Catalytic Reforming of Naphtha in Petroleum Refineries

CHAPTER

5

Catalyti c Reformin g of Naphth a in Petroleu m Refinerie s M . DEAN EDGAR Catalyst Department American Cyanamid Company Houston, Texas

I.

Introduction A. Definition of Catalytic Reforming B. History C. Purpose of This Chapter II. Feed Components and Reactions A. Hydrocarbon Types B. Reactions C. Feedstock III. Process Description A. Unit Classification B. Purpose of Reforming C. Unit Design Variables IV. Catalysts A. Dual Function B. Substrate Form C. Physical Properties D. Promoter Metals E. Benefits of Bimetallic Catalysts F. Start-up and Presulfiding G. Poisons H. Coking Deactivation I. Regeneration J. Total Life V. Operating Variables A. Reactor Inlet Temperature B. Feedstock End Point C. Water-Chloride Balance VI. Future Reforming Growth References

Applie d Industria l Catalysis , Volum e 1

123

124 124 124 124 125 125 125 128 129 129 133 1336 136 136 137 137 138 138 *39 142 142 144 144 144 146 147 147 148

Copyrigh t ' 1983 Academi c Press , Inc . All rights of reproductio n in an y form reserved . ISBN: 0-12-440201-1

124

M. Dean Edgar

I. A.

Introductio n

DEFINITIO N O F CATALYTI C R E F O R M I N G

Catalyti c reformin g is a refiner y proces s in whic h a naphth a feed ( C 5400 F ) is passe d throug h severa l reacto r bed s of catalys t at hig h temperatur e an d moderat e pressur e t o achiev e a n increas e in th e aromati c conten t of th e naphth a or a n increas e in its octan e number . Normall y th e naphth a ha s bee n hydrotreate d t o remov e impuritie s tha t eithe r inhibi t th e reaction s or poiso n th e reformin g catalyst . Th e naphth a ca n b e obtaine d directl y fro m th e crud e uni t or fro m th e fractionate d produc t of anothe r refiner y process , suc h a s a coking . Th e catalys t is generall y a few tenth s percen t platinu m (in admixtur e wit h othe r nobl e metal s an d a halogen ) supporte d on a pur e alumin a base .

B.

HISTOR Y

Th e catalyti c reformin g proces s wa s develope d durin g 1 9 4 7 - 1 9 4 9 [1]. In th e nex t 7 yr , 13 ne w commercia l reformin g processe s wer e develope d an d license d b y variou s petroleu m an d engineerin g companie s [2]. Th e proces s ha s continue d t o evolve throug h th e years , th e lates t developmen t bein g a desig n in whic h th e catalys t move s continuousl y throug h th e reactor , fro m reacto r t o reactor , an d finally t o a regeneratio n vessel. In additio n t o change s in th e proces s design , th e catalys t use d in reformin g ha s bee n modifie d t o offer improve d performance . Th e mos t significan t development , whic h occurre d in th e lat e 1960s, is th e bimetalli c (platinu m rhenium ) reformin g catalyst . Accordin g t o dat a reporte d in th e Oil and Gas Journal [3] a s of Januar y 1, 1981, U.S. reforme r capacit y wa s 4,051,400 barrels/strea m da y (bbl/sd) . Seventy-fiv e percen t of thes e unit s use d a bimetalli c reformin g catalyst , wherea s th e remainin g 25% use d a straigh t platinu m catalyst . Thi s repre › sent s abou t 290 reformin g unit s in th e Unite d States .

C.

PURPOS E O F THI S CHAPTE R

Th e inten t of thi s chapte r is t o presen t a n overvie w of th e reformin g process . It will cover feedstoc k properties , reactions , proces s descriptions , catalysts , an d operatin g variables .

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Catalytic Reforming of Naphtha in Petroleum Refineries

II .

125

Feed Component s and Reaction s

Th e hydrocarbo n feed t o th e reforme r is usuall y a depentanize d strea m wit h a 400 F m a x i m u m A S T M D-86 distillatio n en d point . Thi s feed ca n origi › nat e fro m a crud e distillatio n uni t or fro m th e fractionatio n of product s fro m anothe r proces s uni t suc h a s a coker . Hydrocarbon s in th e feed tha t hav e fewer tha n six carbo n atom s (C 5 ) ar e no t considere d t o b e involve d in th e reactions . It is desirabl e t o remov e the m fro m th e feed, sinc e thei r presenc e physicall y interfere s wit h th e acces s of th e reformabl e hydrocar › bon s t o activ e sites on th e catalyst .

A.

H Y D R O C A R B O N TYPE S

Beside s carbo n number , th e component s of th e feed ar e groupe d b y th e followin g types : paraffins , naphthenes , an d aromatics . Paraffin s ar e saturate d straight - or branched-chai n hydrocarbo n mole › cules . Straight-chai n molecule s ar e calle d norma l paraffins , an d branched chai n molecule s ar e referre d t o a s isoparaffins . Naphthene s ar e saturate d rin g compound s tha t ma y hav e sid e chain s attache d t o th e ring . Aromatic s ar e rin g compound s in whic h th e carbo n atom s ar e bonde d b y resonatin g singl e an d doubl e bonds . A six-carbo n benzenoi d rin g is th e basi c aromati c structur e t o whic h sid e chain s or othe r ring s ma y b e attached . Unsaturate d hydrocarbo n compounds , classe d a s olefins , reac t rapidl y wit h th e reforme r catalys t t o for m cok e in th e reformer . Generall y thes e compound s ar e saturate d durin g th e feed preparatio n ste p in th e hydrotreate r whic h nor › mall y precede s th e reformer .

B.

REACTION S

Th e mai n reaction s tha t occu r durin g th e reformin g proces s ar e naph then e dehydrogenation , naphthen e isomerization , dehydrocyclization , par › affin isomerization , an d hydrocracking . Example s of eac h of thes e reaction s ar e show n in Fig. 1. 1.

Naphthene

Dehydrogenation

Naphthen e dehydrogenatio n is a relativel y fast reactio n in whic h naph › thene s ar e converte d t o aromatics . Mos t of th e naphthen e dehydrogenatio n is complete d in th e first reacto r of th e reformer . Becaus e thi s reactio n is highl y endothermic , ther e is a substantia l reductio n in temperatur e acros s

M. Dean Edgar

126 1. NAPHTHENE DEHYDROGENATION

CH

CH

2

2

^

CH?

CH2 XH

+

3H

CATALYST

2

2

CYCLOHEXANE

BENZENE

2. NAPHTHENE ISOMERIZATION £H

2

CH

2 CH 2

CH

CH CH

2 2

2

"

CH

7 CATALYST

CYCLOHEXANE

I CH

2

CH—CH

2

CH

I

3

2

METHYLCYCLOPENTANE

3. DEHYDROCYCLIZATION

+

CH3(CH ) CH3

24

4H

CATALYST

2

NORMAL HEXANE BENZENE 4. PARAFFIN ISOMERIZATION

-C H 3 CH3(CH ) CH3

24

CH3 CH"CH -CH -CH3

2 2

CATALYST

NORMAL HEXANE

2-METHYLPENTANE

5. HYDROCRACKING

CH3(CH )7CH

2

NONANE

3

+

H

2

7 CATALYST

CH (CH ) CH

3

22 3

BUTANE

+

CH (CH ) CH

3

23 3

PENTANE

Fig. 1. Examples of reforming reactions.

th e firs t reactor . Temperatur e decrease s in excess of 100 F ar e c o m m o n for a midcontinent-typ e naphtha . Thi s reactio n is catalyze d b y th e preciou s meta l portio n of th e catalyst . Thi s reactio n produce s hydrogen , an d its rat e is slowed b y hig h hydroge n partia l pressures . Th e conversio n of naphthene s t o aromatic s produce s a n increas e in produc t density .

5

2.

Catalytic Reforming of Naphtha in Petroleum Refineries

Naphthene

127

Isomerization

Naphthen e isomerizatio n reaction s procee d quickl y b y actio n wit h bot h th e acidi c (halogen ) portio n of th e catalys t and , t o a lesser degree , th e preciou s meta l portio n of th e catalyst . Thi s reactio n produce s a rearrange › men t of th e molecul e wit h n o additio n or loss of hydrogen ; therefore , th e reactio n rat e is virtuall y unaffecte d b y pressure . Th e exothermi c tempera › tur e effects associate d wit h naphthen e isomerizatio n ar e usuall y smal l enoug h t o go undetecte d in a commercia l reformin g unit . 3.

Dehydrocyclization

Dehydrocyclization , a n importan t octane-enhancin g reactio n in whic h paraffin s ar e converte d t o aromatics , is a relativel y slow reactio n catalyze d b y bot h th e preciou s meta l an d th e aci d portion s of th e catalyst . Thi s endothermi c reactio n usuall y occur s in th e middl e t o th e last reactor s of th e reforme r unit . Dehydrocyclizatio n produce s hydrogen , an d its rat e is inhib › ite d b y hig h hydroge n partia l pressure . Dehydrocyclizatio n reaction s in › creas e th e densit y of th e product . 4.

Paraffin

Isomerization

Paraffi n isomerizatio n is a relativel y fast reactio n catalyze d mainl y b y th e aci d functio n of th e catalyst . Lik e naphthen e isomerization , thi s reactio n produce s a rearrangemen t of th e molecula r structur e wit h n o ne t chang e in hydroge n production . Th e rat e of paraffi n isomerizatio n is no t strongl y affecte d b y hydroge n partia l pressure . Exothermi c temperatur e effects asso › ciate d wit h paraffi n isomerizatio n ar e no t usuall y detecte d in a refiner y reformer . 5.

Hydrocracking

Hydrocracking , breakin g long-chai n paraffin s int o smaller-chai n paraf › fins, is mostl y catalyze d b y th e aci d functio n of th e catalyst . Thi s relativel y slow reactio n is generall y undesired , sinc e it produce s excessive quantitie s of light e n d s C 4an d lighte r hydrocarbons an d cok e an d consume s hydro › gen tha t coul d b e use d elsewher e in th e refinery . Th e rat e of hydrocrackin g is enhance d b y hig h uni t pressure . Hydrocrackin g is exothermi c an d normall y occur s in th e last reactor . I n som e case s enoug h hydrocrackin g occur s t o produc e a temperatur e increas e acros s th e last reactor . Hydrocrackin g reaction s reduc e th e densit y of th e product . Th e precedin g wa s a genera l descriptio n of reforme r reactions . A mor e detaile d treatmen t of reforme r reaction s is containe d in th e boo k b y Gate s et al [4].

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C.

M. Dean Edgar

Feedstoc k

I n Tabl e I ar e thre e example s of naphth a feeds. Th e compositio n of th e feed ha s a majo r effect on th e reacto r temperatur e require d t o achiev e a desire d produc t octan e valu e an d o n th e quantit y of reformat e yield ob › tained . Th e highe r th e paraffi n conten t of th e feed , th e harde r it is t o reform . A high-paraffi n naphtha , suc h a s light Arabia n naphtha , require s highe r reacto r temperatures , produce s less reformate , an d cause s shorte r cycle length s tha n th e othe r naphtha s show n in Tabl e I. Ther e ar e variou s way s t o ran k naphtha s a s t o reformability . On e metho d use s a valu e determine d b y addin g th e naphthen e conten t t o twic e th e aromati c conten t an d is referre d t o a s th e N + 2A value . As show n in Tabl e I, N + 2A value s of 6 0 - 6 5 ar e typica l of naphtha s fro m midcontinen t crudes , whic h ar e simila r t o th e naphtha s bein g ru n b y U.S. refiner s unti l th e lat e 1970s. As refiner s move d int o th e 1980s an d wer e require d t o us e less desirabl e crudes , th e N + 2A value s of naphtha s bega n t o decrease . Naphtha s simila r t o th e light Arabia n naphth a in N + 2 A valu e ar e becomin g mor e c o m m o n . Tabl e II illustrate s th e difference s in propertie s betwee n feed an d produc t for a midcontinent-typ e naphth a reforme d t o 95 clear researc h octan e numbe r (RONC) . As ca n b e seen , ther e is a substantia l increas e in th e aromati c conten t of th e reformat e at th e expens e of th e napthenes . Sinc e th e paraffin s ar e difficul t t o reform , ther e is onl y a smal l reductio n in th e paraffi n conten t betwee n feed an d reformate .

TABLE I Feedstock Examples Naphtha properties

West Coast naphtha

Gravity (°API) ASTM D-86 distillation (°F) IBP 10% 30% 50% 70% 90% EP Composition (vol%) Paraffins Naphthenes Aromatics

49

56

66

230 272 287 302 326 350 385

178 221 235 246 264 289 330

176 194 205 216 232 254 298

22 56 22 100 100

45 45 10 100 65

74 19 7 100 33

N + 2A

Midcontinent naphtha

Light Arabian naphtha

5

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Catalytic Reforming of Naphtha in Petroleum Refineries TABLE II Feed versu s Reformat e Compariso n for Midcontinen t Naphth a Properties

Feed

Reformate

Gravity (°API) ASTM D-86 distillation (°F) IBP 10% 30% 50% 70% 90% EP Volumetric average boiling point (°F) Reid vapor pressure (psia) Composition (vol%) Parafins Naphthenes Aromatics

56

46

178 221 235 246 264 289 330

140 200 230 245 270 305 365

251 1.0

250 2.5

45 45 10 100 55

40 5 55 100 95

Octane (RONC)

III . A.

Proces s Descriptio n U N I T CLASSIFICATIO N

Reformin g unit s ar e usuall y classified a s belongin g t o on e of th e followin g thre e categories : semiregenerative , cyclic, or movin g bed . Thes e classifica › tion s reflec t th e manne r an d frequenc y of regeneratio n of th e reformin g catalyst . 1.

Semiregenerative

Units

A bloc k diagra m of a semiregenerativ e uni t is show n in Fig. 2. Pretreate d feed enter s th e uni t throug h a feed-effluen t hea t exchanger . Fro m th e exchanger , th e feed, combine d wit h th e recycl e gas, goes t o a heate r t o increas e its temperatur e t o a rang e of abou t 9 0 0 - 9 8 0 T . Th e feed the n goes throug h thre e or fou r reactor s in series . Sinc e th e reformin g reaction s affectin g th e reacto r temperatur e ar e mainl y endothermic , ther e ar e fur › nace s betwee n th e reactor s t o restor e th e temperatur e of th e feed strea m t o th e desire d level. Th e produc t fro m th e last reacto r goes throug h th e feed-effluen t hea t exchange r an d the n t o a flash drum . At th e flash drum ,

M. Dean Edgar

130

NE T HYDROGE N TO REFINER Y

RECYCL E GA S

n

Pk fs Rx 2

Rx 3

La! UJ FLAS H DRUM NAPHTH A REFORMAT E TO STABILIZE R

Fig. 2. Example of a semiregenerative reforming unit. Rx, Reactor.

th e liqui d produc t is take n off th e botto m of th e vessel, an d th e overhea d is divide d int o produc t hydroge n an d recycl e gas. Th e liqui d produc t is sen t t o a stabilize r t o remov e th e light ends . Thes e unit s hav e t o shu t dow n periodicall y t o regenerat e th e catalys t whic h become s deactivate d a s a resul t of cok e deposition . Generall y refiner s striv e t o operat e thes e unit s t o achiev e at least a 6-mont h tim e interva l betwee n regenerations . Thi s limit s th e m a x i m u m produc t octan e valu e obtaine d t o abou t 100 R O N C . 2.

Cyclic

Units

Figur e 3 is a n exampl e of a cyclic unit . It differ s fro m a semiregenerativ e uni t in tha t it ha s on e additiona l reacto r an d a manifol d syste m tha t allow s th e catalys t in on e reacto r t o b e regenerate d whil e th e catalys t in th e othe r reactor s is processin g feed . Th e additiona l reacto r is know n a s a swin g reactor . Th e swin g reacto r ca n b e substitute d for an y of th e serie s reactors . Thi s desig n ca n tak e advantag e of low uni t pressure s t o gai n a highe r C 5+ reformat e yield an d hydroge n mak e an d ca n b e operate d at hig h octan e levels ( 1 0 0 + R O N C ) tha t woul d resul t in unacceptabl y shor t cycle length s in

5

Catalytic Reforming of Naphtha in Petroleum Refineries

RECYCL E GA S

131 NET HYDROGE N TO REFINER Y ,

Tram •A FLAS H DRUM

-REGENERATIO N PIPIN G

I REFORMAT E ^ TO STABILIZE R

Fig. 3. Example of a cyclic reforming unit. Rx, Reactor.

a semiregenerativ e unit . Cok e mak e is usuall y highe r in th e last reactor s of a reforme r becaus e of thei r highe r averag e be d temperature . Therefor e reac › tor s in thes e position s in a cyclic uni t ar e normall y regenerate d mor e frequentl y tha n th e firs t reactors . In additio n t o allowin g operatio n at highe r severit y or lower pressur e tha n in a semiregenerativ e unit , th e mor e frequen t regeneratio n of individua l reactor s in a cyclic uni t result s in less of a declin e in C 5+ reformat e yield an d hydroge n productio n wit h tim e on strea m whe n compare d t o th e yield s of a semiregenerativ e unit . 3.

Moving Bed

Units

A movin g be d unit , show n in Fig. 4, is a n extensio n of th e cyclic uni t concept . I n th e mos t c o m m o n movin g be d design , th e reactor s ar e stacke d on e ato p th e othe r excep t for th e fourt h (last ) reacto r whic h frequentl y is set besid e th e othe r stacke d reactors . Th e flow pat h of th e feed is simila r t o tha t of th e othe r reforme r design s in tha t th e feed is heate d b y exchange , heate r an d interheaters , an d exit s throug h a flash dru m en rout e t o a stabilizer . Th e reactor s ar e radia l flow in design . Th e catalys t is slowly move d fro m th e firs t

M. Dean Edgar

132

RECYCL E GAS

NET HYDROGE N TO REFINER Y ^

CATALYST CIRCUI T

Fig. 4. Example of a moving bed reforming unit. Rx, Reactor.

(top ) reacto r t o th e botto m reactor . Th e coke d catalys t is sen t t o th e regeneratio n section . Catalys t flows throug h th e fourt h reactor , if present , a s a separat e system . Thes e unit s ca n b e buil t withou t th e regeneratio n sectio n an d operate d as semiregenerativ e units . A less c o m m o n desig n for a movin g be d reforme r ha s th e individua l reactor s place d separately , a s in typica l semiregenerativ e fashion , wit h provision s for movin g th e catalys t fro m th e botto m of on e reacto r t o th e to p of th e nex t reacto r in line . Coke d catalys t is withdraw n fro m th e last reacto r an d sen t t o a regeneratio n vessel. Fres h catalys t or regenerate d catalys t is adde d t o th e to p of th e first reacto r t o maintai n a constan t quantit y of catalyst . Becaus e ther e is a mechanis m tha t prevent s excessive cok e buildu p on th e catalyst , thes e unit s ca n operat e at low pressures , e.g., 100 psig, and/o r hig h severity , e.g., 100+ R O N C . Reformat e yield loss or hydroge n productio n declin e over tim e on strea m is minimize d b y selectin g th e correc t catalys t circulatio n rate .

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Catalytic Reforming of Naphtha in Petroleum Refineries

B.

PURPOS E O F REFORMIN G

Reformer s ar e operate d primaril y t o produc e eithe r moto r fuel or aro › matics . In eithe r case , hydroge n is produce d tha t ca n b e use d in othe r refiner y unit s suc h a s hydrotreaters . Moto r fuel productio n usuall y use s a full rang e or heav y naphth a wit h a n en d boilin g poin t of abou t 400 F . Th e octan e of thi s naphth a is increase d t o a 95 - 1 0 2 R O N C range . Thi s materia l provide s th e high-octan e componen t tha t is blende d wit h othe r refiner y stream s boilin g in th e gasolin e rang e t o produc e th e finished product . Fo r aromatic s production , frequentl y a light naphth a feed wit h a n en d boilin g poin t of 310 t o 340 F is sen t t o th e reformer . Thes e unit s ar e calle d B T X unit s (for benzene , toluene , xylene) . Th e produc t fro m th e B T X reforme r is normall y sen t t o a n aromatic s extractio n unit . Th e benzene , toluene , an d xylene s ar e utilize d a s ra w material s for variou s petrochemica l processes .

C.

U N I T D E S I G N VARIABLE S

Regardles s of it s classificatio n whe n a reformin g uni t is in th e desig n stage , ther e ar e severa l variable s t o consider : liqui d hourl y spac e velocit y (LHSV) ; pressure ; hydrogen-to-hydrocarbo n mola r rati o ( H 2: HC) ; an d type , num › ber , an d size of reactors . 1.

Liquid Hourly Space

Velocity

_ 1

C o m m o n value s of L H S V rang e fro m 1 t o 3 h r . Fo r a uni t of a specified capacity , selectio n of th e LHS V determine s th e volum e of catalys t required . Excep t for ver y low value s of spac e velocity , e.g., less tha n 1.0, whic h ten d t o favor hydrocracking , th e magnitud e of th e spac e velocity’ s effect on yield selectivit y is considerabl y less tha n th e effect of othe r desig n variables . Th e primar y effect of spac e velocit y on uni t operatio n is th e cycle lengt h obtaine d an d th e start-of-ru n temperatur e require d t o achiev e a desire d produc t octan e level. 2.

Pressure

Th e pressur e selecte d for th e operatio n ha s a majo r effect on yield an d cycle length . Pressure s o f 4 0 0 t o 500 psi g favo r a lon g cycle length . However , b y reducin g th e pressure , th e dehydrogenatio n reactio n equilibriu m is shifte d in a directio n favorin g increase d aromatic s yield an d hydroge n

134

M. Dean Edgar

production . Th e decreas e in uni t pressur e reduce s th e likelihoo d of hydro › cracking . Enhancin g th e dehydrogenatio n reaction s an d inhibitin g th e hy › drocrackin g reaction s resul t in a n increas e in th e C 5+ reformat e yield . Unfortunatel y a reductio n in pressur e als o increase s th e rat e of cok e deposi › tio n on th e catalyst , whic h reduce s th e cycle length . Semiregenerativ e unit s coul d no t tak e advantag e of low-pressur e operatio n unti l th e adven t of bimetalli c an d multimetalli c reformin g catalyst s wit h thei r abilit y t o tolerat e highe r cok e levels. Pressure s of 200 t o 250 psig coul d the n b e use d whil e obtainin g cycle length s simila r t o thos e obtaine d wit h straigh t platinu m catalyst s at highe r pressures . Wit h cyclic unit s an d movin g be d units , pressure s a s low a s 85 psi g hav e bee n considere d [4]. 3.

H2:HC

Molar

Ratio

Th e H 2: H C mola r rati o is a measur e of th e recycl e hydroge n flow rate . Th e tren d ha s bee n t o reduc e th e H 2: H C mola r rati o fro m 8: 1 - 1 0 : 1 t o 3: 1 - 5 : 1 . A reductio n in th e H 2: H C rati o reduce s th e compresso r need s of th e unit . However , a reductio n in th e H 2: H C rati o increase s th e rat e of cok e make , whic h reduce s th e cycle length . Fo r example , a reductio n in th e H 2: H C rati o fro m 5: 1 t o 4: 1 reduce s th e cycle lengt h b y abou t 20% , all othe r condition s remainin g th e same . A reductio n in th e H 2: H C rati o reduce s th e hydroge n partia l pressur e and , lik e a reductio n in uni t pressure , enhance s th e dehydrogenatio n reactio n an d inhibit s th e hydrocrackin g reaction . Th e magnitud e of th e effect on yield s of a reductio n in th e H 2: H C rati o fro m 8: 1 - 1 0 : 1 t o 3: 1 - 5: 1 is no t a s grea t a s th e effect of a reductio n in pressur e fro m 500 t o 200 psig. 4.

Reactors—

Type, Number and

Size

Th e reactor s use d in reformin g unit s ar e classified a s eithe r downflo w or radia l flow. Thes e type s of reactor s ar e illustrate d in Fig. 5. I n a downflo w reactor , feed enter s at th e to p an d flows downwar d throug h th e catalys t bed . Th e produc t exit s at th e bottom . I n a radia l flow reactor , feed enter s at th e to p an d produc t exit s at th e bottom , bu t th e feed flows acros s a n annula r catalys t be d t o a cente r pipe . Reformin g unit s ar e comprise d of all radia l flow reactors , all downflo w reactors , or a combinatio n of both . Th e advantage s of th e radia l flow reacto r desig n ar e a low pressur e dro p acros s th e reacto r an d a low foulin g rat e of th e catalys t be d du e t o particu › late s carrie d in wit h th e feed. In unit s in whic h ther e is concer n abou t scale carryove r int o th e firs t reactor , th e firs t reacto r is frequentl y of th e radia l flow design . Th e disadvantage s of th e radia l flow reacto r desig n ar e problem s in characterizin g th e flow patter n of th e feed t o obtai n full utilizatio n of th e catalys t an d possibl e be d settlin g whic h allow s feed t o bypas s th e catalyst .

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Catalytic Reforming of Naphtha in Petroleum Refineries

135

FEED

FEED

PRODUCT

PRODUCT

DOWNFLOW REACTOR

RADIAL FLOW REACTOR

Fig. 5. Examples of reactor design.

Semiregenerativ e reformer s ar e generall y buil t wit h thre e t o fou r reactor s in series . Initiall y th e unit s ha d thre e reactors . A fourt h reacto r wa s adde d t o som e unit s t o allo w a n increas e in eithe r severit y or throughpu t whil e maintainin g th e sam e cycle length . At constan t octan e operatio n ther e doe s no t appea r t o b e a significan t yield advantag e for four reactor s over thre e reactor s whe n th e additio n of th e fourt h reacto r reduce s th e spac e veloc› it y [6]. Th e size of th e first reactor s is smal l compare d t o th e size of th e en d reactors . Th e ver y rapi d endothermi c reaction s tha t occu r in th e first reactor s ca n reduc e th e reacto r temperatur e t o a poin t at whic h furthe r reactio n stops . Additiona l catalys t in th e be d belo w thi s poin t is no t utilize d effectively. Thu s th e smal l reacto r size is favored . Th e reaction s tha t occu r in th e last reactor s ar e relativel y slow, so larg e reactor s ar e required . A typica l reacto r size distributio n is show n in th e accompanyin g tabulation .

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Three-reactor system

Four-reactor system

Reactor no.

Percent

Reactor no.

Percent

1 2 3

20 30 50

1 2 3 4

12 20 28 40

Sometimes , whe n a semiregenerativ e reformer’ s capacit y is expanded , tw o existin g reactor s ar e place d in parallel , an d a new , usuall y smaller , reacto r is added . Frequentl y th e paralle l reactor s ar e place d in th e termina l position . Whe n evaluatin g uni t performance , th e paralle l reactor s ar e treate d a s thoug h the y ar e a singl e reacto r of equivalen t volume . Cycli c unit s typicall y us e five or six reactor s includin g th e swin g reactor . Movin g be d unit s generall y us e thre e or four reactors . Th e precedin g item s ar e set at th e tim e of uni t desig n a s a serie s of compromise s tha t enabl e th e uni t t o handl e th e rang e of feedstoc k type s envisage d t o produc e th e desire d produc t qualit y an d t o provid e a n accept › abl e cycle length .

IV. A.

Catalyst s

DUAL FUNCTIO N

Moder n reformin g catalyst s ar e dua l functio n catalysts . On e functio n is th e hydrogenation-dehydrogenatio n function , an d th e othe r is th e aci d function . As indicate d in Sectio n II.B , som e reaction s requir e on e or bot h of thes e functions . Fo r th e catalys t t o provid e th e desire d product s unde r reformin g conditions , th e tw o catalys t function s mus t b e balanced . Th e balanc e mus t b e maintaine d throughou t th e cycle t o obtai n th e best per › formance . Th e severit y of th e operatio n affect s th e desire d balanc e betwee n th e tw o functions . B.

SUBSTRAT E FOR M

Reformin g catalyst s hav e evolved over th e year s alon g wit h advance s in th e proces s itself. Th e mos t c o m m o n l y use d reformin g catalyst s toda y consis t of on e or mor e preciou s metal s on a n alumin a support . Th e alumin a use d for suppor t is on e of tw o crystallin e forms et a or gamma . Th e et a

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for m ha s a highe r aci d functio n tha n th e gamm a form . Th e eta for m ha s serve d a s th e suppor t for mainl y straigh t platinu m catalysts . ^-Alumin a is characterize d b y a hig h initia l surfac e area . Followin g us e an d regeneration , th e surfac e are a begin s t o decline . Tota l life is limite d b y thi s loss in surfac e are a t o jus t a few cycles. y-Alumin a doe s no t hav e a s hig h a n aci d functio n a s ^/-alumina , bu t it is mor e thermall y stabl e an d retain s mor e of its initia l surfac e throug h re › peate d us e an d regeneratio n tha n th e et a form . y-Alumina-base d catalyst s use d in cyclic reformer s ma y underg o severa l hundre d regeneration s befor e losin g enoug h surfac e are a t o requir e replacement . Th e lower aci d functio n of th e y-alumin a catalys t ca n b e compensate d for b y prope r adjustmen t of th e haloge n conten t of th e catalyst .

C.

PHYSICA L PROPERTIE S

Th e physica l propertie s of reformin g catalysts , regardles s of manufac › turer , ten d towar d simila r ranges .2Th e surfac e are a measure d b y instrumen › ta l analysi s is abou t 1 7 5 - 3 030 m / g . Th e por e volum e measure d b y wate r range s fro m 0.45 t o 0.65 c m / g . Th e catalys t particle s ar e eithe r extrudate s or sphere s rangin g fro m TV t o ^ in . diameter . Th e typica l crus h strengt h of thes e catalyst s is abou t 3 - 7 l b3 / m m . Th e densit y of reformin g catalyst s range s fro m 32 t o 49 lb/ft .

D.

P R O M O T E R METAL S

Variou s promote r metal s hav e bee n use d in reformin g catalysts . Platinu m gaine d widesprea d us e in th e 1950s an d earl y 1960s. In th e middl e t o lat e 1960s, th e additio n of rheniu m t o platinum-containin g catalyst s bega n t o see commercia l use . Th e functio n of rheniu m an d its for m on a catalys t ar e still debated , bu t it appear s t o increas e th e catalyst’ s cok e tolerance . Thi s allow s refiner s t o reduc e th e pressur e or H 2: H C rati o or t o increas e operat › in g severit y whil e maintainin g th e sam e cycle lengt h obtaine d wit h th e straigh t platinum-promote d catalyst . Th e paten t literatur e contain s reference s t o th e incorporatio n of othe r metals , suc h a s tin , germanium , an d lead , ont o th e platinum-supporte d catalyst . Catalyst s containin g platinu m an d incorporatin g othe r metal s hav e bee n use d in commercia l units . Promote r metal s usuall y compris e 1 wt /o or less of th e finished reformin g catalyst .

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E.

BENEFIT S O F BIMETALLI C CATALYST S

Th e benefit s of bimetalli c catalyst s ar e primaril y relate d t o thei r enhance d cok e tolerance . Bimetalli c reformin g catalyst s allo w operation s t o continu e unti l levels of 20 wt% cok e on th e catalys t in th e last reacto r ar e reached . Th e bette r cok e toleranc e ha s bee n use d mainl y t o allo w a reductio n in reacto r pressur e an d thereb y t o obtai n a yield advantag e for ne w an d revampe d units . Othe r way s t o tak e advantag e of th e enhance d cok e toleranc e ar e a reduce d H 2: H C ratio , increase d feed rate , increase d produc t octan e or aromatic s yield , an d reduce d catalys t volum e in ne w uni t designs .

F.

START-U P A N D PRESULFIDIN G

Fres h reformin g catalys t readil y pick s u p moisture . Thi s moistur e ca n b e fro m rai n or hig h humidit y presen t durin g loadin g or fro m wate r tha t ha s collecte d in th e low spot s in th e unit . Becaus e wate r remove s som e of th e chlorid e place d on th e catalyst , thi s coul d upse t th e balanc e betwee n th e preciou s meta l functio n an d th e aci d function . Therefor e car e shoul d b e take n t o dr y th e unit , drai n low spots , an d loa d th e catalys t unde r condition s a s dr y a s possible . Reformin g catalyst s ar e supplie d wit h preciou s metal s eithe r in th e oxid e for m or alread y in th e reduce d an d presulfide d form . If th e catalys t is in th e oxid e form , it mus t b e reduce d and , especiall y for bimetalli c catalysts , presulfide d befor e feed is introduce d int o th e reforme r unit . After loadin g th e catalyst , th e uni t is closed an d pressure-teste d wit h a nitroge n atmosphere . After determinin g tha t th e uni t is leak-free , heatin g of th e catalys t begins . A good practic e is t o maintai n a temperatur e differentia l amon g th e reactor s of 50 F betwee n th e first an d last reactor , wit h th e last reacto r havin g th e highes t temperature . Thi s prevent s condensatio n of an y wate r drive n off durin g heatin g fro m occurrin g in a downstrea m reactor . Whe n th e reacto r temperature s lin e ou t at abou t 7 0 0 - 8 0 0 T , reductio n of catalys t in th e oxid e for m take s plac e whe n th e nitroge n atmospher e is displace d b y high-purit y hydrogen . Th e catalyst , especiall y a bimetalli c catalyst , is in a highl y reactiv e stat e followin g reduction . If feed is introduce d at thi s point , methanatio n reac › tion s an d excessive hydrocrackin g ar e likel y t o occur . Thes e reaction s ar e exothermi c an d coul d resul t in a temperatur e runaway , wit h subsequen t damag e t o th e catalyst . T o preven t thi s fro m occurring , th e refine r shoul d temporaril y deactivat e th e catalyst ; mos t c o m m o n l y thi s is don e wit h sulfur . Followin g reduction , a typica l procedur e migh t call for 0.06 wt% (base d on th e weigh t of th e catalyst ) sulfu r t o b e introduce d int o th e individua l

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reactors . Hydroge n sulfid e ha s bee n widel y use d for thi s purpose , althoug h som e refiner s prefe r dimethy l sulfid e becaus e of easier handling . After injectio n of th e sulfidin g agen t is completed , th e recycl e gas strea m is checke d for th e presenc e of H 2S a s a positiv e indicatio n tha t th e sulfu r ha s worke d its wa y throug h th e catalys t beds . Onc e th e catalys t ha s bee n reduce d an d presulfided , it is read y t o accep t th e feed. Sweet naphth a is introduce d at reacto r be d temperature s rangin g fro m 700 t o 8 5 0 T . Use of th e lower temperatur e whe n feed is introduce d ma y requir e mor e tim e t o achiev e th e desire d octane , bu t it reduce s th e chance s of a temperatur e runaway . Ther e is als o som e latitud e in th e reacto r temperature s whe n feed is introduced , dependin g on th e naphtheni c con › ten t of th e feed. Th e larg e temperatur e dro p du e t o a hig h naphtheni c conten t ma y allo w highe r reacto r temperature s t o b e used . Onc e th e feed is introduce d an d ther e is n o evidenc e of temperatur e instability , reacto r temperature s ma y b e raise d slowly, abou t 25 F/hr . Ther e ar e variou s plateau s at whic h furthe r increase s in th e reacto r temperatur e ar e hel d u p unti l moistur e an d sulfu r in th e recycl e ga s com e dow n t o specified levels. Withi n thes e limitations , th e reacto r temperatur e is raise d t o th e level require d t o provid e th e desire d produc t octan e or aromatic s yield .

G.

POISON S

Durin g processing , upset s in th e feed pretreatmen t or contaminant s in th e feed ca n poiso n th e reformin g catalyst . Th e effect of thes e poison s ca n b e eithe r temporar y or permanent . C o m m o n temporar y poison s includ e sul› fur , nitrogen , an d chloride . Permanen t poison s typicall y encountere d ar e lea d an d arsenic . Tabl e II I is a list of c o m m o n reforme r catalys t poisons . 1.

Sulfur

Sulfu r poisonin g usuall y result s fro m a n upse t in or failur e of th e reforme r feed pretreatmen t system . Althoug h straigh t platinu m reformin g catalyst s ca n tolerat e a few part s pe r millio n of sulfu r in th e feed on a stead y basis ,

TABL E II I Commo n Reforme r Catalys t Poison s Temporary poisons Sulfur Nitrogen Chloride

Permanent poisons Lead Arsenic

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bimetalli c reformin g catalyst s generall y requir e less tha n 1 pp m an d som e requir e less tha n 0.2 ppm . Sulfu r poisonin g affect s th e hydrogenation - de › hydrogenatio n functio n of th e catalyst . Indication s of sulfu r poisonin g in a uni t ar e loss of C 5+ an d H 2 yields ; loss of activit y (a nee d t o increas e th e reacto r inle t temperatur e rapidl y t o maintai n produc t octane) ; reductio n in th e magnitud e of th e temperatur e dro p acros s th e unit , especiall y in th e firs t reactor ; an d th e presenc e of H 2S in th e recycl e gas . Th e activit y loss du e t o sulfu r poisonin g ca n b e recovere d b y removin g th e sulfu r fro m th e feed an d continuin g t o proces s th e feed. Unfortunately , whe n sulfu r poisonin g occurs , it is no t alway s recognize d as such . Conse › quentl y th e refine r rapidl y increase s th e reacto r temperatur e t o offset th e loss of activity . Th e reacto r temperatur e increas e produce s a mor e rapi d cok e deposition . Thi s cok e depositio n result s in a shortene d cycle lengt h even if th e sourc e of th e sulfu r is remove d at a late r time . Whe n sulfu r poisonin g occur s an d is correctl y identified , a good practic e is t o reduc e th e reacto r inle t temperatur e t o abou t 900 F . Thi s reductio n in operatin g severit y avoid s additional , unnecessar y cok e deposition . Th e sourc e or caus e of th e sulfu r contaminatio n shoul d b e locate d an d correcte d quickl y t o minimiz e th e lengt h of tim e th e uni t is ru n at reduce d severity . After th e feed sulfu r specificatio n is onc e agai n achieved , th e uni t shoul d b e kep t at th e reduce d severit y unti l th e quantit y of H 2S in th e recycl e gas drop s t o a few part s per million . At tha t point , th e uni t ca n b e returne d t o th e desire d operatin g severit y wit h little , if any , loss of cycle lengt h or activity . Onc e th e sourc e of sulfu r contaminatio n ha s bee n eliminated , a quicke r wa y t o reduc e th e sulfu r on th e reformin g catalys t tha n low-severit y feed processin g is t o stri p wit h hydrogen . Fee d is cu t ou t of th e unit , an d reacto r temperature s ar e raise d t o th e 950 - 975 F range . Hydroge n is swep t throug h th e catalys t bed s for a perio d of time , usuall y no t exceedin g 24 hr . Th e H 2S in th e off-gas is monitore d t o determin e whe n t o en d th e hydroge n sweep .

2.

Nitrogen

Nitroge n poisonin g is als o associate d wit h a n upse t in or failur e of th e feed pretreatment . Nitroge n poisonin g affect s th e aci d functio n of th e catalys t b y formin g ammoni a an d neutralizin g catalys t acidit y an d b y formin g am › m o n i u m chlorid e an d strippin g chlorid e off th e catalyst . Th e frequenc y of nitroge n poisonin g is less tha n tha t of sulfu r poisoning , sinc e feeds ar e usuall y low in nitroge n conten t even befor e feed pretreatment . Notabl e exception s t o thi s ar e naphtha s fro m Wes t Coas t crude s an d fro m syntheti c crude s suc h as shal e oil. Mos t catalyst s requir e feed nitroge n levels of less tha n 1 pp m t o less tha n 0.5 ppm . Indicator s of nitroge n poisonin g ar e a shift in catalys t selectivit y t o lower

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quantitie s of th e iso for m of th e hydrocarbon s a s th e resul t of a reductio n in isomerizatio n reaction s an d increase d difficult y in maintainin g produc t octan e bu t withou t muc h loss in th e temperatur e dro p acros s th e unit . If thi s conditio n continue s for a sufficien t perio d of time , a m m o n i u m chlorid e will precipitat e in th e feed-effluen t hea t exchanger s wit h a resultan t loss in efficiency. As wit h sulfu r poisoning , if th e refine r attempt s t o offset th e activit y loss b y increasin g th e reacto r temperature , th e rat e of cok e deposi › tio n will b e increased . After th e sourc e of nitroge n poisonin g is eliminated , th e nitroge n in th e reforme r is remove d b y continue d feed processing . Becaus e of th e loss of chloride , it will b e necessar y t o replenis h th e remove d chlorid e t o retur n th e catalys t t o th e desire d balanc e betwee n th e hydrogen › ation-dehydrogenatio n functio n an d th e aci d function . 3.

Chloride

Althoug h it is desirabl e t o maintai n th e chlorid e level on th e catalys t at a predetermine d value , occasionall y chlorid e enter s wit h th e crud e oil fro m certai n oil well cleanin g or enhance d recover y techniques . Whe n thi s hap › pens , th e chlorid e level on th e reformin g catalys t increase s an d unbalance s th e dua l function s of th e catalyst . Excessiv e chlorid e on th e catalys t shift s yield selectivit y an d frequentl y result s in excessive hydrocracking . Evidenc e of a hig h chlorid e level is a loss in hydroge n yield , increas e in recycl e ga s gravity , loss in C 5+ yield , reductio n in th e temperatur e dro p in th e last reacto r (sometime s even a temperatur e ris e in case s of excessive hydrocracking) , an d a n increase d level of HC 1 in th e recycl e ga s stream . Th e effects of hig h levels of chlorid e in th e feed ca n b e partiall y offset b y addin g wate r or alcoho l t o th e feed t o was h off th e excess chloride . Th e hea t exchanger s ca n b e water-washe d t o remov e th e a m m o n i u m chlorid e de › posits . 4.

Lead

Lea d poisonin g usuall y result s fro m th e feed becomin g contaminate d eithe r fro m bein g transporte d in a tanke r or barg e tha t previousl y containe d leade d gasolin e or fro m rerunnin g leade d gasolin e fro m refiner y slop . Whe n th e quantit y of lea d in th e feed exceed s th e abilit y of th e pretreate r t o remov e it, lea d enter s th e reforme r an d permanentl y deactivate s th e catalys t b y interferin g wit h th e hydrogenation-dehydrogenatio n function . Lea d poisonin g is characterize d b y a loss of activit y in th e first reactor , mos t easil y detecte d b y observin g a decreas e in th e temperatur e dro p acros s th e reactor . Althoug h eliminatin g th e caus e of th e lea d poisonin g limit s furthe r deactivation , th e activit y tha t ha s bee n lost is no t recoverable . Lea d tend s t o buil d u p on th e first catalys t it contact s befor e movin g deepe r int o

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th e catalys t be d or int o downstrea m reactors . Replacemen t of th e catalys t in th e first reacto r is frequentl y require d in case s of lea d poisoning . A guidelin e for acceptabl e feedstoc k lea d levels for bimetalli c reformin g catalyst s is less tha n 10 ppb . 5.

Arsenic

Arsenic , containe d in som e crudes , act s lik e lea d in poisonin g reformin g catalysts . Th e sensitivit y of reformin g catalyst s t o arseni c poisonin g is even greate r tha n t o lea d poisoning . A guidelin e t o acceptabl e feedstoc k qualit y for arseni c conten t is less tha n 2 ppb . H.

COKIN G

DEACTIVATIO N

As th e ru n progresses , even if n o poisonin g problem s occur , th e catalys t eventuall y loses activit y a s a resul t of cok e deposition . Th e rat e of cok e depositio n is a functio n of th e feedstoc k quality , th e operatin g severit y of th e unit , th e uni t pressure , th e LHSV , H 2: H C mola r ratio . As th e carbo n level on th e catalys t increases , reacto r inle t temperature s hav e t o b e raise d t o offset th e loss in activity . In semiregenerativ e units , carbo n characteristicall y deposit s lowest in th e first reacto r an d highes t in th e last reactor . Th e level of carbo n continue s t o buil d unti l th e en d of th e cycle is signale d b y th e unit’ s heater s inabilit y t o rais e th e reacto r inle t temperatur e furthe r at a constan t feed rate . Becaus e of change s in th e selectivit y of th e catalys t a s th e quantit y of cok e increases , som e refiner s find economi c justificatio n in endin g a cycle befor e reachin g th e unit’ s heate r limitation . Fo r th e bimetalli c reformin g catalys t in a semiregenerativ e unit , carbo n levels in th e last reacto r of a s muc h a s 20 - 25 wt% hav e bee n reporte d at th e en d of th e cycle [7]. Althoug h cycle length s ar e highl y variabl e an d depen › den t upo n mos t of th e unit’ s operatin g condition s an d feedstoc k quality , mos t refiner s wit h semiregenerativ e unit s attemp t t o achiev e m i n i m u m cycle length s of 6 month s t o a year . I n cyclic unit s an d in movin g be d units , th e carbo n levels on th e catalys t ar e hel d t o muc h lower values . I. 1.

REGENERATIO N

Purpose

Regeneratio n is th e proces s of restorin g a reformin g catalys t t o its origina l activit y b y carefull y removin g accumulate d cok e deposits . Durin g th e regeneratio n process , measure s ca n b e take n t o overcom e th e effects of

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temporar y catalys t poison s if th e uni t wa s shu t dow n befor e thei r deactivat › in g effects wer e eliminated . 2.

Procedure

A typica l regeneratio n procedur e is a multiste p oxidatio n process . After feed is remove d fro m th e unit , th e catalys t is swep t wit h recycl e ga s for a perio d of tim e t o remov e heav y hydrocarbon s left in th e unit . Th e heaters , reactors , an d recycl e ga s syste m ar e isolate d fro m th e remainde r of th e unit . Th e hydroge n atmospher e is replace d wit h a nitroge n atmosphere . At 700 t o 800 F , a smal l quantit y of oxygen is admitte d t o th e syste m t o initiat e burnin g of th e coke . Temperature s withi n th e reactor s ar e carefull y moni › tore d durin g th e bur n t o avoi d excessive temperature s whic h coul d damag e th e catalyst . Th e carbo n is remove d in a serie s of step s in whic h eithe r th e temperatur e or th e oxygen concentratio n of th e regeneratio n ga s is increase d unti l ther e is n o furthe r evidenc e of cok e combustion . Althoug h mos t refiner s introduc e oxygen int o th e firs t reacto r of a semire › generativ e uni t an d continu e th e regeneratio n sequenc e t o th e last reactor , som e refiner s introduc e oxygen t o bot h th e firs t an d last reactor s at th e sam e time . Thi s is don e a s a time-savin g step , sinc e th e larg e quantit y of cok e in th e last reacto r take s th e longes t tim e t o remove . However , th e oxygen conten t of th e ga s enterin g th e last reacto r mus t b e carefull y monitored . U p o n completio n of th e combustio n of cok e in th e reacto r ahea d of th e last reactor , th e oxygen n o longe r bein g consume d enter s th e last reactor , whic h whe n combine d wit h th e oxygen conten t alread y presen t ca n resul t in excessively hig h temperatures . Whe n th e uni t mus t b e shu t dow n for maintenance , maintenanc e is don e eithe r befor e or afte r a proo f bur n whic h is a bur n conducte d at a hig h temperatur e an d oxygen content . After maintenanc e or th e proo f burn , th e uni t is read y for reduction , presulfiding , an d th e introductio n of feed. Betwee n regeneratio n an d reduction , th e refine r ma y appl y othe r proprie › tar y procedures , a s recommende d b y eithe r th e proces s licenso r or catalys t vendor , t o ensur e good preciou s metal s dispersio n o n th e catalyst . Th e regeneratio n procedur e for cyclic an d movin g be d unit s als o follows th e proces s licensor’ s recommendations . 3.

Source of Oxygen for

Regeneration

I n semiregenerativ e units , th e oxygen use d in regeneratio n ca n c o m e fro m air or fro m liquefie d oxygen . Regeneratio n at low pressure , e.g., 100 psig, wit h air a s th e sourc e of oxygen is th e mos t c o m m o n method . By usin g pur e oxygen in a nitroge n atmosphere , a refine r ca n regenerat e

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th e catalys t at a highe r pressure . Thi s reduce s th e tim e require d for th e regeneratio n procedure , althoug h th e hazard s associate d wit h usin g pur e oxygen mus t b e carefull y considered . 4.

Ojfsite

Regeneration

As a resul t of th e succes s in othe r catalyti c processes , som e refiner s ar e investigatin g havin g thei r reformin g catalyst s regenerate d b y merchan t catalys t regeneratio n services . In thi s case , th e coke-lade n catalys t is re › move d fro m th e reactor s an d shippe d t o th e merchan t regenerator’ s plan t wher e th e carbo n is remove d b y oxidation . Th e regenerate d catalys t is the n returne d t o th e refine r for reloading . Merchan t regeneratio n ma y offer som e advantage s becaus e of th e generall y bette r temperatur e contro l afforde d as compare d wit h in situ regeneration . J.

T O T A L LIF E

If a reformin g catalys t is protecte d fro m permanen t poisons , it ca n b e returne d t o its origina l activit y b y carefu l regeneration . Wit h prope r care , in semiregenerativ e units , catalyst s hav e bee n use d for at least 10 cycles befor e bein g replaced . Bimetalli c reformin g catalyst s hav e laste d for over 10 yr of operatio n or 800 bbl/l b in som e reformer s [8]. In cyclic units , it is no t u n c o m m o n for catalyst s t o b e subjecte d t o severa l hundre d regeneration s befor e bein g replaced . Whe n it ha s bee n determine d tha t a reformin g catalys t ca n n o longe r hav e its activit y restore d b y regeneration , th e catalys t is replaced . Th e spen t catalys t is usuall y sen t t o a metal s reclaime r for recover y of th e platinu m and , if present , othe r promote r metals , especiall y rhenium . Th e recovere d preciou s metal s ar e normall y returne d t o th e catalys t manufacture r for incorporatio n int o th e nex t productio n of fres h catalys t for th e refiner .

V. A.

Operatin g Variable s

REACTO R INLE T T E M P E R A T U R E

Th e operatin g variabl e use d mos t b y uni t operator s is th e reacto r inle t temperature . Althoug h th e size of th e uni t an d th e capacit y of th e pump s an d compressor s ar e fixed at th e desig n stage , th e reacto r inle t temperatur e ca n b e varie d t o achiev e an d maintai n th e desire d produc t properties . Severa l temperatur e profile s ar e c o m m o n l y used , as diagramme d in Fig. 6.

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DESCENDING TEMPERATURE PROFILE

EQUAL TEMPERATURE PROFILE

I

I Rx1

Rx2

Rx3

Rx1

Rx2

J

Rx3

Inlet Temperatures

Inlet Temperatures

MODIFIED ASCENDING TEMPERATURE PROFILE

ASCENDING TEMPERATURE PROFILE

I

I I Rx1

I Rx2

I Rx3

I Rx4

Rx1

Rx2

Rx3

Inlet Temperatures Inlet Temperatures Fig. 6. Example of reactor inlet temperature profiles. Rx, Reactor.

Probabl y th e mos t c o m m o n profil e is on e in whic h all th e reacto r inle t temperature s ar e th e same . As th e cok e on th e catalys t build s up , th e inle t temperatur e of eac h reacto r is increase d equally . Thi s procedur e is followed unti l th e limi t of th e unit’ s abilit y t o achiev e highe r temperature s is reached . Anothe r profil e use d ha s descendin g reacto r inle t temperatures . I n thi s case , th e lea d reacto r temperatur e is highe r tha n tha t in th e nex t reacto r in lin e an d so on t o th e last reactor , whic h ha s th e lowest reacto r inle t temperature . As wit h th e even temperatur e profile , th e individua l reacto r

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inle t temperature s ar e increase d t o compensat e for catalys t deactivatio n whil e maintainin g th e descendin g profile . Th e rational e for th e descendin g profil e is t o kee p th e last reacto r temperatur e lowest t o minimiz e cok e deposition , sinc e wit h a n even temperatur e profil e carbo n depositio n is highes t in th e last reactor . Unfortunately , even wit h a 5 - 10 F differenc e in inle t temperatur e betwee n reactors , th e averag e be d temperatur e of th e last reacto r is still highe r tha n th e averag e be d temperatur e of th e lea d reactor . Thus , althoug h possibl y helpin g t o prolon g ru n length , th e averag e be d temperature , whic h affect s carbo n deposition , is still highe r in th e last reacto r an d so th e carbo n depositio n is highe r also . Anothe r profil e encountere d ha s th e firs t tw o of four reactor s hel d at start-of-ru n condition s for mos t of th e cycle. T o maintai n produc t qualit y th e reacto r inle t temperatur e of th e thir d an d fourt h reactor s is increase d as th e catalys t deactivates . Thi s profil e is typica l of unit s wit h a split recycle . In thes e unit s a smal l quantit y of recycl e ga s is fed t o th e firs t tw o reactor s wher e dehydrogenatio n reaction s normall y occur . Sinc e th e averag e be d tempera › tur e in th e firs t tw o reactor s is relativel y low, an d sinc e th e reaction s occurrin g ar e no t th e majo r contributor s t o coking , a hig h H 2: H C mola r rati o is no t require d t o minimiz e cok e formation . Recycl e ga s fro m th e firs t tw o reactor s combine s wit h additiona l recycl e ga s injecte d ahea d of th e thir d reactor . Thi s provide s a highe r H 2: H C mola r rati o in reactor s in whic h th e majo r coke-producin g reaction s occur . Eve n in unit s withou t a split recycle , occasionall y a n ascendin g reacto r inle t temperatur e profil e is tried . Base d on th e consideratio n of lon g cycle length , th e ascendin g profil e appear s t o b e th e least desirable , sinc e it aggravate s th e alread y hig h averag e be d temperatur e in th e last reactor .

B.

FEEDSTOC K E N D POIN T

Th e refine r ma y hav e som e flexibilit y in th e choic e of feedstoc k en d boilin g point . Th e en d boilin g poin t of th e feed is somewha t set b y th e purpos e for whic h it is reformed . Sinc e th e reformat e ha s a n en d boilin g poin t abou t 1 5 - 2 0 F highe r tha n tha t of th e feed , th e moto r fuel reforme r feed is usuall y limite d t o a m a x i m u m en d poin t of 400 F t o mee t gasolin e boilin g specifications . Feed s t o B T X reformer s generall y hav e a lower en d boilin g poin t of abou t 300 F . I n th e cas e of moto r fuels reformer s in whic h th e feed en d boilin g poin t ma y exceed 400 F , th e portio n of th e feed tha t boil s abov e 400 F greatl y increase s cokin g of th e catalyst . Becaus e of th e rapi d cokin g tendency , mos t refiner s d o no t proces s feed wit h a n en d boilin g poin t exceedin g 400 F .

5

Catalytic Reforming of Naphtha in Petroleum Refineries

C.

WATER-CHLORID E

147

BALANC E

A n operatin g variabl e tha t ha s a majo r impac t on th e performanc e of th e reformin g catalys t is th e moistur e level in th e unit . After pretreating , th e naphth a enterin g th e reforme r typicall y contain s 3 - 5 pp m b y weigh t water . I n th e reforme r th e wate r in th e feed vaporize s an d interact s wit h th e chlorid e on th e reformin g catalys t an d will stri p th e chlorid e fro m th e catalys t if given enoug h time . If th e wate r level in th e feed is ver y high , e.g., 50 ppm , it ca n remov e chlorid e fro m th e catalys t quickl y an d upse t th e dua l functio n balanc e of th e catalyst . T o monito r th e quantit y of moistur e in th e unit , refiner s us e recycl e gas moistur e analyzers . Wit h informatio n fro m thes e moistur e analyzers , refiner s ca n ad d chlorid e t o th e reforme r feed at levels prescribe d b y th e proces s licenso r or th e catalys t vendo r t o maintai n th e prope r m e t a l - a c i d functio n balance . Som e proces s licensor s an d catalys t vendor s feel tha t thei r catalyst s wor k bes t in a n essentiall y moisture-fre e atmospher e wit h 1 pp m b y volum e wate r or less in th e recycl e gas. Wher e grea t effort s ar e mad e t o avoi d contactin g thes e catalyst s wit h water , chlorid e additio n is no t required . Becaus e moistur e displace s chlorid e fro m th e catalyst , man y refiner s tende d t o slu g chlorid e int o th e uni t whe n it neare d its end-of-ru n tempera › tures . I n som e case s thi s increase d th e produc t octan e an d allowe d th e cycle t o b e stretche d for severa l week s befor e regeneratio n wa s required . I n som e case s th e slu g of chlorid e wa s to o muc h an d it upse t th e operation . In general , optimu m performanc e is achieve d b y addin g smal l quantitie s of chlorid e continuousl y durin g th e ru n t o establis h a water-chlorid e balanc e tha t maintain s th e desire d chlorid e level on th e catalyst . Thes e balance s ar e usuall y specific for eac h catalys t typ e an d considere d proprietar y informa › tio n b y th e proces s licenso r or catalys t vendor .

VI.

Futur e Reformin g Growt h

Althoug h catalyti c reformin g is a firmly establishe d refiner y proces s tha t will continu e t o b e used , its additiona l growt h in capacit y is tie d t o gasolin e consumption , automobil e engin e design , an d petrochemical s demands . Th e switc h t o lead-fre e gasolin e tha t starte d in th e mid-1970s , necessitate d b y governmen t regulatio n of automobil e exhaus t emissions , straine d refiners ’ abilit y t o provid e enoug h high-octan e blendin g component s for thei r gaso › lin e pools . Reformer s wer e ru n t o highe r octan e levels an d in som e case s expande d in anticipatio n of futur e demands .

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Currently , however , becaus e of a combinatio n of th e e c o n o m i c pressure s of hig h gasolin e cost s an d automobil e manufacturers ’ improvement s in automobil e mileage , gasolin e consumptio n ha s droppe d considerabl y fro m previousl y projecte d values . Automobil e maker s ar e committe d t o manu › factur e even smaller , mor e fuel-efficien t automobile s in th e future . Conse › quently , it appear s tha t th e nee d for expansio n of reforme r capacit y will b e limite d for at leas t th e nex t severa l years .

Reference s 1. W. L. Nelson, "Petroleum Refinery Engineering," 4th edition, p. 810. McGraw-Hill, New York (1958). 2. F. G. Ciapetta, and D. N. Wallace, Cat. Rev. 5(1), 67-158 (1971). 3. L. R. Aalund, Oil Gas J. 79(13), 6 3 - 6 5 (1981). 4. B. C. Gates, J. R. Katzer, and G. C. A. Schuit, "Chemistry of Catalytic Processes." McGraw-Hill, New York (1979). 5. W. H. Hatch, S. J. Cohen, and R. Diener, Modern Catalytic Reformer Designs Help Reduce Cost of Low-Lead Gasoline. Presented at NPRA Annual Meeting, A M - 7 3 - 3 3 (1973). 6. NPRA Q&A Session on Refining and Petrochemical Technology, pp. 2 8 - 2 9 . Petroleum Publishing, Tulsa (1976). 7. NPRA Q&A Session on Refining and Petrochemical Technology, p. 90. Farrar and Asso­ ciates, Tulsa (1977). 8. NPRA Q&A Session on Refining and Petrochemical Technology, pp. 2 3 - 2 4 . Petroleum Publishing, Tulsa (1976).