduced, less reflux liquid i c available for the oolumn resulting in a warmer column overhQad te~perature.

Ethane reoavery then dkcIe3ses.

This is illustrated in thY followLng example.

Exam~le 5 m e prooess of the present invention is also applicable for processinq gas streams whQn it is desirable to rec~ver only the C3 c~rpcnenes and heavier hydrocarbon c~rpcDents (rejection of C2 ocmponents and lighter ccrponends to the residue gas).

Such an e~bodiment of the present invention is shown in FIG. 10. m e feed gas oomposition and oonditions considered in the process of FIG. 10 are the same as those in FIGS. 1 through 6.

Eecause of the warmer prooess operating conditions associated with propane reoovery (ethane rejection) operation, the inlet gas cooling scheme i8 slightly different than for the ethane recovery cases dlscs~sed previously.

Referring to FIG. 10, inlet gas entera the prccess at 120F and 1040 ~psia as atream 21.

This stream is cooled to 23-F (stream 21a) by heat exchange with cool resi~ue gas at -15F ~stream 29a) in exchanger 10.

From this e~changer, 6tream 21a is ~ r cooled to 12F (stream 2Ib) by the expander outlet 6treEm 27a.

The feed stream 21b then enters the high pressure separator 14 at 1025 psia where the vapor (stream 24) is separated from the condensed liquid (~tream 28).

~he vapor (stream 24) from separator 14 is divid3d into gaseous first and seoand streams, 25 and 27.

Stream 25, oontaining akcut 35 percent of the tatal vapor, is ~nbined with ~e ~separator liquid (s~m 28). ~e combined stream 26 then passes through heat exdhang3r 15 in heat exchange relation with the -116F cold resi~ gas fraction 29 reQ~ltir~ in cooli~ and subrt3rti3i cnndensation of the combined stream.

The cwt~tankially con~ensed stream 26a at -112-F is then expan~ed thrcu~h an aFpropriate 20 20~3~97 e~pansian devioe, ~h as expansian valve 16, to a pre~ of ~tely 41~ p6ia. ~ng expansion, a por~i~ or ~he ~m will ~ orize, re~lt~ng in cool~r~ o~ the total stream.

5 rea~hes a ten~erature of -138F and flows to ~e heat e~ 31.

Ihe mi~ d p~ase s ~ 26b ~s warmed ~n the e ~ to -135F and partially vaporized as it providbs cool ~ , resulting m ~Db~t3rei~1 condensation of a portion (stream 40a) of distillation ætream 39 leaviny the top of tcwer 19.

m e wzrmed stream 26c then enters the deethanizer distillation column in fractionation tcwer 19 a~ a mid-column feed position.

m e s~bstantially ccD~enNel stream 40b leaving exchanger 31 is then ~xpandbd thrcugh an appropriate ~ sion dbvi oe, subh as expansion valve 33, to the deethanizer operating pressure. During expansion a portion of the ~tream is vaporized, re~ulting in cooling of the total stream. In the process illustrated in FIG. 10, the expanded stream 40c leaving e~pansion v lve 33 reaches a temperature of -141F, and is supplied to the fractionation t~wer as the top feed. ~he vapor portion of stream 40c ccmbines with the vapors rising ~rom the top fractionation stage of the column to form distillation ~tream 39, which exits the tcp of the t~wer.

~his stream i6 then divided into two streams. Qne portion, stream 29, is the volatile residue gas fraction. The ckher portion, recycle stream 40, is cocprosso~ to a pressure of abaut 485 psia, i.e. about 78 psi higher than ~he deethanizer, in cold recycle compreasor 32. ffl is ccrpcessed stream 40a, now at about -95-F, tlhen flows to heat eYoh~nger 31 where it is oooled and subctantially condensed by heat exchange witlh m~xed phase st~eam 26b.

Returnlng to gaseous second st~eam 27, the remzini*a 65 percent of the vapor fr~m ~eparator 14 enters an expansion devioe such as work expansion machine 17 ~c described in earlier cases. The expansion machine 17 expands the vapor to a pressure of about 412 psia ther~by oooling the expanded stream to a temperature of -72F (stre~m 27a). The expanded and partially c~Y dkcsed stream 27a then flows to Y=tz*gsr 11 where it is warmed as it prcvide6 oooling of the inlet gas stream. The warm expanded -21 - X0031~ stream 27b, naw at a ten~e~a~re of abaut -56F, is ~e~ ~lied to the deethanizer a~ a seco~ mid-ool~ feed poeition.

me ~izer ~lu~es a r~boiler 12 ~ic~h hea~s and va~orizes a pt~rtiorl of the liquids flawir~ ~awn the col~ to pravide the st~ippirlg rejection), the tcwer reboiler ~emperatures are ~ignifican~ly ~ r than ~hen operating as a demethanizer (ethane reoovery).

Generally this mak2s it Lmpossible to reboil the tower using plant inlet f~c~ as is typi~lly done for ethane reoovery cperation.

Therefore, an external source for reboil heat is normHlly employed.

The liquid produst stream 30 exits the bottom of the tower at 220F, ~ased on a typical specification of an ethane to propane ratio of 0.025:1 on a molar basis in this liquid product. me cold distillation overhead streEm 39 exits the column at abcut -116F and is divided into two streams (29 and 40), as ~ CCP~ above.

The cold residue gas str~am 29 is warmed to about 110F as it provides cooling of the inlet gas stream in exchangers 15 and 10.

It is then oo~pressed in t~D stages by co~pressor 18, driven by expansion machine 17, and ccmpresLor 20, driven by a supplemental power source.

A sumnary of the stream flow rates and energy u ptic~ for the prooess illustrated in FIG. 10 is set forth in the follcwing table:

2003~9~ (FIG. 10) Stream Flow Summary - lLb.

Mbles/Hr) Stream ~thane Ethane Pr~Qane Eutanes+ lctal 2125381 1161 362 332 27448 2425323 1149 352 264 27298 28 58 12 10 68 150 258863 402 123 92 9554 2716460 747 229 172 17744 3927754 1260 0 0 29245 402373 108 0 0 2500 2925381 1152 0 0 26745 0 9 362 332 703 Reooveries* Ethane l-~qc than 1% Prqpane 99.90% R~+ 100. 00% Horsepower Residue Compression 13,562 Cold Rcycle Compression ~0 I~kal Horsepower 13,682 *(Eased on un-romded flow rates) As can be seen frcm the akove table, more than 99 percent of the ethane is rejected to the residue gas stre~m while still ~aint~in1ng in Pxroc~ of 99.9 percent prqpane recovery. qhis requires a lesser quantity of recycle stre~m 40 than was used for ethane reccvery qperatiQn.

Althcu3h it i8 nok necessarily the optimum approach, it should also be reoognlzed that the flow sdheme presonecd in Figure 10 can also be e~ployed when ethane tC2 cowçonent) r q is desired.

Okher Embcdlnents Ihe high pressure liquid s~ream 28 in FlGS. 3 through 6 an~ 10 need not be combined with the portion of the Eeparator vapor (stre~m 25) flowing to exnh~nqer 15. Alternatively, ~trsæm 28 (or a portion ther~cf) may be expanded thrcugh an apprcpriate~, ~ ion dbvi oe, sudh ~.~ an, expansion, valve or expansicn machine, and fed tD a third mid-oclumn feed point an the distillation colun,n. Stream 28 may also be used for inlet gas -23 200~g7 cooling or okher heat exchange servioe before or after the expansion step prior to Elowi~ ~o the demethanizer.

In inst3rcY~ whre the inlet gas is richer than that heretofore described, an embodi~eDt such as that depicted in FIG. 9 s y be employed.

Ccodensed stream 28 flows through xchonger 4~ whre it is ~ oled by heat exchange with tbe cooled ~tream 42a fi~ expansion valve 43. The sukcooled li ~ d (stream 28a) is then divided into two portions. The first portion (stream 42) flows thrcugh expansion valve 43 where it urlbrgces expansion and flash vaporization as the pressure is reduced to abcut the pressure of ~he fractionation tower. me cold stre~m 42a fram e~pansion valve 43 then flows through e#rho~ger 45, ~here it is used to subcool the liquids fram separator 14. Fram exchanger 45 the stream 42b flows to the distillation column in fractionation to~er 19 as a lower mid-column feed. m e second liquid portio,n, stre~m 41, 6till at high pressure, is either: (1) combined with portion 25 of ~he vapor stream fram separator 14, (2) combined wit!h subst~neiclly ccndcrE~d 6tream 26a, or (3) expan~ed in expansion valve 44 and thereafter either Q lied to the distillation column at an upper midcolumn fe~d po6ltion or co~bin3d with e~panded straam 26b. Alternatively, portions of stream 41 may follow more than one and indeed all of the flow paths heretofcre described and depicted in FIG. 9.

In acoordbnco with this invention, the splitting of the vapor feed may be accompli~hed in several ways. In the preoeose3 of FIGS. 3 throu~h 6, 9 and 10, ths splitting of vapor occurs following cooling and separation of any liqyids which may have been formed. qhe high pressure gas may be split, 25 haw~, prior to ar~ cooling of the inlet gas as sha~ in E~[G. 7 or after the coolirlg of the gas and prior to any E;~paration stages as sha~ in EIG.

8.

In sane ~i~t~, vapor splitting may be effected in a s~ r.

Alternatively, ~e ~epa~tor 14 in the proo~ses s21aw.n in FIGS. 7 and 8 may be uu~y if ~3e inlet Tc is relatively lean. ~5areaver, the use of 30 e~ct~nal ~efrigeraticn to supplement the coolir~ ava;lable to ~e inlet gas fm m other ~ strea~ m~y be e~ployed, partia~larly in ~e r~ce of an inlet gas ridher than that used in Example 1. The use and distribution of desethanizer liquids for process heat exchange, and the particular -24 2~)03~97 ~rr~ of heat ex~rs for inlet gas a~ol~ must also be ~valuat~dfor each partiallar a~lication, as well as ~e c~oice or p~Of~ for ~;pecific heat ex~an~ services.

For e~le, the seocr~ stream depicted in ~IG. 8, s~n 25, may be cooled after division of the inlet 5 st~ and prior to expansion of the seo~d stre~.

It will also be ~ zed that the relative am ~ t of feed fa in each branch of the split vap~r feed will depend on several fact4rs, Lncluding feed y pressure, f~ed gas composition, the amount of heat which can eoonomlcally be extracted frcm the feed and the quantity of horsepower available. MDre feed to the tcp of the oolumn may ~ recovery while decreasing power recovered frcm the expander thereby increasing the reo~pression horsepawer requirements. Increasing feed lower in the oolumn reduoes the horsepower consumpkion but may also reduce prcduct recovery.

me mid-column feed positions depicted in FIGS. 3 through 6 and lO are the preferred feed locations for the prcoess cperating conditions described.

Hbwæver, the relative locations of the mid-column feeds may vary de}csd~ng on inlet oomposition and ckher factors such as desired reoovery levels and amcunt of liquid formed during inlet gas cooling. Mkreover, two or more of the feed s~reams, or portions thereof, m~y be oombined depending on the relative temperatures and quantities of the individual streams, and the combined stream then fed to a mid-column feed position. FIGS. 3 through 6 and 10 are the preferred en~xxiimenks for the cowpositions and pressure conditions ~hown. Although individual stream expansion is depicted in particular expansion devices, alternative expansion means may be employed where o~plu~riate. For example, conditions may warrant w~rk expansion of the sub6tantiPlly o~n~ensed portion of the feed 5tream (26a m FIG. 3) or the sub~t~nei~lly condensed recycle stream (40b in FIG. 3).

me e~bodiments shown in Figures 3 tbrough 9 can also be us~d when it is dbæirable to reccver only t~e C3 oompon~nts an~ heavier c=lçonents (C ccrpcnent rejection). This is acoomplished by appropriate adjustment o~ the column feed rates and conditions.

While there have been dbscribed what are kPlieved to ke preferred e~LcriDYnes of the invenkion, those skilled m the art will recognize that -25- ` 20~31~7 ather and furth~ Drldifications may be made ~ereto, e.g. to a~t the inv~tiaal to varia~s canditiorls, types of reea, or otner ~ir~ wi~out departi~ fr~n the ~irit of tlle pres~t i3n~entia~ ~ defined ~y the follawi~ claims:

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