FORM 2
THE PATENTS ACT, 1970
(39 of 1970)
&
THE PATENTS RULES, 2003
COMPLETE SPECIFICATIOIS
(See section 10, rule 13)

"PROCESS FOR CATALYTIC CONVERSION OF
HYDROCARBONS"

INDIAN OIL CORPORATION LIMITED
a Public Limited Company, having its head office at G-9, Ali Yavar
Jung Marg, Bandra (East), Mumbai-400051, India
The following specifïcation particularly describes the invention and the
manner in which it is to be performed.


FIELD OF THE INVENTION
The present invention relates to a process for conversion of hydrocarbon streams with
95% true boiling point less than 400oC, and preferably below 250°C to high yield of LPG
(olefins < 20 wt% of LPG) and high octane gasoiine with significantly lower olefins (< 2
wt%) and sulfur content using fluidizable solid acidic catalyst in a continuously
circulating fluidized dense bed reactor.
DESCRÏPTION OF THE PRIOR ART
tn recent years, significant attention is given on improvement of quality of fuels, both
gasoiine and diesel, to meet the stringent specifications. One of the requirements for
improving the fuel quality is to reduce the olefins content, which are photo-chemically
reactive and a major factor in the smog problem. Olefins are also undesirable due to
higher gum-forming tendency and also relatively lower motor octane number (MON).
Environmental regulations have also restricted the use of streams with higher sulfur and
aromatics (particularly benzene) as fuel. Due to the above factors along with the lower
vapor pressure requirements of gasoiine, some of the streams such as visbreaker / coker
naphtha, benzene rich light naphtha from reformer feed, high sulfur olefinic FCC
gasoiine, etc. no longer qualify for blending into gasoiine pool. At the same time, demand
of straight run naphtha is in declining trend due to its substitution by natural gas in
fertilizer and power sectors owing to obvious reasons. Hence, utilization of the above
streams is a problem to the refmers worldwide, which will further aggravate in days to
come.
On the other hand, the demand of LPG is in increasing trend in countries like India and
other countries of Asia e.g., China, Philippines, etc. The prime application is as domestic
cooking gas. LPG (rich in C3 and C4 paraffins) is emerging as a popular automobile fuel
due to its several advantages. The total number of vehicles run on LPG the world over is
estimated at four million, which is likely to increase further in coming days. As a result,
the demand of saturated LPG for use as auto-grade fuel will also increase. Tn such
2

situation, a process for conversion of low value naphtha streams to products like LPG is
going to be highly attractive to the refmers.
Conventionally, naphtha streams are thermally cracked in steam crackers at high
temperature (above 800°C) to produce light olefïns for use as petrochemical feedstocks.
However, since the cracking process is thermal in nature, the yield of dry gas (H2, C1 &
C3) including ethylene (50 wt% of feed) is much higher as compared to LPG (25 wt% of
feed). Although steam cracking is widely used, the process is energy intensive and not
very selective towards the LPG, particularly saturated LPG for automotive application.

Catalytic cracking is an alternate route for selective conversion of naphtha to light
olefins. In this regard, conventional fluid catalytic cracking (FCC) units can be adapted to
convert naphtha to light olefins either through injection of naphtha in the same riser along
with main feed, normally vacuüm gas oil (U.S. Patent. NOS. 6,538,169, 6,238,548 and
5,389,232) or through incorporation of a second riser (U.S. Patent. No.5,372,704 and
4,918,256). Most of these references deal with olefin rich naphtha streams e.g.
visbreaker, coker and FCC naphtha and does nbt include the straight run naphtha.
Hsing et al discloses a process in U.S. Patent. No. 5,637,207 for converting light paraffin
naphtha (C2 - C10) to light olefins (C2 - C5) and naphtha of enhanced octane through its
use as a lift fluid along with an inert gas at bottom of a conventional FCC riser. The
conversion of naphtha (C5-) was reported to be 52.84 vol% of naphtha feed with 2 wt%
ZSM-5 additive in Y-zeoIite based catalyst inventory.
The gasoline produced in the above methods has high octane with high sulfur and high
olefin content. Under typical FCC conditions, the naphtha conversion is not very high.
Also, the amount of naphtha processed is only fraction of the total FCC feed (<5 wt%)
due to the limitations in hardware design and catalyst activity dilution.
There are some processes disclosed exclusively for selective conversion of naphtha to
lower olefins (C2-C4) in fluid bed riser, dense bed reactor or fixed bed reactor using
ZSM-5 based catalyst. U.S. Pat. NOS. 6,548,725, 5,171,921 and 6,222,087 propose
3

catalytic cracking processes for conversion of naphtha to light olefins on catalysts
comprising phosphate doped ZSM-5 zeolite with or without promoter metal in a short
residence time riser or dense bed reactor. Yield of C3 plus C4 using catalytically cracked
light naphtha as feed was reported to be about 36 wt% of feed in U.S. Pat. No. 6,222,087.
U.S. Pat. No. 5,167,795 describes a process for the conversion of hydrocarbon feedstocks
consisting of C4-C7 paraffins, naphtha and light gas oils by catalytic cracking in riser and
quenching with similar activity catalyst to produce light olefins and aromatics, especially
benzene. The yield of LPG was reported to be 46.7 wt% using FCC gasoline as feed.
However, the dry gas produced was also high (15.4 wt% of feed).

U.S. Pat. NOS. 6,455,750, 6,153,089 and 6,602,403 mention processes for the upgradation
of catalytically or thermally cracked naphtha to light olefins (C3 - C4) and aromatics rich
and / or high octane gasoline using ZSM-5 based and large pore zeolite catalyst. U.S.
Pat. No. 6,153,089 by Das et al reports a process for producing 30-60 wt% of C3 and C4
hydrocarbons comprising olefins more than 50% at very high reactor temperature (above
570°C) and pressure similar to that of conventional FCC (0.5-2.5 atm(g)). One of the
products of these naphtha conversion processes is gasoline with higher aromatics and
hence higher octane no. Under the present and the emerging scenario, such gasoline
needs further pretreatment for reduction of sulfur and olefins before blending into
gasoline pool.
Some conventional processes attempt to reduce sulfur and olefins concentration in
naphtha by employing hydroprocessing stage subsequent to catalytic cracking. Such
hydroprocessing results in reduction of octane number. U.S. Pat. No.6,315;890 discloses
a two-step process for converting high octane naphtha having higher olefins and sulfur
like FCC naphtha to a gasoline having a reduced concentration of sulfur without
substantial reduction of octane number wherein the first step comprises cracking an
olefinic naphtha and the second step comprises a mild hydroprocessing. However, such
processes cannot upgrade the low octane coker and visbreaker naphtha.
4

In similar way, U.S. Pat. No.3,758,628 by Strckland et al discloses a two step process for
converting low octane parafinic naphtha to high octane gasoline. First step comprises
hydrocracking of paraffinic naphtha and the second step comprises a catalytic cracking.
The UOP hydrocracking process converts naphtha to high yield of saturated LPG with
production of low sulfur and zero olefin content gasoline. Since the octane number of
gasoline is substantially lower, it cannot be blended directly into a gasoline pool. Also,
such processes require higher capital investment and higher operating cost due to
requirement of external hydrogen and overall it becomes costlier to handle coker and
visbreaker naphtha.
To summarize, in the above processes via catalytic routes, maximum LPG yield is
reported to be 60 wt% using highly olefmic naphtha feedstock. These processes also
produce yery high dry gas yield. We could not find any process for catalytic conversion
of naphtha towards maximum production of highly saturated LPG along with a gasoline
of high octane. Therefore, there remains a need for a new process for production of
saturated LPG together with a high octane gasoline with substantially lower olefins and
sulfur which can be directly incorporated in refmery gasoline pool without additional
treatment using low value naphtha streams as feedstocks irrespective of their sources.
OBJECTS OF THE INVENTION
In light of the above background, it is the main object of the invention to derive a process
wherein naphtha, light gas oil irrespective of their source, in particular, straight run
naphtha as well as olefmic naphtha e.g., visbreaker naphtha, coker naphtha, FCC gasoline
in an operating refmery can be converted to value added products such as LPG and
gasoline.
It is another objective of the process to have required reactions of substantial cracking
along with reforming, alkylation, hydrogen transfer and Isomerization to produce high
yield LPG comprising predominantly C3 and C4 alkanes for its use as automobile grade
fuel and or other application such as cooking gas without using external hydrogen supply.
5

It is yet another objective of the present invention to produce a gasoline product with
substantially higher octane but lower quantity of olefms and sulfur without desulfurizing
the feed before handling.
Tt is still another objective of the invention to provide a single process wherein saturated
LPG and high octane gasoline can be produced in a single step catalytic process with
adequate flexibility to change the ratio of LPG to gasoline make, substantially at ease.
SUMMARY OF THE INVENTION
In distinction to the prior art processes, the present invention provides a process for
conversion of hydrocarbon streams with 95% true boiling point less than 400°C, and
preferably below 250°C using a solid fluidizable catalyst comprising a medium pore
crystalline alumino-silicates with or without Y-zeolite, non crystalline acidic materials or
combinations thereof in a continuously circulating dense fluidized bed reactor to produce
high yield of LPG (45 - 65 wt% of feed) and high-octane gasoline (RON > 92). The LPG
produced in the process of the invention is highly saturated with olefms content less than
20 wt%. The product gasoline is rich in aromatics having RON more than 92 with
substantially lower olefm content, less than 2 wt%. The catalyst system and the process
conditions of the present invention also favors very high degree of desulfurization
without use of external hydrogen resulting less than 5 wt% of feed sulfur as sulfur in
gasoline.
In accordance with the invention, the hydrocarbon feed is contacted with a hot
regenerated catalyst in a high velocity riser, which is connected to a dense fluidized bed
reactor for simultaneous cracking along with reforming, alkylation, hydrogen transfer and
Isomerization of the feed hydrocarbons under the operating conditions of temperature
range of 400 to 550°C, pressure range of 2 to 20 kg/cm2 (g) and WHSV range of 0. l to 20
hour"1. Spent catalyst is transported into a catalyst stripper from the reactor where steam
stripping is performed to remove entrained hydrocarbons from the spent solid catalyst.
Regeneration of the spent catalyst is performed in a fluidized bed regenerator in the
presence of air and or oxygen containing gases at a temperature ranging from 600 C to
6

700 C to burn off the coke and provide a regenerated catalyst with coke content of less
than 0.05 wt% at the bottom of the riser.
DETAILED DESCRIPTION OF THE INVENTION
Accordingly, the present invention provides a process for catalytic conversion of
hydrocarbon feed streams having 95% true boiling point less than about 400°C to LPG
comprising C3 and C4 hydrocarbons in the range of 30 to 65 wt% of the fresh
hydrocarbon in high yield and gasoline having octane number greater than about 90, said
process comprising:
(a) contacting the hydrocarbon feed stream with hot activated micro-spherical solid
fulidizable catalyst composition comprising medium pore crystalline alumino-
silicates and optionally "Y" type zeolite and non crystalline acidic materials in a
riser;
(b) transporting the mixture of hydrocarbon feed stream and the catalyst into a dense
bed reactor operating with weight hourly space velocity (WHSV) in the range of
0.1 to 20 hr '!, hydrocarbon feed stream residence time being greater than 5
seconds and catalyst residence time being greater than or equal to 60 seconds for
cracking the hydrocarbon feed stream at a temperature in the range of 400 to
550°C and pressure in the range of 2 to 20 kg/cm2 (g) thereby obtaining LPG
comprising C3 and C4 hydrocarbons with olefm content less than 20% (wt/wt)
and propane to butane ratio of more than 2 (wt/wt) and having propane in the
range of 50 to 70% by wt.
In an embodiment of the present application, the micro-spherical solid fulidizable catalyst
comprises 5 to 40 wt% of medium pore crystalline aïumino-silicates, O to 10 wt% of "Y"
type zeolites, O to 5 wt% of non crystalline acidic materials and remaining being non-
acidic components and binder.
In another embodiment of the present application, the medium pore crystalline aïumino-
silicates used comprises shape selective pentasil zeolite such as ZSM-5, ZSM-11 with
pore diameter in the range of 0.5 to 0.6 nanometers.
7

In yet another embodiment of the present application, the micro-spherical solid
fuüdizable catalyst comprises 10 to 30 wt% of the shape selective pentasil zeolite.
In still another embodiment of the present application, the "Y" type zeolite used is
selected from ReY or QSY zeolite.
In one more embodiment of the present application, the micro-spherical solid fuüdizable
catalyst composition comprises O to 5 wt% of ReY or USY zeolite.
In one another embodiment of the present application, the non-crystalline acidic material
is selected from the group consisting of alumina, silica-alumina, siiica-magnesia, silica
zirconia, silica thoria, silica-beryllia and silica titania.
In a further embodiment of the present application, the micro-spherical solid fulidizable
catalyst composition comprises O to 2 wt% of the non-crystalline acidic material.
In further more embodiment of the present application, the micro-spherical solid
fulidizable catalyst composition comprises 70 to 80 wt% of the binder.
In another embodiment of the present application, prior to contacting the micro-spherical
solid fulidizable catalyst composition with the hydrocarbon feed stream, the micro-
spherical solid fulidizable catalyst composition is activated by treating the same with
saturated steam under a temperature of about 550°C for a time period of about 3 hour.
In yet another embodiment of the present application, wherein in step (a), the ratio of the
activated micro-spherical solid fulidizable catalyst composition to hydrocarbon feed
stream is in the range of 2 to 10 wt/wt.
In still another embodiment of the present application, the hydrocarbon feed stream has
95% true boiling point less than about 250°C.
In one more embodiment of the present application, the hydrocarbon feed stream
comprises straight run or cracked components produced by catalytic processes such as
8

hydropossessing, FCC or thermal cracking processes like coking and visbreaking, and or
mixture thereof
In one another embodiment of the present application, wherein subsequent to step (b), the
process further comprises:
(c) separating the spent catalyst from the hydrocarbon product vapors thus formed at
a top portion of the dense bed reactor;
(d) passing the spent catalyst from the reactor into a catalyst stripper where the
catalyst is stripped to remove entrained hydrocarbons using steam, and
(e) burning the stripped catalyst of step (d) in a turbulent or fast fluidized bed
regenerator in presence of air and/or oxygen containing gases at a temperature in
the range of 600 to 700°C to burn off coke and provide a regenerated catalyst with
coke content less than 0.05 wt% at the bottom of the riser, which is re-circulated
to the riser.
In .one further embodiment of the present application, the catalyst is continuously
circulated between the fluidized bed regenerator, riser, dense bed reactor and stripper via
standpipe and slide valves.
In another embodiment of the present application, the yield of gasoline is in the range of
30 to 50% (wt/wt).
In yet another embodiment of the present application, the sulfur content in the product
gasoline is also reduced by about 90 to 95 % wt/wt to that of the hydrocarbon feed.
In still another embodiment of the present application, the olefm content of the gasoline
is less than or equal to 2 wt% irrespective of the feed olefm content and type of olefins in
the feed.
In a further embodiment of the present application, the ratio of ethane to ethane plus
ethylenè expressed in wt/wt in product is in the range of 0.65 to 0.80.
In conformity of the present inventiori, hydrocarbon streams with 95% true boiling point
less than 400°C, and preferably below 250°C is converted to very high yield of LPG
9

containing more than 80% saturates and high octane gasoline with substantially lower
olefin and sulfur content. The gasoline produced in this process is mostly olefin free (<2
wt%) irrespective of the olefin content of the feedstock.
In accordance with the present invention, the hot regenerated catalyst is injected at the
bottom of a high velocity up-flow riser wherein the hydrocarbon feed is injected through
a nozzle along with dispersion and atomization gas. The velocity in the riser is
maintained at a sufïlciently high value so that there is little or no slippage between the
hydrocarbon and catalyst flowing through the riser. The primary purpose of providing the
riser is to achieve proper mixing of the feed hydrocarbons and the regenerated catalyst.
The said riser is terminated into a large inventory of catalyst operating in bubbling bed or
preferably dense bed with WHSV in the range of 0.1-20 hr"1 at temperature in the range
of 400-550°C. The overhead pressure on the dense bed catalyst inventory is maintained
in the range of 2 - 20 kg/cm2(g). The hydrocarbon product effluent passes through a
conventional cyclone system to separate the cataiyst fïnes contained therein and is
discharged to a fractionator. The hydrocarbons separated from the catalyst are primarily
lighter gaseous components (C3 to C4hydrocarbons) and gasoline.
The carbonized spent catalyst is transported to a separate vessel acting as stripper by
maintaining a particular level in the dense bed reactor. Steam is introduced into the
catalyst stripper to remove any entrained hydrocarbons in the cataiyst. Stripped
hydrocarbons along with associated steam enter into the reactor top for recovery of
hydrocarbons.
The stripped catalyst is passed through a lift line to a dense or turbulent fluidized bed
regenerator where the coke on catalyst is burnt in presence of commercial Carbon
monoxide (CO) combustion promoter by air and or oxygen containing gases to achieve
coke on regenerated catalyst (CRC) lower than 0.05 wt%. Air and or oxygen containing
gases is also used as media to lift the catalyst into the regenerator for achieving partial
burning of coke in the lift line itself Regenerated catalyst is circulated back to the bottom
of the riser.
10

In the present invention, the delta coke (defined as the difference in CSC- wt% of coke
on spent catalyst and CRC) is low due to lower coke make in the cracking reactions,
which is expected to keep the regenerator temperature at relatively lower level as
compared to the conventional FCC operation. However, lower catalyst to oil ratio is
likely to compensate this effect and thereby maintain the regenerator temperature at least
to the same level as required for burning of coke on catalyst in presence of CO
combustion promoter. Flue gas leaving the regenerator catalyst bed is passed through
cyclones system for the separation of catalyst fmes and then discharged for pressure
reduction and energy recovery before venting through stack.
Besides the heat provided by the hot regenerated catalyst, external heat is supplied into
the riser through higher feed preheat temperature to achieve the desired temperature in
the riser and the dense bed reactor, which is preferably above 400°C. With a given feed
preheat temperature; the temperature at the top of the riser is controlled by the catalyst
flux into the riser.
Further details of feedstock, catalyst and products of the process of the present invention
are described below;
Feedstock
Feedstock for the present invention includes hydrocarbon fractions having 95% true
boiling point less than 400°C. The fractions could be straight run or cracked components
produced by catalytic processes, as for example, hydropossessing, FCC or thermal
cracking processes like coking, visbreaking, etc. and or mixture thereof. The conditïons
in the process of the present invention are adjusted depending on the type of the
feedstock to maximize the yield of LPG. The LPG yield, gasoline RON, aromatics yield
and extent of desulfurization, etc. are maximized if 95% true boiling point is lower than
250°C, Details of the feedstock properties' are outlined in the examples given in the
subsequent section of the patent. The above feedstock types are for illustration only and
the invention is not limited in any manner to only these feedstocks.
11

The following nomenclatures are generally applicable in all the examples cited here.

SRN Straight run naphtha
LCO Light Cycle oil
CN Coker naphtha
FCCN FCC gaseoline
MSN Mixed naphtha predominantly containins, SRN
MCN Mixed naphtha containing 50 wt% CN
MCFN 90 wt% MCN + 10 wt% FCC gasolïne
Catalyst
Catalyst employed in the process of the present invention predominantly consists of
pentasïl shape selective zeolites. Other active ingredients, as for example, Y zeolite in
rare earth and ultra stable form, non-crystalline acidic rnaterials or combinations thereof
are also added to the catalyst formulation to a limited extent for prqducing synergistic
effect towards maximum LPG production. Tt may be noted that conventional FCC
catalyst mainly consists of Y zeolite in different forms as active ingrediënt to accomplish
catalytic cracking reactions. Ranges as well as typical catalyst composition for the
process of the present invention and FCC process are summarized in Table-1 on weight
percentage.
Table-1: Catalyst composition of the present invention and conventional FCC
Components Process of the present
invention Conventional
FCC
Range Preferred range Range Typical
Shape selective pentasil zeolite 10-40 15-30 0-3.0 1.0
ReY / USY-zeolite 0-10 0-5 8-25 15.0
Non-crystalline acidic material 0-5 0-2 - -
Non-acidic components & binder 60-85 70-80 70-91 80.2
From the Table-1, it is seen that the catalyst composition in the process of the present
invention is markedly different in terms of pentasil zeolite and Y-zelite content as
compared to FCC catalyst. Examples of non-crystalline acid rnaterials are, alumina,
silica-alumina, silica-magnesia, silica-zirconia, silica-thoria, silica-beryllia, silica-titania.
Non-crystalline rnaterials contain about 10 to 40 wt. % alumina and rest is silica with or
without other promoters. Examples of rare earth components are lanthanum and cerium
in oxide form.
12

The pore size range of the active components namely, pentasil and Re-USY zeolite are in
the range of 0.5-0.6 and 0.8-1.1 nanometers respectively. The active components ïn the
catalyst of the process of the present invention, as for example, pentasil zeolite, Y zeolite,
etc. are supported on inactive materials of silica/alumina/silica alumina compounds
including kaolinites. The active components could be all mixed together before spray
drying or separately bound, supported and spray dried using conventional state of the art
spray drying technique and conditions used to produce FCC catalyst micro-spheres.
These spray-dried micro-spheres are then washed, rare earth exchanged and flash dried
following conventional methods to produce the finished catalyst particles. The fmished
micro-spheres containing active materials in separate particles are physically blended in
desired proportion to obtain a particular catalyst composition.
The typical physico-chemical properties of the fmished micro-spheres containing active
materials such as pentasil zeolite and Y-zeolite are given in Table-2 & 3 respectively.
Table-2: Physico-chemical properties of the pentasil zeolite based catalyst

Surface area, mVgm Fresh
Steamed 65-80 '
75-90
Crystallinity, wt% Fresh
Steamed 14-20
12. -18
Pore volume, cc/gm 0.30-0.40
Chemical Analysis, wt%
Al2O3
Na2O
Fe ' 25-35
0.35-0.45
0.4-0.6
Table - 3: Physico-chemical properties of the Y-zeolite based catalyst

Surface area, m2/g Fresh
steamed 110-180
100-140
% Crystallnity Fresh
Steamed 10-15
8-12
Unit Cell Size, oA Fresh
Steamed 24.35-24.75
24.2-24.6
Micro-pore area, m2/g Fresh
Steamed 65 - 100
60-90
Meso-pore area, m2/g Fresh
Steamed 45-80
40-50
Pore volume, cc/gm 0.25-0.38
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The preferred range of major physical properties of the finished fresh catalyst which are
required for the process of the present invention are summarized below:
Partiële size range, micron ; 20 -130
Partiële below 40 microns, wt% : < 20
Average partiële size, micron : 60-100
Average bulk density, micron : 0,6 - 1.0
Typically, the above properties and other related physical properties, e.g., attrition
resistance, fludizability etc. are in the same range as used in the conventionai FCC
process. Although pentasil zeolite materials such as zeolite ZSM-5, ZSM-11 have been
published as hydrocarbon cracking catalyst, the present invention is directed to specific
use of pentasil zeolite catalyst system for selectively cracking naphtha to produce light
saturates.
Products
The main product in the process of the present invention is LPG comprising C3 and C4
hydrocarbons, which is obtained with yield in the range of 45 to 65 wt% of feed. The
other important product is aromatic rich high-octane gasoline, which is almost olefm free,
A very small part of the feed is converted to coke and deposited on the circulating
catalyst system. The coke on catalyst is burnt in the regenerator and the exothermic heat
thus produced is utilized in the reactor. The typical range of the products obtained trom
the process of the invention is given in Table-5.
Table-5: Typical product yields obtained by the process of the present invention

PRODUCT Yield, wt% of feed
Dry Gas (H2+C1 +C2) 1-10
LPG(C3 + C4) 45-65
Gasoline (C5 ~200°C) 20-40
Coke 0.5-2.2
By changing the process conditions and design of catalyst, it is quite possible to alter the
gas to liquid product ratio to a significant extent.
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The LPG produced from the process of the invention is highly saturated vvith olefins less
than 20 wt%. The propane and butane percentages in corresponding C3 and C4 fractions
are more than 80 and 75 (w/wt) respectively. Typical LPG composition in the process of
the present invention is given below.
Table-6: Typical composition of LPG obtained by the process of the present
invention
Components Composition (wt % of LPG)
Propane 56-69
Propylene 5-6
Total C3 61-74
Isobutane 11.5-14.5
n-Butane 11.5-14.5
Isobutylene 1 - 3
1-Butene 0.4-0.8
t-2-Butene 1.25-2
cis-2-Butene 0.75-1.5
Total C4 saturates 23-29
Total C4 26-39
This shows that the LPG from the process of the present invention is highly saturated and
therefore suitable for its use as automotive fuel.
The dry gas is also highly saturated with 70 wt% of ethane in C2 fraction. Typical dry gas
composition in the process of the present invention is given below.
Table-7: Typical composition of dry gas obtained by the process of the present
invention
Components Composition (wt % of dry gas)
Hydrogen 19.1-21,5
Ethane 58-62
Ethylene 22.9-16.5
Ethane / total C2 69-78
Ethane / dry gas 55-62

l One of the important aspects of our invention is that olefm can be directly converted to
the product in the reactor itself unlike conventional reforming process where olefms are
totally saturated in a separate reactor before entering into the reforming reactor. The
catalyst and the contacting system in the present invention are capable to handle as much
15

olefins in the feed, without adding any external hydrogen. The gasoline produced in this
process is mostly olefin free irrespective of the olefin content of the feedstock.
The other important benefit of the invention is its flexibility to produce gasoline with
high octane rating but with signifïcantly lower olefin content as compared to the
conventional FCC gasoline. Aromatics such as toluene, xylenes, etc. are maximized in
the process. Table-8 shown below compares the typicat distribution of saturates, olefms
and aromatics along with benzené, toluene, xylene and ethyl benzene of the liquid
products of the present invention with that of FCC gasoline and reformate.

nable-8: Comparison of properties of liquid products of different processes
Feed Liquid Product of
present invention FCC gasoline Reformate
Wt%
Saturates 24-47 35-20 30-25
Olefms 1.2-2.2 50-55 Nil
Aromatics 51-74 15-25 70-75
Benzene 4.5-7.0 0.5-0.6 0.2-0.5
Toluene 18.-24 3.0-5.0 25-30
Ethyl benzene 2.0-3.8 0.5-1.0 5.0
m-p Xylene 9.0-16 2.5-3.5 25.0
o-Xylene 2.5-5.0 0.5-1.0 3.0
RON 92-98 90-95 >98
The benzene in gasoiine produced from the process of the invention can be maintained
less than 0.5 wt% by splitting the benzene rich light cut. The light cut can be routed to
ethylene cracker after extracting benzene or to naphtha isomerization unit. The typical
properties of the benzene rich light cut and lean cut after splitting the liquid product of
the process of the present invention are given below in table 9.
Table-9: Properties of benzene rich light & benzene lean cuts

PROPERTY Benzene rich cut Benzene lean cut
Benzene, wt% 38. S 0.5
Aromatics, wt% 38.8 58.4
Olefms, wt% 10.4 <0.5
RON 85 93
16

Therefore, the benzene lean cut of the liquid product obtained from the process of the
present invention can be directly blended into refinery gasoline pool without requiring
any additional pre-treatment
Unlike conventional FCC process, the process of the present invention desulfurises the
liquid products more than 90 wt% without requiring any external hydrogen. The
distribution of the sulfur in liquid product of the process of the present invention is
compared with that of the feed in table 10 given below. The total sulfur content of the
liquid product is less than 5 wt% of sulfur in the feed.
Table-10: Distribution of sulfur in liquid product
Feed Liquid product
Mercaptans Thiophenic Mercaptans Thiophenic
Sulfur compounds, ppm
High sulfur olefïn rich
naphtha 183 734 14 41
Low sulfur paraffm rich
naphtha 48 192 12 38
Thus the sulfur content of the liquid product of the process of the present invention is also
substantially lower, thereby allowing the direct into gasoline pool directly after extracting
the benzene.
The following examples will demonstrate flexibility of the present invention towards
various feedstocks and the quantum of LPG yield that can be produced from this process
along with other associated advantages. These examples are to be considered ülustrative
only and are not to be considered as limiting the scope of the present invention.
EXAMPLE-1: High yield of LPG
This example illustrates the important features of the process of the present invention to
produce very high LPG yield from various naphtha range feedstocks. Catalyst used in-this
example is medium pores pentasil zeolite and Re-USY zeolite based having properties as
shown in the Table-2 & 3. Initially, experiments were conducted in a circulating fluidized
bed riser pilot plant of 1.5 kg/hr feed capacity under very high reaction severity. The
17

crackability of naphtha range feedstock as well as the LPG selectivity under conventional
circulating fluidized bed riser conditions was found to be not much attractive.
We have discovered that in distinction to prior art processes for production of light
olefins and / or high octane gasoline using naphtha range hydrocarbon feeds, completely
different reaction conditions are needed for maximized production of LPG comprising
predominantly saturated alkanes. We have found that higher residence time of
hydrocarbon vapors above 5 seconds is essentiai for converting the naphtha range
hydrocarbons to saturated light paraffins. The higher residence time of hydrocarbons is
achieved by providing a dense bed reactor with very low WHSV. Higher reactor pressure
than that of conventional circulating fluidized bed catalytic processes commonly under
practice is found to favor the higher conversion of naphtha towards LPG. In distinction to
prior art processes for conversion of naphtha range hydrocarbon feeds, a lower
temperature is desirable to attain the objectives of the invention as outlined above.
In accordance with the present invention, a dense bed reactor with WHSV in the range of
0.1-20 hr"1 with higher contact time between feed and catalyst under higher pressure in
the range of 2-20 kg/cm2 (g) and relatively lower temperature in the range of 400-5 50°C
is provided to obtain very high yield of LPG. The comparison of product yields and
operating conditions of conventional riser system with higher reaction severity and the
present process of invention using similar feedstock is presented in Table-11.
TabVe-11: Comparison of conventional viser vtactov with dense bed reattor of

Riser pilot plant Dense bed reactor
Feed SRN SRN
Temperature, °C 570 480
Pressure, kg/cm (g) 1 5
Catalyst /Oil ratio (wt/wt) 23 4.4
WHSV, hour-1 150 1.5
Hydrocarbon residence time, sec. <1 >5
Products (Wt% fresh feed)
Dry gas (C3-) 1.31 7.84
LPG(C3 + C4) 14.48 44.64
Coke 0.1 1.01
18

It is seen from the Table-11, although high temperature and high catalyst to oil ratio vvere
maintained in riser reactor, the SRN feed could not be cracked much. The severity of the
reaction in terms of temperature, catalyst to oil ratio, WHSV and pressure was entirely
different in case of dense bed reactor. The reaction severity in terms of WHSV and
pressure are more dominating and responsible for high yield of LPG in dense bed reactor.
Therefore, the process of the present invention is distinct in application of combination of
reaction severity parameters to obtain very high LPG yield in the range of 45-65 wt% of
feed comprising more than 80% of €3 and €4 alkanes.
EXAMPLE-2: Catalyst composition
This example illustrates the importance of cataiyst composition in obtaining maximized
yield of LPG. Numerous experiments were conducted with different catalyst
compositions having composition given in table 12 using MCN feed in a stationery dense
fluidized bed reactor unit of 200 gm catalyst inventory operated in batch mode for
reaction, stripping and regeneration. The reactor pressure could be maintained upto 50
bar using a pressure control valve. All catalyst systems mentioned below are steamed at
550°C for 3 hours in presence of 100% steam before using in experiments.
Table-12: Catalyst systems of different composition

Catalyst system Cat-1 Cat-2 Cat-3
Shape selective pentasil zeolite 22.5 5.4 30
Re-USY-zeoIite 2.5 19 5
Non-crystaliine acidic components 1 1 I
Non-acidic components & binder 74 74.6 64
Table-13: Operating conditions

Temperature °C 480 .
Pressure kg/cm2 (g) 6
WHSV Hour-1 1.5
Residence time of catalyst in the reactor Minutes 10
Residence time of Hydrocarbons in the reactor Seconds 15
Similar operating conditions as summarized in Table-13 were maintained for all catalyst
systems mentioned above in table 12. The yields of LPG, dry gas and coke obtained with
these catalyst systems is given here below in the form of table 14:
19

Table-14: Catalyst systems of different composition

Catalyst system Cat-1 Cat-2 Cat-3
Yields, wt%
Dry gas 7.84 4.8 12.0
LPG 46.5 25.0 41.3
Coke 0.5 2.1 1.4
It is seen in above Table-14, LPG yield is lowest for Cat-2, which is having minimum
concentration of shape selective pentasil component. In creasing the concentration of
shape selective pentasil component is promoting dry gas formation. However, excessive
presence of shape selective pentasil component decreases the dry gas formation. Also, it
can be seen that minimum or excessive presence of shape selective pentasil component
increases the amount of coke being formed. In view of the above, it is also very important
to control the amount of the shape selective pentasil component added to the catalyst
composition. This example demonstrated that there is an optimum catalyst composition,
which gives maximum LPG yield with moderate coke and dry gas yield.
EXAMPLE-3: Optimum operating parameters
This example demonstrates that selection of the operating conditions is very important for
producing maximum LPG and minimum dry gas and coke. The effects of operating
conditions, particularly, vapor residence time, temperature, pressure, WHSV on product
yield pattern were tested with a particular catalyst having similar composition to Cat-1
using SRN as feedstock. The results are summarized below in table 15.
Table-15

Vapor residence time, seconds 5 7 10 5
Temperature, °C 480 480 480 520
Yields, wt%
Dry gas 6.82 7.02 7.15 9.0
LPG 42.2 46.5 42.6 39.9
Coke 1,5 1.7 2.57 3.50
WHSV was kept constant in the above runs. Vapor residence time was varied by
changing the reactor pressure. It is seen that LPG yield increases from 42.2 to 46.5 wt%
with increase in residence time from 5 to 7 seconds. When residence time is increased
20

further to 10 seconds, LPG yield reduces whereas with significant increase in coke yield.
The yield of dry gas also increases marginally. Even at residence time of 5 seconds, with
increase in temperature to 520°C from 480°C, LPG yield decreases with simultaneous
increase in both dry gas and coke.
We have found that for all the process parameters, there exists an optimum, which vary
dependïng on the hydrocarbon composition in feed and the catalysts system applied. The
Applicants have surprisingly found that in direct contradiction to the prior art process for
production of light olefins and / or high-octane gasoline using naphtha range hydrocarbon
feeds, lower temperature and higher pressure are desirable in the present invention to
attain the objectives of higher LPG yield and higher octane of gasoline product.
EXAMPLE-4: Processing of different types of naphtha
This example illustrates the capability of the process of the present invention to process
various naphtha range feedstocks containing different quantity of olefins as well as
sulfur A series of experiments were conducted using different hydrocarbon streams
namely SKN, CN, FCCN and mixture of these streams. The physico-chemicat properties
of the feeds used are summarized in Table-16.
The yields of LPG and dry gas with different feed streams are shown in Table-]7. It is
seen from Table-16 that the LPG produced is in the range of 45-67 wt%. Process is able
to handle all types of naphtha avaüable in an operating refmery to convent it to very high
yield of LPG. Also, the LPG yield increases with increase in olefins content in feed.
Table-16: Properties of naphtha feed stocks

Average boiling point, °C - (10% + 2* 50% + 90%) / 4
Table-17: Typical product yields of different feedstocks

Feed SRN MSN MCN r MCFN CN
Products, wt% of feed
Dry gas (H2, C1 & C2) 7.84 8.43 8.39 10.29 9.91
LPG (C3 + C4) 45.64 49.11 50.15 56.71 67.12
Coke yield 0.50 0.75 O.&O 1.20 1.7
21

F eed SRN MSN MCN MCFN CN
Density, gm/cc@150C 0.74 0.7326 0.7295 0.73 0.72
Sulrur, ppm 18 240 917 - 1600
Saturates 85.0 86.0 64.4 58.4 41.3
Olefins Nil 1.3 23.6 29.9 49.4
Aromatics 15.0 12.7 ' 12 11.7 9.3
RON 85.4 66.1 69.5 88.5 74
Average boiling point, °C 101.6 116.5 109.3 111 78.5
EXAMPLE-5: Liquid product composition and quality
This example illustrates the composition and quality of the liquid product obtained in the
process of the present invention. The distribution of hydrocarbon types, i.e., olefins,
aromatics and saturates in the liquid product obtained from different type feeds are given
below in Table-18:
Table-18: Saturates / olefins / aromatics distribution in liquid products

Feed SRN MSN MCN MCFN CN
Wt% in liquid product
Saturates 47.0 44.1 36.2 28.7 24.2
Olefins 2.2 1.3 2.5 2.1 2.2
Aromatics 50.8 54.6 61.3 69.2 73.6
On comparison with the feed composition as shown in Table-16 in ExampIe-4, it is seen
that above 95 wt% of the olefm reduction based on total olefin content in feed is
achievable in the process. In context of requirement of gasoline specifications with
respect to olefins content, this specific attribute of olefin reduction in the process of
invention is a distinct advantage.
The aromatics content in the liquid product is more than 50 wt%. The distribution of
benzene, toluene, xylene and ethyl benzene in the liquid products produced from different
feedstocks are shown in Table-19.
Table-19: Liquid product properties

Feed SRN MSN MCN FCCN CN
Wt% in liquid product
Benzene 7.01 5.68 6.3! 7.46 4.51
Toluene 17.86 18.8! 20.58 21.85 24.58
22

Ethyl benzene 2.06 2.91 2.87 3.02 3.69
m-p Xylene 9.16 12.10 h 12.46 15.28 15.91
O-Xylene 2.52. 3.41 3.45 4.34 4.97
The toluene and xylene contents in the liquid products of the propess of the invention are
quite high, which can be recovered as aromatics for use as petrochemical teedstocks. The
RON of the liquid products obtained from different feedstocks is compared with the RON
of feed in Table-20. The minimum RON of the liquid product is obtained trom SRN feed,
which does not contain any olefins. As the feed olefin content increases, the RON
increases. It is also seen that the RON of the liquid product was more than 92 irrespective
of the nature of the feedstocks.
Table-20: Research Octane Number of liquid product

Feed SRN MSN MCN MCFN CN
RON 85.4 66.1 69.5 88.5 74 -
RON 92.5 92.9 94.9 | 96.8 97.4
EXAMPLE-6: Olefin and sulfur reduction in liquid product
This example illustrates the capability of the process of the invention to convert the sulfur
in feed to hydrogen sulfide and thereby reduce the concentration of sulfur in the liquid
product.
The sulfur distribution of feed and liquid product are obtained by GC-PFPD / sulfur
analyzer. The sulfur distribution in products obtained from MSN and MCN feeds under
the process conditions similar to that given in Table-13 is shown in Table-21.
Table-21: Sulfur distribution in products

Feed MSN MCN
Total sulfur in feed, ppm 240 916
Weight percent of feed sulfur
Dry gas 70.36 72.76
LPG 17,62 18.48
Liquid 9.15 4.20
Coke 2.87 4.57

23
The total sulfur content of the feed in low sulfur naphtha (MSN) and high sulfur naphtha
(MCN) were 240 ppm and 917 ppm respectively. About 80% of the feed sulfur content
was in the form of thiophene and thiophene derivatives. The product of low sulfur
naphtha (MSN) had a sulfur content of 55 ppm by weight only. In case of high sulfur
naphtha feed (MCN), sulfur content in the product was 117 ppm. Total sulfur reduction in
the liquid product in all the experiments was in the range of 90 to 95 wt%. Reported
sulfur compounds in the liquid product contain about 25 wt% mercaptan compounds.
Significant part of the feed sulfur is being converted to hydrogen sulfide, which can
easily be removed from dry gas.
The Applicants respectfully submit that the process of the present invention should not be
understood as mere optimization of the operating parameter of known processes. The
Applicants would like to emphasize here that in addition to optimizing the operating
parameters, the applicants have also found the ideal catalyst composition which would
provide the necessary results. The Applicants have for the first time been able to arrive at
a method which is applicable to all types of naphtha / light gas oils. Further, for the first
time the Applicants have been able to arrive at a process that simultaneously converts all
types of hydrocarbon feed streams having 95% true boiling point less than about 400°C to
LPG comprising C3 and C4 hydrocarbons in the range of 30 to 65 wt% of the fresh
hydrocarbon in high yield and gasoline having octane number greater than about 90. Here
the applicants would like to highlight that till date no body has been provide a process
which can simultaneously produce LPG and gasoline in such high yield from even a
single feed, leave alone from a variety of feed streams.
The Applicants would also like to emphasize here that the process of the present
invention should be considered in its entirety. The various stages / steps of the process
(along with their respective operating parameters) should not be split and compared on an
individual basis with existing prior art documents. The Applicants have been able to
arrive at the unexpected and improved results after much trial and error and it is not
possible to theoretically predict that varying a particular parameter in the entire process
will result in improved result. As can be seen from our earlier experiments, varying any
24

individual parameter beyond a certain extent will only adversely affect the results and
will not give any improved results.
ADVANTAGES OF THE PRESENT INVENTION:
The important advantages of the process of the present invention are summarized below:
(i) Possessing of all types of naphtha/ light gas oils is possible.
(ii) Process uses circulating fluidized dense bed reactor-regenerator with adequate
flexibility of changing the LPG to gasoline ratio in the products.
(iii) Reactor pressure is higher than the conventional FCC.
(i v) Temperature of the reactor is quite low.

(v) High yield of saturated LPG is produced.
(vi) Highly saturated dry gas is produced.
(vu) High-octane gasoline product with substantïally lower olefin content.
(viii) In-situ desulfurisation resulting less than 5% of föed sulfur in gasoline product.
(ix) Low yield of coke and lower regenerator temperature.
25

WE CLAIM:
l. A process for catalytic conversion of hydrocarbon feed stream having 95% true
boiling point less than about 400°C to LPG in the range of 30 to 65 wt% of the
hydrocarbon feed stream, the said LPG comprising C3 and C4 hydrocarbons, and
gasoline having octane number greater than about 92, said process comprising:
(a) contacting in a riser the hydrocarbon feed stream with hot activated micro-
spherical solid fluidizable catalyst composition consisting of 5 to 40 wt% of medium
pore crystalline alumino-silicate, O to 10 wt% of "Y" type zeolite, O to 5 wt% of non
crystalline acidic material and the remaining being non-acidic component and binder; and
(b) transporting the mixture of hydrocarbon feed stream and the catalyst into a
dense bed reactor operating with weight hourly space velocity (WHSV) in the range of
0.1 to 20 hr" , hydrocarbon feed stream residence time being greater than 5 seconds and
catalyst residence time being greater than or equal to 60 seconds for cracking the
hydrocarbon feed stream at a temperature in the range of 400 to 500°C and pressure in
the range of 2 to 20 kg/cm (g) thereby obtaining LPG comprising C3 and C4
hydrocarbons with olefin content less than 20% (wt/wt) and propane to butane ratio of
more than 2 (wt/wt) and having propane in the range of 50 to 70%;
wherein the hydrocarbon feed stream is substantially devoid of hydrocarbons
having less than or equal to 4 carbon atoms and comprises straight run naphtha (SRN),
light cycle oil (LCO), coker naphtha (CN), FCC gasoline (FCCN): mixed naphtha
predominantly containing straight run naphtha (MSN), mixed naphtha containing 50 wt%
coker naphtha [[(MCN)]] CN, a feed stream comprising 90 wt% of MCN and 10 wt%
FCCN [[and]] or mixtures of two or more thereof.
2. A process as claimed in claim l, wherein the medium pore crystalline alumino silicates
used comprises shape selective pentasil zeolite such as ZSM-5. ZSM-1 l with pore
diameter in the range of 0.5 to 0.6 nanometers.
3. A process as claimed in claim 2, wherein the micro-spherical solid fluidizable catalyst
comprises 10 to 30 wt% of the shape selective pentasil zeolite.
26
4. A process as claimed in claim l, wherein the "Y" type zeolite used is selected from
ReY Or USYzeolite.
5. A process as claimed in claim 4, wherein the micro-spherical solid fluidizable catalyst
composition comprises O to 5 wt% of ReY of USY zeolite.
6. A process as claimed in claim l, wherein the non-crystalline acidic material is selected
from the group consisting of alumina, süica-alumina, silica- magnesia, silica zirconia,
silica thoria, silica-beryllia and silica titania.
7. A process as claimed in claim l, wherein the micro-spherical solid fluidizable catalyst
composition comprises O to 2 wt% of the non-crystalline acidic material.
8. A process as claimed in claim l, wherein the micro-spherical solid fluidizable catalyst
composition comprises 70 to 80 wt% of the binder.
9. A process as claimed in claim l, wherein prior to contacting the micro-spherical solid
fluidizable catalyst composition with the hydrocarbon feed stream, the micro-spherical
solid fluidizable catalyst composition is.activated by treating the same with saturated
steam under a temperature of about 550°C for a time period of about 3 hour.

10. A process as claimed in claim l wherein in step (a), the ratio of the activated micro-
spherical solid fluidizable catalyst composition to hydrocarbon feed stream is in the range
oflto 10 vvt/wt.
11. A process as claimed in claim l, wherein the hydrocarbon feed stream has 95% true
boiling point less than about 250°C.
12. A process as claimed in claim l, wherein the hydrocarbon feed stream comprises
straight run or cracked components produced by catalytic processes such as
27

hydropossessing, FCC or thermal cracking processes like coking and visbreaking, and or
mixture thereof,
13. A process as claimed in claim I wherein subsequent to step (b), the process further
comprises:
(c) separating the spent catalyst from the hydrocarbon product vapors thus formed at a
top portion of the dense bed reactor;
(d) passing the spent catalyst from the reactor into a catalyst stripper where the catalyst is
stripped to remove entrained hydrocarbons using steam. and
(e) burning the stripped catalyst of step (d) in a turbulent or fast fluidized bed regenerator
in presence of air and/or oxygen containing gases at a temperature in the range of 600 to
700°C to burn off coke and provide a regenerated catalyst with coke content less than
0.05 wt% at the bottom of the riser.
14. A process as claimed in claim l, wherein the catalyst is continuously circulated
between the fluidized bed regenerator, riser. dense bed reactor and stripper via standpipe
and slide valves.
15. A process as claimed in claim l, wherein the yield of gasoline is in the range of 30 to
50% (wt/wt).

16. A process as claimed in claim l, wherein sulfur content in the product gasoline is also
reduced by about 90 to 95 % wt/wt to that of the hydrocarbon feed.
17. A process as claimed in claim l, wherein the olefin content of the gasoline is less than
or equal to 2 wt% irrespective of the feed olefin content and type of olefms in the feed.
18. A process as claimed in claim l, wherein the ratio of ethane to ethane plus ethylene
expressed in wtlwt in product is in the range of 0.65 to 0.80.
Dated this 25th day of June, 2004