Use of Zeolite Containing Catalysts in Hydrdcracking

93

B. Imelik et al. (Editors), Catalysis by Zeolites © 1980 Elsevier Scientific Publishing Company, Amsterdam - Printed in The Netherlands

USE OF ZEOLITE CONTAINING CATALYSTS IN HYOROCRACKING. C. MARCILLY. J.P. FRANCK INSTITUT FRANCAIS OU PETROLE - B.P. 311

92506 - RUEIL-MALMAISON CEOEX

(FRANCE)

INTRODUCTION The beginning of hydrocracking may be traced back to the year 1927 when the first BERGIUS plant for hydrogenating brown coal at

LEUNA (GERMANY) was put on stream. Hydrocracking

only appeared a few years later in the U.S.A. where this process is now more widespread than elsewhere. At the present time. hydrocracking is used for converting various petroleum cuts. from light naphtas to atmospheric or deasphalted vacuum residua. to lighter and more valuable products like propane. butane. gasoline. jet fuels, diesel or heating oils and, in some cases, lubricating oils (1). fuel production.

It has essentially been developed in the U.S.A. for motor

In other countries. hydrocracking is not as common and is directed rather

to the production of middle distillates (jet fuels and gas oils). The first zeolite based catalysts were used in 1964.

Compared with conventional amorphous

catalysts, their advantage is especially apparent when a maximum gasoline is aimed for. Hydrocracking is now the second largest use for zeolite containing catalysts. I. THE CONVENTIONAL HYDROCRACKING ON AMORPHOUS CATALYSTS. According to the required products and the treated feedstock. the process scheme involves one or two stages (2, 3).

The first one essentially frees the feedstock from the nitrogen

and sulfur compounds as NH and SH and hydrogenates most of the aromatics without any 3 2, serious cracking of the hydrocarbons. In the second one, the main reactions are cracking and hydrogenation of the effluents coming out from the first stage.

Each of them contains

a specific catalyst specially adapted to the desired transformations. 1.1

The catalysts They are of bifunctional type combining a hydrogenating and an acidic function. For the first stage, the hydrogenating function is produced from cobalt

o~nickel

sulfi-

des combined with molybdenum or tungsten sulfides. The atomic ratio of metals from the th th group to those from the VI group is near 0.25 (4). For the second stage, that func-

VIII

tion usually comes from nickel or palladium (5). The acidic function is prOVided by an acidic carrier. like alumine, silica-alumina, silica-magnesia. etc ..•

It gives the catalyst its cracking activity.

The acidity level

depends on the feedstock and the required transformation : it must be higher in the second step than in the first one. The selectiVity of the transformation largely depends on the balance between both functions.

If the hydrogenation function is high, in comparison to the acidic one, the main

reactions. besides deazotation and desulfurization. are hydrogenation and isomerization.

If

94

it is low, the main reaction is cracking.

That can be seen in figure 1 r6J which compares

selectivities of two catalysts having the same hydrogenation activity for the n heptane transformation: the most acidic catalyst, with the highest heat of ammonia sorption, has the highest cracking activity,

This explains the difference in the products distribution

obtained from the same feedstock, on the one hand by catalytic cracking, and on the other hand by hydrocracking. A1

As shown in figure 2 (7), hydrocracking of hexadecane on Pt/Si0 2gives products with carbon numbers varying between 3 and 12-13 through a maximum bet-

203 ween 4 and 11, whereas catalytic cracking on silica-alumina gives products with carbon numbers between 2 and 7 and a greater maximum at about 4-5.

The catalytic hexadecene cracking

results in a distribution of products very similar to that obtained from hexadecane hydrocracking: replacing hexadecane by hexadecene is comparable with providing an infinite hydrogenation function to the catalyst.

These results explain why catalytic cracking is

better for gasoline production, whereas hydrocracking is better for middle-distillates production, (jet fuel, diesel or heating oil). Ref. 8 gives the respective levels of hydrogenation and acidity that are to be reached for various industriel cases.

100·...-----==-=-=:-:-::==::::-:l NH3 HEAT OF ADSORPTION CAL/G CARRIER 1 ,10 CARRIER 2,16

i\

! !

90 ~

o::< 60

/

o

d >=

,

I

30

nY6 CONVERSION = 53%

\ \

CATALYTIC CRACKING OF HEXAOECANE

~ CATALYTIC CRACKIt\G

/

'){\OFHEXAO:::::RACKING OF HEXAOECANE

, _

(P1/Si02AI2~)

"...) O~_~_+

o

Fig. 1 (6) : hydrocracking of n-heptane (Ni-Mo sulfides/Si0

1.2

2-A1 203)

-----;-~=""-'!! __~

15

Fig. 2 (71 : comparison between hydrocracking and catalytic cracking

The process Hydrocracking is a very flexible process allowing the treatment of various feedstocks.

from naphtas to residua, and the acquirement of a large diversity of products.

Operating

conditions are within the following limits

temperature from 350 to 450°C, partial pressure 1• molar ratio H near 20. Com2/HC pared with catalytic cracking, the use of high partial pressures of hydrogen and of relati-

of hydrogen from 100 to 150 bar, LHSV from 0.3 to 1.5 h-

vely low temperatures decreases the kinetics of coke formation and favors the hydrogenation of aromatic compounds.

It results in a fixed bed technology, a very exothermic reaction. an

important chemical consumption of hydrogen and, contrary to catalytic cracking. nearly nonolefinic and non-aromatic products. The process outline usually requires only one stage when the desired transformation is not important : for instance, production of middle-distillates from vacuum distillates or heavy coking distillates or deasphalted vacuum residua.

It often involves two stages when

the required transformation is important : production of a maximum gasoline from vacuum

95 distillates or deasphalted vacuum residua. figurffi3 and 4 (3, 6). hydrogen recycle.

Both processes' schemes are respectively shown in

In figure 4, it can be noticed that each step possesses its own

This results from the necessity of keeping the second step catalyst's

environment free of NH 3 and SH 2 and, therefore, of getting rid of both compounds at the outlet of the first step.

SEPARATORS

SINGLE STAGE REACTOR

FRACTIONATOR

GASOLINE

JETOR

DIESEL FUEL

o--",~)"'-

. . ._ _

~....

LIQUID OR RECYCLE ADDlTIONAL

....

!~E!f.l.

Fig. 3 : single stage conventional hydrocracking (amorphous catalyst)

2nd STAGE

lsi STAGE REACTOR

FRACTIONATOR

REACTOR

SEPARA10RS

SEPARATORS

FUELGAS GASOUNE

JETOR

DIESEL FUEL

LIQUID

RECYCLE OR L.

Fig. 4

ADDITIONAL

..PRODUCT •

two-stage conventional hydrocracking (amorphous catalysts)

Table 1 (8) gives an example of results performed in a two stage industrial unit with the possibility of liquid recycle, when products heavier than those aimed at are to be avoided. The main products that can be obtained are light [3-[4 gases, light gasoline [5-80°[, (essentially [5-[6 hydrocarbons), naphta 80-185°[, jet fuel 185-230 0 [ and gas oil.

The

following points deserve to be emphasized : - the light gasoline has a high research octane number due to its high isoparaffinic character, - the heavy gasoline is a very good feedstock for catalytic reforming because of its high napht:e ne content, - the burning qualities of jet fuel and gas oil, respectively measured as smoke point and

96

diesel index, are excellent as a result of their low aromatic content, - the pour point and freezing point of jet fuel and gas oil are good owing to their low content of linear paraffinic compounds. TABLE 1 (ref. 8) Yield and quality of eroducts from Ilydrocracking middle-east vacuum distillate. (d 2o ; 0,905 ASTM 36o-48o°C S ; 2,41 % N ; 900 ppm) 4

Initial point of recycled residual oil (OC) Yield (wt %) Gas ...................•.............•.. Light gasoline ......••................. Naphta .•........•......•.•.....•...•... Jet fuel ••..•....•..•••.•...•••.•••..•. Gas oil .•.....•....................•.•. Residuum •..........................•.•.

280

200

350

Without recycle

Gasoline

Jet fuel

Gas oil

Gas oil

16.4 22.0 64.9

10.7 10.3 21.7 60.3

7.7 6.3 13.4

6.7 5.6 13.0

74.8

69.0 8.0

103.0

102.2

102.3

103.3 Characteristics :

20 4 NOR NOR + 0.5 %0 20 Naphta ..••........ d 4 PNA Jet fuel •.....•.•. d 20 4 S (ppm) Smoke point (mm) Ar (% pds) Cristallisation point (OC) Gas oil ••.•....•.• d~O Light gasoline •.•• d

0.670

0.668

82 93

75 89

0.742 56/34/10

0.675 74 88

0.665 76 91

0.742

0.770

0.765

51/37/12

48/39/12

49/38/13

0.802 <:10 25 15

<-

S (ppm)

60

0.827 30 62 - 45 70

nr

Freeze point (-ci Flash point (OC)

0.814 40 70 - 33 74

II. HYORoCRACKING ON ZEOLITE BASED CATALYSTS. The development of new highly active zeolite based catalysts has resulted in the modification of the conventional process scheme.

The new process which was first used in 1964 at the

UNION OIL Co. of California refinery in Los Angeles. is essentially characterized by a first stage with two different catalysts : the first one is for classical hydrotreating. the second one contains a zeolite and promotes a substantial cracking of the feedstock.

This new design

combines. in one stage. the two stages of the conventional process without any intermediate separation.

It was proposed by UNION OIL Co. of California and Esso Research and Eng. Co.

as Unicracking-JHC integral hydrotreating-hydrocracking process (3, 9). 11.1

The catalysts

11.1.1 Composition and preparation

97

The hydrogenating function is provided by metals from the Vlllth group, particularly palth groups (10, 11). Reaching good

ladium, or by combining metals from the Vlll t h and Vl

noble metal dispersion and macroscopic distribution into the zeolite is of great importance.

);+.

This aim can be reached by ion exchanging palladium as Pd(NH Techniques for deposing 3 of non noble metals are not well known. Some of them are described in patents (12, 13, 14). The acidic function comes from a zeolite mixed with a refractory oxide like alumina, silica-alumina, silica-magnesia, etc ...

The zeolite content is from about 20 to 80 wt %.

Among the most open structures, faujasite and large pore mordenite, the former is more selective, and therefore more suitable, than the latter for producing gasoline (15, 16). That clearly appears in figure 5 (16) which compares the distribution of light products on Pd HY and Pd HMordenite and shows that the former gives more gasoline than the latter.

In-

dustrial catalysts contain Y zeolite in H form or exchanged with bivalent metals, like magnesium (17). Contrary to the first catalysts, the new ones which appeared at the end of the 60 t h (18) seem to contain a hydrothermally stabilized form of Y zeolite. The trend would now be towards increasing use of ultrastable Y (19) like Z14US (20). TABLE 2 (ref 36)

40...----------...,

Comparison between zeolite containing catalyst

o Pd/MORDENITE

at 35

and amorphous silica-alumina for maximal diesel oil production. Vol %/feed C4 CS-82°C 82-166°C 166-371°C

Cl

HZ consump-

~ 25

Zeolite containing Amorphous catacatalyst lyst 4.0 8.0 20.5 81.5

20

2.2 6.4 6.7 97.2

15 10

112:5

Total C~ han Kmol/m3

• Pd/FAUJASITE

3 30

10.15

NOR clear (C5-82°C) IJ3. U 54.0 Cetane number (166-371°C)

5

10.6B 74.0 60.0

Fig. 5 (16) : hydro cracking of an atm. Gas oil (P.I. = 1BO°C) conversion 17r

The stabilizing treatments leading to ultrastable form of Y zeolite, increase the Si0 A1

ratio of the framework and modi fie the acidity. 203 such stabilized forms may be classified in 4 groups :

The main techniques for obtaining

21

(1) - Self-steaming Me

Daniel and Maher (21) have proposed 2 techniques from NaY.

One of them involves 4

steps : 2 ion-exchanges of sodium with ammonium, each of them preceding a calcination, the first one at 540°C. the second one at B15°C under self-steaming conditions. tions the unit-cell parameter is 24.3

Ainstead

In these condi-

of 24.65 ~ for NaY. The calcination of

NH 4Y in deep-bed conditions under static atmosphere. reported by KERR (22), belongs to that group too. (2) - Controlled steaming Hansford (23) and Eberly (24, 25) describe other techniques in which NH or HY are cal4Y cined from about 370 to 800°C under dynamic atmosphere containing air and 3 to 100 vol % of water vapor.

The unit-cell parameters of the thus obtained stabilized HY is down to about

98

° 24.3 A. (3) - Thermal treatment by halogen or halogenic acid in a vapor phase In such a technique, potented by 8.A.S.F. (26), dehydrated HY is calcined under chlorine or hydrochloric acid between 400 and 700°C. (4) - Treatments in liquid phase One method consists of exchanging sodium of NaY in an autoclave containing an ammonium salt solution, the pH of which is from 2.5 to 5, at a temperature between 150 and 230°C (27).

In another method, the zeolite is refluxed in a H EDTA solution (28, 29, 3D, 31) 4

or in acetylacetone (3D, 31, 32).

11.1.2 The balance between hydrogenating and acidic functions The performances of the catalyst greatly depend on the balance between both functions. The possibility of modifying the acidity of zeolites by suitable ion exchanges or treatments, and the hydrogenating function. allows the preparation of "tailored" catalysts which are adapted to a refiner's specific need (18).

Concerning noble metal. the hydrogenating acti-

vity depends on the reduction extent. on the accessibility and the size of metallic particles in zeolite structure.

So far as it is known, no thorough study, similar to that of

GALLEZOT and al. (33, 34) on PtY. has been published, concerning the positions of palladium atoms or particles in stabilized Y structure according to the calcination and reduction conditions. Some modifications of the balance between both functions may occur during operation. They seem to be reversible when caused by the presence of poisons like NH reacting atmosphere.

or SH in the 3 2 In some cases, they may be irreversible and require. as it will be

seen further. "rejuvenation" of catalysts. 11.1.3 Advantages of zeolitic catalysts over amorphous catalysts Zeoli tic catalysts present 3 major advantages over amorphous catalysts. (1) They have a greater acidity and so a greater cracking activity. (2) They are much more selective for producing a maximum gasoline.

Fig. 6 (35) presents

the true boiling point distillation curve of products from hydrocracking a vacuum gas oil

l

it appears that a zeolitic catalyst gives a yield of about 65 % of C gasoline instead 5-180oC of a maximum of 40 % with an amorphous catalyst. As shown in table 2 (36). this advantage disappears when the production of a maximum gas oil is required and. in this particular case, the use of a zeolitic catalyst is not justified.

The use of a zeolitic catalyst may be

profitable when a great flexibility is looked for. for example with maximal gasoline production in summer and maximal gas oil production in winter. (3) They have a much better resistance to nitrogen and sulfur compounds. One of the best advantages of zeolite based catalysts compared with amorphous catalysts is their much higher resistance to nitrogen poisons.

That superiority. ascribed to the

greater acid sites number in the zeolite. is clearly apparent in figure 7 (37) which presents the evolution of n-heptane conversion as a function of the quinoline

content, for respec-

tively a zeoli tic Bnd an amorphous catalyst; in the presence of 2000 ppm of quinoline the decrease of conversion is about 20 % for the zeolitic catalyst instead af 80 % for the amorphous one.

In other respects. the greater acidity of the zeolite enhances the resistan-

ce of the hydrogenating function to poisoning by sulfur compounds (38, 39. 40). As it will be seen further. these advantages result in a large saving for the refiner who wants to produce a maximum gasoline.

99

- - - INCREASING 380OC- CONVERSION

~

400

w

a:: ;:)

ti 300 a:: ~ ::E w

~200 Q

ti

-'

1-15%HY(0.4'%Pt)in Si02

f= 100

2-Si02-AI203(o.4%PI)

-'

~

OL-

o

-=":-::-

--;::~

1000 2000 NITROGEN W PPM (from Quinoline )

Fig. 7 (37l influence of N content on hydrocracking of n-heptane (T = 400°C, P = 25 barJ

Fig. 6 (35) hydrocracking of a vacuum gas oil (P.I. = 380°C) 11.1.4 Catalyst regeneration and rejuvenation

In an industrial unit, the activity of the catalyst slowly decreases, in spite of high hydrogen partial pressures.

For palladium containing catalyst, deactivation may have

different reasons : - classical coke formation on the active surface. - aging of the hydrogenating function (41), - inhibition of the acidic function (41). Coke and simultaneously highly basic nitrogen and sulfur compounds. fixed on the surface, are burned by classical combustion in the presence of poor oxygen containing air (regeneration) • Aging of the hydrogenation function is essentially due to the metal sintering.

This

phenomenon should rapidly occur when the catalyst is put into contact with high water vapor partial pressures, at temperatures above about 250°C. either during regeneration or. accidentally, during operation (42).

Contacting the catalyst with a neutral gas. such as ni-

trogen, at high temperatures, would also result in decreasing the hydrogenation activity (43).

Different rejuvenation techniques have been devised according to the metallic parti-

cle sizes : (1) particles with diameters

«

100-200 A (macro agglomeration)

Succession of oxidation-reduction (44, 45) or sulfurization-oxidation (46) cycles, usually performed "in situ" after coke combustion, enables metal redispersion towards particles with diameters slightly smaller than 50 (2) Particles with diameters

< 50

~.

~.

The regenerated and either oxidized or sulfided catalyst is contacted with a mixture of water vapor and ammonia gas or with an aqueous ammonia solution, under 150°C, in order to fill up the micropores.

Metallic particles are dissolved as Pd (NH

grate again in the structure towards ion exchange sites (42. 47 to be performed "ex-situ" (51).

a

2+

3)4 50).

cations which miThis technique seems

100

Causes of the acidic function inhibition are less clear.

It would seem that accidents

which are responsible for metal sintering. are also responsable for the redistribution of non-reducible metallic ions (Na. Ca, Mg, etc ..• ) on acid sites; metal ions which initially occupy sites in small cavities [sodalite cage and hexagonal prism) would migrate towards the acidic sites in the supercage (42) and this would decrease the cracking activity. sides, after use. catalysts are found to contain iron scale and sulfate ions (52).

Be-

Most of

these ions and impurities can be eliminated by controlled ion exchange with ammonium ions, at room temperature, at a pH from 4.5 to 6.5.

Such a treatment which enables the elimina-

tion of 60-70 % of sodium ions. 90 % of iron

and 80 % of sulfate ions (52), would be per-

formed~x situ' (42. 47. 48. 52). catalyst with H EDTA solution or 4

Another technique consists of contacting the regenerated with an ammonium salt of H EDTA (53). 4

11.2 The process The hydrocracking process using zeolite based catalysts involves one or two stages according to the feedstock and the desired products. 11.2.1 Description of one-stage and two-stage processes As shown in figure 8 (41). the one-stage scheme combines. in one step. the two steps of the conventional process. i.e. hydrotreating and hydrocracking. without any intermediate separation of NH

3

and SH . 2

catalyst to these poisons.

That is made possible by the high resistance of the cracking In the second reactor. the conversion is about 40 to 70 %.

After separation. the unconverted feedstock is recycled to the cracking reactor.

In spite

of its high ammonia resistance. the presence of large quantities of this compound [up to about 10

4

ppm) results in a definite suppression of the cracking activity which can be com-

pensated by an increase of temperature.

In such an "acidity controlled" environment. the

total activity is limited by the cracking activity of the catalyst. lyst must exhibit a very high acidity.

Therefore. this cata-

This may be done by adjusting the zeolite content

in the catalyst (54).

SINGLE STAGE REtlCTORS

SEPARATORS

FRACTIONATOR

UNTREATED FEED Fig. 8

single stage Unicracking-JHC unit (zeolitic catalyst)

The two-stage scheme involves one more hydrocracking reactor. As shown in figure 9 (41). the total effluent from the first stage is combined with that from the second stage and

101

sent to the fractionator where the following products are separated : propane and butane, light and heavy gasolines, jet fuels and gas oils.

The heavier residual fraction is recy-

cled to extinction in the second stage at 50-80 % conversion per pass (55, 56).

The recycle

gas from the combined high pressure separator is stripped of ammonia. so that the second stage catalyst operates under an atmosphere essentially free of ammonia [55, 56 J. ; "us, the environment is essentially "hydrogenation controlling" and the catalyst must exhibit a higt hydrogenation activity.

1st STAGE REACTORS

2ndsrAGE REACTOR

Such a catalyst probably consists of palladium on Y zeolite (41).

SEPARAiORS

FRACTIONAiOR

JETOR DIESEL FUEL

~~LEOR

'"'--~L ....._ UNTREATED FEED

ADDITIONAL FURNACE i--_~~I- _ _J.:~Q.D~9"_.

Fig. 9 : two-stage Unicracking-JHC unit (zeolitic catalysts)

II.2.2 Comparison of the new hydrocracking process with the conventional process The advantages of zeolitic catalysts (seen in part II.1.3) over amorphous catalysts results in a large saving for the refiner who wants to produce a maximum gasoline.

At cons-

tant production, a conventional two-stage process gives a C gasoline with NOR of 84 5-80oC instead of 86 with a first stage integral hydrotreating-hydrocracking process. This allows the increase of the premium/regular ratio and so results in a high profit. Moreover, the first version requires one more fractionment stage and one more furnace. Thus there is important saving of money on the total investment [10 to 20 % for a 10 6 t/year capacity) and on the operating cost [6). 11.2.3 Comparison of one-stage with two-stage processes The reasons for choosing one version rather than the other are not very simple.

They

depend upon the required operating flexibility, the characteristics of the feedstock and the yields and qualities of the desired products. (1) Flexibility: the two-stage scheme offers a better flexibility than the one-stage scheme. (2) Characteristics of the feedstock: the one-stage scheme is preferally used when the feedstock is not too difficult to transform, such as those containing few highly basic nitrogen compounds and with not a too high final boiling point, or when the required transformation is not very important.

The two-stage scheme is more suitable for difficult feed-

stocks, such as highly aromatics or rich nitrogen ones (9). (3) Yields and quality of products: to produce gasoline with a high octane number, it is desirable not to hydrogenate completely the unsaturated c0mpounds, in order to keep some aromatics in the heavy gasoline.

Table 3 [18) shows that the one-stage integrated process,

102 where the hydrocracking catalyst operates under atmospheres

rich in NH

temperature, is suitable for obtaining a good quality gasoline.

3

and SH

2

and at high

In two stages, the

obtai-

ned naphta is free of aromatics, sulfur and nitrogen and is then an excellent catalytic reforming feedstock. TABLE 3 (ref 1B) Unicracking-JHC of Los Angeles Basin Heavy Virgin Gas Oil for Gasoline and Turbine Fuel. Feedstock properties Gravity, °API

Product yields and properties 20.3

Distillation, 0-1160, of IBP

520

10

641

30

6B9

50

728

70

762

90

820

EP

890

Sulfur, wt %

1.33

Nitrogen, total, ppm

2270

Aromatics, wt %

35

Heterocyclics, wt %

21

Yields

One Stg Cat A

C 146 1-C 3, scf/b feed Liquid Products, vol % of feed Butanes 12.1 C5-C6 light gasoline 23.5 C7-plus gasoline 40.6 Turbine fuel 45.0 Total C 121.2 4-plus H Consumption, scf/b 2 feed 1750

One Stg Cat B

Two Stg Cat C

50

110

B.2 18.3 34.1 61 .1 121.7 1950

B.6 16.5 34.4 61.3 120.8 2110

Product properties C5- C6 Octane, F-1+3 ml TEL 99.0 C7-plus gasoline Octane, F-1+3 ml TEL 87.0 Naphtenes + Aromatics 70.1 Turbine Fuel Gravity, "API 37.5 Aromatics, vol % 34 Smoke point, mm 13.6 Luminometer number 25 Flash point, OF 117 - 51 Freeze point, OF 352-514 ASTM Di s t , , OF 10 %-95%

98.0

96.5

84.7 68.2

82.8 68.3

39.7 19 20.1 44 176 -58 358-510

41.4 2.0 29.7 67 120 -66 356-512

By contrast, gOOd quality jet fuels and gas oils must contain as few aromatics (smoke point sp8cific3tionsJ and linear paraffins (pour point or freeze points specifications) as possible; a~d

The two-stage process. which allows comolete hydrogenation of aromatics compounds

obtention of velY isopnraffinic products, is particUlarly well adapted for such

production (Table 3J. aromatics

saturat~on

When operating in one stage, an edditicnal

~

separate unit of

may be necessary (18J.

CONCLLBION Hyd~ocracking

on zeo· ita-based catalysts is a young process which has chiefley develop-

ped in the U.S.A. for gasoline and jet fuel production.

In Europe, where the need for such

products is less drastic, its development is much more limited and operating units are rather of a conventional type using amorphous catalysts and are oriented towards maximal middle distillates production.

Anywhere a great operating flexibility is not aimed at,

hydrocracking suffers severe competition from catalytic cracking.

The main advantage of

103 hydrocracking is indeed its versatility which allows, on the one hand. the treatment of heavy feedstocks rich in polynuclear aromatics. and particularly resistant to catalytic cracking. and.

on the other hand. very varied products of good quality. especially for

jet fuels and gas oils.

It has the drawback of requiring more expensive investment

(high pressure) and operating cost (hydrogen consumption ..• ). Future prospects seem to be more favorable for hydrocracking in Europe.

It is fairly

well established that on the horizon of 1985. the need for petroleum products will be characterised by a decrease in heavy fuel consumption and an increase of gasoline consumption.

So it will be necessary to transform the excess of heavy black cuts into

lighter white products.

So, hydrocracking will find its place in refining schemes for

producing middle-distillates and gasoline and will continue, in most cases, to compete with catalytic cracking.

It may be noticed that constraints coming from the struggle

against pollution may, to a certain extent, impede the development of catalytic cracking in Europe.

Besides. other way of progress are now open to hydrocracking.

Thus, production

of 80-160°C naphtas transformable in BTX aromatics for petrochemical use, seems to have a promising future.

The development of more active and more stable new catalysts would

allow the decrease of the temperature and hydrogen partial pressure, the increase of LHSV and, therefore, the decrease in the reactor-size, the cost of which is about 35 % of the total unit. REFERENCES 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32

A. Billon. J.P. Franck and J.P. Peries, Hydroa. Proc., sept. 1975, p. 139. P. Trambouze, A. Billon and J.P. Peries, Conference on hydrocracking, Institut BeIge du Petrole, Bruxelles, oct. 1968. N. Chaudhary and O.N. Saraf, Ind. Eng. Chem., P.R.D .• vol 14 (2), 1975, pp. 74-83. S.P. Ahuja, M.L. Derrien and J.F. Le Page, P.R.D., vol. 9 [31, 1970, p. 272. Petroleum Times, july 1970, p. 5. J.P. Franck, Rev. I.F.P •• vol XXXIII (4), 1978, 597-612. H.L. Coonradt and W.E. Garwood. Ind. Eng. Chem., P.D.D •• vol 3 (1), 1964, p. 38. J.W. Scott and A.G. Bridge, ACS/CIC Conference, Toronto, may 26, 1970. J.W. Ward, Hyd. Process, sept. 1975, 101-106. U.S. 3.132.089, 1964. U.S. 3.159.568, 1964. U.S. 3.835.027, 1974. U.S. 3.890.247, 1975. U.S. 4.120.825, 1978. M. El Malki, these Paris 1978. J.P. Franck [I.F.P.), unpublished results. U.S. 3.173.853. 1965. R.P. Vaell. J.L. Lafferty and J. Sosnowski, Proceed. Div. Refin. 1970, 1011-1021. D.E.W. Vaughan, Symp. on Propert. and Applicat. of zeolites, London. April 18-20, 1979. U.S. 3.293.192, 1966. C.V. Mc Daniel and P.K. Maher, 1 st Intern. Congress on Molecular Sieves. London, 1967, 186-195. G.T. Kerr, J. Catal., 1969. 15, 200-204. U.S. 3.354.077. 1967. U.S. 3.506.400, 1970. U.S. 3.591.488, 1971. Fr 2.303.764, 1976. U.S. 4.058.484, 1977. G.T. Kerr, J. Phys. Chem., 1968. 72, 2594. G.T. Kerr, J. Phys. Chem., 1969, 73, 2780. R. Beaumont and D. Barthomeuf, J. Catal., 1972, 26, 218. R. Beaumont and D. Barthomeuf, J. Catal., 1972, 27, 45. D.W. Breck. 'Zeolite Molecular Sieves', John Wiley and Sons, 1974, pp. 505-506.

104

33 34 35 36 37 38 39 40 41 42 43 44 45 46 -47 48 49 50 51 52 53 54 55 56

P. Gallezot. A. Alarcon-Diaz. J.A. Dalmon and B. Imelik. C.R. Acad. Sci .• Paris. 1974. t , 278. 1073-1075. P. Gallezot. P.A. Alarcon-Diaz. J.A. Dalmon. A.J. Renouprez and B. Imelik. J. Catal .• 1975. 39. 334. J.P. Franck and A. Billon (I.F.P.). unpublished results. Iranian Petroleum Institute. Bulletin 66. 1977. 17. J. Gonzalez Salas. these Paris. 1974. P. Gallezot. J. Datka. J. Massardier, M. Primet and B. Imelik. Vlth Congress on Catalysis. London. 1976. vol. 2. 696-704. G.D. Chukin. M.L. Landau. V. Ya. Kruglikov. D.A. Agievskii, B.V. Smirnov, A.L. Belozerov. V.S. Asrieva. N.V. Goncharova. E.D. Radchenko, D.O. Konolvachikov and A.V. Agafonov. Vlth Congress on Catalysis. London. 1976, vol. 2, 668-676. M.V. Landau. V. Ya. Kruglikov. N.V. Goncharova. 0.0. Konoval'Chikov. G.D. Chukin, B.V. Smirnov and V.I. Malevich. Kin. Cat .• 1976. 17. 1281-1287. A.D. Reichle, L.A. Pine. J.W. Ward and R.C. Hansford. Oil G.J .• july 29. 1977, 137-143. U.S. 4.139.433. 1979. U.S. 4.116.867. 1978. U.S. 3.197.399, 1965. U.S. 3.357.915, 1967. U.S. 3.287.257, 1966. U.S. 3.899.441, 1975. U.S. 3.835.028, 1974. U.S. 3.849.293. 1974. U.S. 3.692.692. 1972. Oil G.J •• july 29. 1974. U.S. 4.055.482, 1977. U.S. 4.148.750, 1979. J.W. Ward, R.C. Hansford, A.D. Reichle and J. Sosnowski, Oil G.J., may 28, 1973, 69-73. J.H. Duir. Oil G.J., august 21. 1967. 74-77. J.A. Rabo. R.D. Bezman and M.L. Poutsma. Proc. on the Symp. on Zeolites, Szeged (Hungary). sept. 11-14, 1978, 39-52.