U.S. patent application number 16/971743 was filed with the patent office on 2020-12-17 for selective conversion of paraffinic naphtha to propylene in the presence of hydrogen.
The applicant listed for this patent is Total Research & Technology Feluy. Invention is credited to Raoul Dethier, Wolfgang Garcia, Nikolai Nesterenko, Valerie Vanrysselberghe.
Application Number | 20200392418 16/971743 |
Document ID | / |
Family ID | 1000005100977 |
Filed Date | 2020-12-17 |
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United States Patent
Application |
20200392418 |
Kind Code |
A1 |
Nesterenko; Nikolai ; et
al. |
December 17, 2020 |
Selective Conversion of Paraffinic Naphtha to Propylene in the
Presence of Hydrogen
Abstract
The invention relates to a process of catalytic conversion by
hydrocracking of paraffinic and naphthenic hydrocarbons from a
naphtha feedstock (1) to propylene, the process comprising the
steps of providing a naphtha feedstock (1) containing one or more
paraffins comprising 4 to 10 carbon atoms; and contacting said
naphtha feedstock (1) with a catalyst composition in the presence
of hydrogen in a reaction zone under hydrocracking conditions;
wherein the catalyst composition consists of one or more zeolite
catalysts comprising acid 10-membered ring channels.
Inventors: |
Nesterenko; Nikolai;
(Nivelles (Thines), BE) ; Dethier; Raoul;
(Schaerbeek, BE) ; Vanrysselberghe; Valerie;
(Aalbeke, BE) ; Garcia; Wolfgang;
(Braine-l'Alleud, BE) |
|
Applicant: |
Name |
City |
State |
Country |
Type |
Total Research & Technology Feluy |
Seneffe |
|
BE |
|
|
Family ID: |
1000005100977 |
Appl. No.: |
16/971743 |
Filed: |
February 21, 2019 |
PCT Filed: |
February 21, 2019 |
PCT NO: |
PCT/EP2019/054341 |
371 Date: |
August 21, 2020 |
Current U.S.
Class: |
1/1 |
Current CPC
Class: |
C10G 2400/28 20130101;
C10G 2300/4018 20130101; C10G 65/043 20130101; C10G 2300/4081
20130101; C10G 2300/4012 20130101; C10G 47/16 20130101; C10G
2400/20 20130101; C10G 2300/4006 20130101 |
International
Class: |
C10G 65/04 20060101
C10G065/04; C10G 47/16 20060101 C10G047/16 |
Foreign Application Data
Date |
Code |
Application Number |
Feb 22, 2018 |
EP |
18158160.4 |
Feb 22, 2018 |
EP |
18158175.2 |
Claims
1.-15. (canceled)
16. A process of catalytic conversion by hydrocracking of
paraffinic and naphthenic hydrocarbons from a naphtha feedstock to
propylene, the process being characterized in that it comprises the
following steps: a) providing a naphtha feedstock containing one or
more paraffins comprising 4 to 10 carbon atoms; and b) submitting
said naphtha feedstock to a hydrocracking step by contacting said
naphtha feedstock with a catalyst composition in the presence of
hydrogen in a reaction zone under hydrocracking conditions to
produce an effluent; c) submitting the effluent to a separation
step to recover propane; and d) submitting said propane to a step
of dehydrogenation into propylene in a propane dehydrogenation
reactor; e) collecting hydrogen produced in the step of
dehydrogenation into propylene, and recycling said hydrogen back to
the hydrocracking reaction zone in order to perform the
hydrocracking step; and in that the catalyst composition of the
hydrocracking step comprises one or more zeolite catalysts
comprising one or more acid 10-membered ring channels.
17. The process according to claim 16, characterized in that the
naphtha feedstock comprises at least 10 wt % of naphthenes as based
on the total weight of the naphtha feedstock.
18. The process according to claim 16, characterized in that the
one or more zeolite catalysts have a Si/Al molar ratio ranging from
10 to 100.
19. The process according to claim 16, characterized in that the
hydrocracking conditions of the hydrocracking step comprise: a. the
naphtha feedstock being contacted with the catalyst composition at
a temperature ranging from 200 to 600.degree. C., and/or b. the
naphtha feedstock being contacted with the catalyst composition at
a pressure ranging from 1 to 10 MPa.
20. The process according to claim 16, characterized in that the
catalyst composition contains no added noble metal and no added
transition metals.
21. The process according to claim 16, characterized in that the
catalyst composition contains at most 50 ppm wt of noble metal and
less than 0.05 wt % of transition metals as based on the total
weight of the catalyst composition.
22. The process according to claim 16, characterized in that the
hydrocracking conditions of the hydrocracking step comprise the
naphtha feedstock being contacted with the catalyst composition at
a WHSV (feed) of at least 0.1 h.sup.-1
23. The process according to claim 16, characterized in that, in
the hydrocracking step, hydrogen is provided to the naphtha
feedstock at a molar ratio H.sub.2/Naphtha ranging from 1000:1 to
1:1.
24. The process according to claim 16, characterized in that the
catalyst composition comprises at least 60 wt % of one or more
zeolite catalysts comprising one or more acid 10-membered ring
channels.
25. The process according to claim 16, characterized in that the
catalyst composition comprises one or more zeolite catalysts
selected from the list comprising ZSM-5, silicalite-1, ZSM-11,
silicalite-2, SSZ-46, MCM-68, CIT-1, SSZ-33, ZSM-8, Ferrierite,
FU-9, ZSM-35, ZSM-23, ZSM-22, Theta-1, NU-10, ZSM-50, EU-1, ZSM-57,
SAPO-11 and ZSM-48.
26. The process according to claim 16, characterized in that the
one or more zeolite catalysts comprising one or more acid
10-membered ring channels comprise or are zeolites catalysts of the
MFI-type.
27. The process according to claim 16, characterized in that the
catalyst composition comprises a binder selected from silica,
alpha-alumina, gamma-alumina, clays, alumina phosphates, mullite,
zirconia, titania, yttria, silicon nitride, silicon carbide, iron,
bronze and stainless steel, glass, and carbon.
28. The process according to claim 16, characterized in that the
step of dehydrogenation into propylene is performed at a
temperature ranging from 500 to 800.degree. C. and a partial
pressure of propylene below one atmosphere.
29. The process according to claim 16, characterized in that the
separation step comprises recovering a C4 paraffins fraction from
the effluent of step b) and a step of recycling back the C4
paraffins fraction to the hydrocracking reaction zone, and/or the
separation step comprises recovering a C5+ hydrocarbon fraction
from the effluent of step b) and a step of valorization of said C5+
hydrocarbon fraction as gasoline.
30. The process according to claim 16, characterized in that it
comprises a step of recovering the unconverted propane after the
step of dehydrogenation into propylene and recycling it to the
propane dehydrogenation reactor to be submitted to a further step
of dehydrogenation into propylene.
31. The use of one or more zeolite catalysts comprising the one or
more acid 10-membered ring channels in a process according to claim
16 of catalytic conversion by hydrocracking of paraffinic and
naphthenic hydrocarbons from a naphtha feedstock to propane, and
further dehydrogenation of said propane to propylene, characterized
in that the zeolite catalysts contains at most 50 ppm wt of noble
metal and less than 0.05 wt % of transition metals as based on the
total weight of the catalyst composition.
Description
FIELD OF THE INVENTION
[0001] The invention relates to a process for producing propane
from a naphtha feedstock comprising paraffinic and naphthenic
hydrocarbons, and to further produce propylene from the propane
recovered. The invention also relates to the use of catalyst
compositions to improve propane selectivity in a process for
producing propane from a naphtha feedstock comprising paraffinic
and naphthenic hydrocarbons.
BACKGROUND OF THE INVENTION
[0002] World demand for propylene is expected to continuously grow
up at an average annual rate of 4-5%. Currently, propylene is
mainly produced as a by-product of liquefied petroleum gas
(LPG)/naphtha steam crackers and Fluid Catalytic Cracking (FCC)
units. With the startup of additional ethane-based ethylene
capacities (lighter fractions like ethane or LPG are considered
very advantageous naphtha feedstock), the production of propylene
and aromatics via the steam cracker declines. Furthermore, the
demand for propylene is actually growing faster than that of
ethylene. To some extent, the propylene production can be optimized
by transforming an FCC unit into a petrochemical FCC after a major
revamp. However, a significant amount of dry gases and low value of
by-products will be produced. In addition, an important revamp of
the existing FCC unit and significantly higher process severity
would be required to crack the naphtha to propylene and
ethylene.
[0003] All the aforementioned factors have created an imbalance of
supply and demand for propylene: a gap is being established between
the available propylene supplies to meet ongoing demand growth.
While the markets have evolved to the point where modes of
propylene by-product production can no longer satisfy the demand
for propylene, the traditional classification of propylene as a
"by-product" begun to evolve into more of classification as a
"coproduct" or even a "primary product".
[0004] Propane dehydrogenation (PDH) is the fastest on-purpose
growing technology today to bridge the gap. The technology offers a
mean of producing propylene as a single product near consumer
markets. PDH system is suitable as stand-alone facilities not
requiring integration with a nearby olefin or refinery unit but
requires a secured supply of a big amount of propane (>0.5
Mt/y).
[0005] The PDH units used to be only appropriate for a limited
amount of geographical regions, where propane was highly available.
Today, the majority of the PDH projects are based on the imported
propane from the United States of America (US). A direct propane
import from the producers in the US or the Middle East has many
limits on volume and requires sophisticated logistics. On the
contrary, the market expects an important surplus in the
availability of naphtha due to a massive shift to ethane as
feedstock to steam-crakers and low oil prices. Therefore, there is
a need for the development of a technology to transform naphtha to
propylene with significant yield advantages versus naphtha steam
cracker and petro FCC.
[0006] The majority of the catalytic processes to crack paraffinic
naphtha is dedicated to direct production of propylene in absence
of hydrogen at a temperature higher than 550.degree. C. and low
partial pressure of hydrocarbon below atmospheric. These conditions
reduce hydrogen transfer reactions and favor propylene and
aromatics vs propane (D5943, Applied Catalysis A: General 398
(2011) 1-17). Propane is formed but as an undesirable side product.
One should mention also that under high-temperature conditions, a
lot of ethylene is produced which, in the context of growing
ethylene supply from ethane crackers, reduces significantly
attractiveness of the process.
[0007] The isomerization and hydrotreating processes of naphtha are
well known to produce clean fuel from linear paraffinic species
existing in light naphtha. Both processes occur in the presence of
hydrogen at moderate pressure on the metal-containing catalysts
with a purpose to increase fuel qualities. Light formation is big
drawback for both reactions. A post-reforming process known as
selectoforming was commercialized in the 1960s for raising the
octane rating of reformates while producing propane as a by-product
(Chen et al., 1968; Burd and Maziuk, 1972).
[0008] Some articles report naphtha hydrocracking, for example:
"High Yields of LPG Via n-Hexane Hydrocracking Using Unloaded
Acidic Zeolite Catalysts" in Journal Petroleum Science and
Technology, Volume 33, 2015--Issue 12, includes investigating the
hydrocracking of n-hexane, as a low-octane naphtha component to
high-octane gaseous motor fuel (LPG) in a pulse flow atmospheric
microreactor using untreated and steam-treated H-MOR, H-BEA, or
H-ZSM-5 zeolite catalysts. All zeolites catalysts were metal-free
and their bi-functionality depended only on the Bronsted zeolitic
acid sites. Zeolites catalysts H-ZSM-5 and St-H-ZSM-5 acquire very
low catalytic activities mainly due to their narrow pore structure,
as well as due to their partial pore filling by Al debris in case
of St-H-ZSM-5.
[0009] In "Conversion of light naphthas over sulfided nickel
erionite" in Industrial and Engineering Chemistry Research Volume
32, Issue 6, 1993, Pages 1003-1006, a natural zeolite erionite has
been exchanged with ammonium and nickel salts to yield a Ni/H
erionite catalyst that is active and stable for selectively
hydrocracking only the n-paraffins from light straight-run
naphthas. The primary product is a C5+ liquid that has 15-20 octane
numbers higher than the feed and a propane- and butane-rich gas
by-product. Results from a 110-day pilot plant run demonstrated
that a catalyst life of more than 1 year should be possible.
Naphthenes, aromatics, and isoparaffins are neither produced nor
consumed in this process, resulting in a C5+ liquid product that is
lower in benzene and total aromatics than attainable by catalytic
reforming of these feeds. Although no further work is planned with
this catalyst, a naphtha-upgrading process based on shape-selective
zeolitic hydrocracking could provide an attractive alternative to
catalytic reforming or isomerization for these hard-to-upgrade
naphthas. It should be particularly attractive in areas where the
propane and butane by-products have good value.
[0010] WO2012/071137 describes the use of a catalyst comprising at
least one zeolite having 10-membered ring channels and at least one
group VIb, VIIb and/or VIII metal, in the presence of hydrogen at
elevated temperature and elevated pressure. The process allowed to
convert at least 40 wt % of the paraffins comprising from 4 to 12
carbon atoms based on the total weight of paraffins in the feed to
ethane and/or propane to obtain a hydrocracked gas cracker naphtha
feedstock comprising ethane and/or propane. The process was
focussed on maximizing the production of ethane. The drawback of
the invention is an important formation of methane due to the
presence of the noble metals in the catalyst at about 6-20 wt
%.
[0011] US 2017/058210 discloses a process for producing BTX
comprising pyrolysis, aromatic ring opening and BTX recovery.
[0012] US 2016/369191 discloses a process for cracking a
hydrocarbon feedstock in a steam cracker unit where a liquid
hydrocarbon feedstock is fed to a hydrocracking unit, the stream
hydrocracked are then separated into a high content aromatics
stream and a gaseous stream comprising C2-C4 paraffins, hydrogen
and methane, the C2-C4 paraffins are then separated form said
gaseous stream and fed to a steam cracker unit.
[0013] US 2017/369795 discloses a process for producing C2 and C3
hydrocarbons, comprising a) subjecting a mixed hydrocarbon stream
to first hydrocracking in the presence of a first hydrocracking
catalyst to produce a first hydrocracking product stream; and b)
subjecting the first hydrocarbon product stream to C4
hydrocracking.
[0014] WO 98/56740 discloses a process for improving the conversion
of a hydrocarbon feedstock to light olefins comprising the steps of
thermally converting the hydrocarbon feedstock to produce an
effluent; quenching the effluent to produce a quenched effluent;
and contacting the quenched effluent with a light olefin-producing
cracking catalyst.
[0015] It was discovered that the naphtha may selectively be
transformed into propane on the metal free catalyst (i.e. noble and
transition metal free catalyst) with very low selectivity to
methane (from 0.3 to 1.5 wt % CH4). This opens an opportunity for a
high efficient process of production of propylene from naphtha via
a combination of hydrocracking of naphtha to propane followed by
dehydrogenation of propane to propylene. The dehydrogenation
process produces hydrogen, which could be used at least partially
for the first step.
SUMMARY OF THE INVENTION
[0016] According to a first aspect, the invention provides a
process of catalytic conversion by hydrocracking of a paraffinic
and naphthenic hydrocarbons from a naphtha feedstock to propylene,
wherein the process comprises the following steps: [0017] a)
providing a naphtha feedstock containing one or more paraffins
comprising 4 to 10 carbon atoms with preferably no olefins i.e. an
olefin content of less than 1 wt % preferably less than 0.1 wt %
even more preferably less than 0.01 wt % relating to the total
weight of said naphtha feedstock; and [0018] b) submitting said
naphtha feedstock to a hydrocracking step by contacting said
naphtha feedstock with a catalyst composition in the presence of
hydrogen in a reaction zone under hydrocracking conditions to
produce an effluent; [0019] c) submitting the effluent to a
separation step to recover propane; and [0020] d) submitting said
propane to a step of dehydrogenation into propylene in a propane
dehydrogenation reactor; [0021] e) collecting hydrogen (21)
produced in the step of dehydrogenation into propylene (23), and
recycling said hydrogen (21) back to the hydrocracking reaction
zone in order to perform the hydrocracking step (3) and wherein the
catalyst composition of the hydrocracking step comprises one or
more zeolite catalysts comprising an acid 10-membered ring
channels.
[0022] Surprisingly, it was found by the inventors that the
catalyst composition comprising acid 10-membered ring channels was
particularly resistant to deactivation. Indeed the hydrogen
recycled from the dehydrogenation step to the hydrocracking
reaction zone is not necessarily of the highest purity. This
hydrogen can contain methane or other impurities like sulphur at
relatively high concentration that deactivate traditional
hydrocracking catalyst. The use of the catalyst composition
comprising acid 10-membered ring channels is resistant to
deactivation induced by the methane contained in hydrogen. Such
catalyst allows avoiding expensive purification of the hydrogen.
The hydrogen feed stream can for instance contain up to 5 wt % of
methane, preferably up to 10 wt %, even more preferably up to 20 wt
% of methane based on the total weight of the hydrogen feed stream.
Other impurities may also be present such as sulphur compounds,
such as H.sub.2S, at concentration for instance up to 0.1% wt,
preferably 1% wt, even more preferably up to 5 wt % based on the
total weight of the hydrogen feed stream. In a preferred
embodiment, the one or more acid zeolite catalysts or catalyst
composition are metal-free, containing less than 1000 ppm of noble
metal and less than 1% of transition metals. The content of the
noble metals is below 1000 ppm, preferably below 500 ppm, more
preferably at most 250 ppm even more preferably at most 50 ppm wt,
the most preferred being no noble metal at all, that is to say
below the detection limit. In exceptional cases, traces of noble
metals (Pt, Pd) may be introduced to increase the stability of the
zeolite and hydrogenate the coke precursors. However, higher
concentration of the metals may lead to changes of the reaction
mechanism. Presence of a significant amount of noble metals
(>0.2 wt %) on the catalyst leads to a different reaction
mechanism, which involves activation of the hydrogen on the metal
(hydrogen spillover) followed by subsequent cracking. This leads to
a different selectivity pattern and formation of higher amount of
ethane and methane. It is therefore preferred to have no noble
metals at all.
[0023] As used herein the terms "metal-free" indicate that the one
or more acid zeolite catalysts are also devoid of any transition
metal selected from the groups of Fe, Ni, Co, W, Mo, Ga, Zn. The
content of the transition metals is below 1.0 wt %, preferably
below 0.5 wt % more preferably below 0.05 wt % even more preferably
below 0.01 wt %. Traces of these metals may be present on the
catalyst as impurities from the binder, e.g. a component of the
clays. Surprisingly, it was found by the inventors that metal-free
zeolite catalysts comprising acid 10-membered ring channels showed
a high selectivity to propane and at the same time, stable
performance in the presence of the aromatics presents in the
naphtha feedstock. The use of said catalysts in catalyst
compositions allows a direct cracking in presence of hydrogen of a
naphtha feedstock comprising paraffinic and naphthenic hydrocarbons
without deactivation of the catalyst.
[0024] It is noted that U.S. Pat. No. 4,061,690 discloses a method
of catalytic conversion of a butane cut to propane by means of a
catalyst consisting of acid mordenite in the presence of hydrogen,
with a partial hydrogen pressure higher than 0.5 MPa. The results
show that the use of a large pore mordenite exchanged by protons
(Zeolons H) at 400.degree. C. and at a pressure of 3 MPa had a
propane yield of 75.0 wt % for a rate of conversion of 92%, which
corresponds to a propane selectivity of approximately 69%. The
weighted analysis of the catalyst was SI: 40.6%; Al 6.2%; Na: 0.2%.
The feedstock comprised a mixture of n-butane (60 wt %) and of
iso-butane (40 wt %). These results were promising, however, it was
found that MOR catalysts, having 12-membered ring channels,
deactivate in the presence heavier feedstock with the number of
atoms sufficient to form aromatics or even containing the aromatics
(the naphtha feedstock). Thus, it was not possible to use MOR
catalysts in direct hydrocracking of naphtha (DCN). By way of
examples, it was shown that MOR deactivates very fast naphtha
feedstock and is not suitable.
[0025] With preference, one or more of the following embodiments
can be used to better define the inventive process: [0026] The
naphtha feedstock comprises C4-C10 hydrocarbons with a potential
presence of the aromatics. The feedstock contains preferably no
olefins i.e. an olefin content of less than 0.1 wt % relating to
the total weight of said naphtha feedstock. The C5+ hydrocarbons
lead to a significant deactivation of the 12-members ring zeolites
catalysts. The small pore erionite could crack only a part of the
feedstock. The solutions allow treating the total amount of the
feedstock without significant deactivation of catalyst. [0027] The
naphtha feedstock comprises at least 10 wt % of naphthenes as based
on the total weight of the naphtha feedstock. [0028] The naphtha
feedstock comprises at least 2.0 wt % of aromatics of five or more
carbon atoms as based on the total weight of the naphtha feedstock,
preferably at least 2.2 wt %. [0029] The naphtha feedstock
comprises at most 10.0 wt % of aromatics of five or more carbon
atoms as based on the total weight of the naphtha feedstock,
preferably at most 9.0 wt %. [0030] The catalyst composition is
metal-free, containing less than 1000 wt ppm, preferably less than
50 ppm wt, more preferably at most 5 ppm wt of noble metal and less
than 1 wt %, preferably less than 0.05 wt % even more preferably
below 0.01 wt % of transition metals as based on the total weight
of the catalyst composition. The catalyst composition is not
sensitive to the presence of sulfur and provides low C1-02
selectivity. [0031] The hydrocracking conditions of the
hydrocracking step comprise the naphtha feedstock being contacted
with the catalyst composition at a temperature ranging from
200.degree. C. to 600.degree. C., preferably ranging from
250.degree. C. to 550.degree. C., more preferably ranging from
300.degree. C. to 450.degree. C. [0032] The hydrocracking
conditions of the hydrocracking step comprise the naphtha feedstock
being contacted with the catalyst at a pressure ranging from 1 to
10 MPa, preferably in the range of 2 to 6 MPa, more preferably from
3 to 5 MPa. [0033] The hydrocracking conditions of the
hydrocracking step comprise the naphtha feedstock being contacted
with the catalyst at a WHSV (feed) of at least 0.1 h.sup.-1,
preferably is ranging from 0.1 h.sup.-1 to 10.0 h.sup.-1, more
preferably from 0.5 h.sup.-1 to 8.0 h.sup.-1, even more preferably
from 1.0 h.sup.-1 to 6.0 h.sup.-1, and most preferably from 1.5
h.sup.-1 to 5.0 h.sup.-1. [0034] In the hydrocracking step,
hydrogen is provided to the naphtha feedstock at a molar ratio
H.sub.2/Naphtha ranging from 1000:1 to 1:1, preferably from 100:1
to 1:1, more preferably from 20:1 to 1:1. The hydrocracking step is
particularly advantaging in that hydrogen with a low purity can
preferably be used. That is to say that the hydrogen feed stream
can for instance contain up to 5 wt % of methane, preferably up to
10 wt %, even more preferably up to 20 wt % of methane based on the
total weight of the hydrogen feed stream. Other impurities may also
be present such as sulphur compounds, such as H.sub.2S, at
concentration up to 0.1% wt, preferably 1% wt, even more preferably
up to 5 wt % based on the total weight of the hydrogen feed stream.
[0035] The hydrocracking step is performed in a single reactor.
[0036] With preference, one or more of the following embodiments
can be used to better define the catalyst used in the inventive
process: [0037] The one or more zeolite catalysts have a Si/Al
molar ratio ranging from 10 to 100, preferably from 20 to 80, more
preferably from 30 to 60. [0038] At least 50 wt % of said one or
more zeolite catalysts comprising an acid 10-membered ring channels
are in their hydrogen form as based on the total weight of the
zeolites catalysts. [0039] The catalyst composition comprises one
or more zeolites catalysts comprising an acid 10-membered ring
channels selected from the list comprising ZSM-5, silicalite-1,
ZSM-11, silicalite-2, SSZ-46, MCM-68, CIT-1, SSZ-33, ZSM-8,
Ferrierite, FU-9, ZSM-35, ZSM-23, ZSM-22, Theta-1, NU-10, ZSM-50,
EU-1, ZSM-57, SAPO-11 and ZSM-48 being preferably under their
acidic form or H-form. [0040] Optionally, the one or more zeolites
catalysts comprising an acid 10-membered ring channels may be
modified by phosphorous, steaming, leaching, mesoporization,
dealumination or modification with alkali-earth or rare-earth
metals. [0041] The preferred zeolite catalyst structure is MFI
(H-ZSM-5). [0042] The catalyst composition comprises a binder
selected from silica, alpha-alumina, gamma-alumina, clays, alumina
phosphates, mullite, zirconia, titania, yttria, silicon nitride,
silicon carbide, iron, bronze and stainless steel, glass, and
carbon, preferably the binder is alumina.
[0043] In an embodiment, the step of dehydrogenation into propylene
is performed at a temperature ranging from 500 to 800.degree. C.
and a partial pressure of propylene below one atmosphere and/or the
process further comprises a step of collecting hydrogen produced in
the step of dehydrogenation into propylene, and a step of recycling
said hydrogen back to the hydrocracking reaction zone in order to
perform the hydrocracking step.
[0044] In an embodiment, the separation step comprises recovering a
C4 paraffins fraction from the effluent of step b) and a step of
recycling back the C4 paraffins fraction to the hydrocrakrting
reaction zone, and/or the separation step comprises recovering a
C5+ hydrocarbon fraction from the effluent of step b) and a step of
valorization of said C5+ hydrocarbon fraction as gasoline.
[0045] In an embodiment, the process comprises a step of recovering
the unconverted propane after the step of dehydrogenation into
propylene and recycling it to the propane dehydrogenation reactor
to be submitted to a further step of dehydrogenation into
propylene.
[0046] According to a second aspect, the invention provides the use
of one or more zeolite catalysts comprising acid 10-membered ring
channels in a process as defined according to the first aspect of
catalytic conversion by hydrocracking of paraffinic and naphthenic
hydrocarbons from a naphtha feedstock to propane, and further
dehydrogenation of said propane to propylene, wherein the one or
more zeolites catalysts comprising an acid 10-membered ring
channels are metal-free.
[0047] With preference, the one or more zeolites catalysts
comprising an acid 10-membered ring channels comprise or are
zeolites catalysts of the MFI-type, preferably the one or more
zeolites catalysts comprising an acid 10-membered ring channels are
or comprise H-ZSM-5.
DESCRIPTION OF THE FIGURES
[0048] FIG. 1 illustrates the process according to the
invention.
DETAILED DESCRIPTION OF THE INVENTION
[0049] For the purpose of the invention the following definitions
are given:
[0050] Naphtha is mainly a mixture of straight-chain, branched and
cyclic aliphatic hydrocarbons. Naphtha is generally divided into
light naphtha having from 4 to 10 carbon atoms per molecule and
heavy naphtha having from 7 to 12 carbons per molecule. Typically,
light naphtha contains naphthenes, such as cyclohexane and
methyl-cyclopentane, and linear and branched paraffins or alkanes,
such as hexane and pentane. Light naphtha typically contains 60% to
99% by weight of paraffins and cycloparaffins.
[0051] The term "metal free" as used herein means that in the
course of the preparation of the catalyst composition, no metal is
willingly added. It is always possible that some metals (noble or
transition) are present as pollution or traces, but it shall be
understood that they are present at a very low concentration if not
below the detection limit. As a preferred embodiment, a metal-free
catalyst composition contains at most 50 ppm wt, more preferably at
most 5 ppm wt even more preferably at most 1 ppm wt of noble metal
and less than 0.05 wt % even more preferably less than 0.01 wt % of
transition metals as based on the total weight of the catalyst
composition.
[0052] The term "alkane" or "alkanes" as used herein describes
acyclic branched or unbranched hydrocarbons having the general
formula C.sub.nH.sub.2n+2, and therefore consisting entirely of
hydrogen atoms and saturated carbon atoms; see e.g. IUPAC.
Compendium of Chemical Terminology, 2nd ed. (1997). The term
"alkanes" accordingly describes unbranched alkanes
("normal-paraffins"or "n-paraffins" or "n-alkanes") and branched
alkanes ("iso-paraffins" or "iso-alkanes") but excludes naphthenes
(cycloalkanes).
[0053] The term "aromatic hydrocarbons" or "aromatics" relates to
cyclically conjugated hydrocarbon with a stability (due to
derealization) that is significantly greater than that of a
hypothetical localized structure (e.g. Kekule structure). The most
common method for determining aromaticity of a given hydrocarbon is
the observation of diatropicity in the .sup.1H NMR spectrum.
[0054] The terms "naphthenic hydrocarbons" or "naphthenes" or
"cycloalkanes" as used herein describes saturated cyclic
hydrocarbons.
[0055] The term "olefin" as used herein relates to an unsaturated
hydrocarbon compound containing at least one carbon-carbon double
bond. Preferably, the term "olefins" relates to a mixture
comprising two or more selected from ethylene, propylene,
butadiene, butylene-1, isobutylene, isoprene, and
cyclopentadiene.
[0056] The term "LPG" as used herein refers to the well-established
acronym for the term "liquefied petroleum gas". LPG, as used
herein, generally consists of a blend of C2-C4 hydrocarbons i.e. a
mixture of C2, C3, and C4 hydrocarbons.
[0057] One of the petrochemical products which may be produced in
the process of the present invention is BTX. The term "BTX" as used
herein relates to a mixture of benzene, toluene, and xylenes.
[0058] As used herein, the term "C# hydrocarbons", wherein "#" is a
positive integer, is meant to describe all hydrocarbons having #
carbon atoms. C# hydrocarbons are sometimes indicated as just "C#".
Moreover, the term "C#+ hydrocarbons" is meant to describe all
hydrocarbon molecules having # or more carbon atoms. Accordingly,
the term "C5+ hydrocarbons" is meant to describe a mixture of
hydrocarbons having 5 or more carbon atoms. The term "C5+ alkanes"
accordingly relates to alkanes having 5 or more carbon atoms.
[0059] The term "zeolite catalyst" refers to a molecular sieve
aluminosilicate material. Reference herein to a zeolite catalyst
having acid 10-membered ring channels is to a zeolite or
aluminosilicate having 10-membered ring channels in one direction,
optionally intersected with 8, 9 or 10-membered ring channels in
another direction.
[0060] The terms "comprising", "comprises" and "comprised of" as
used herein are synonymous with "including", "includes" or
"containing", "contains", and are inclusive or open-ended and do
not exclude additional, non-recited members, elements or method
steps. The terms "comprising", "comprises" and "comprised of" also
include the term "consisting of".
[0061] The recitation of numerical ranges by endpoints includes all
integer numbers and, where appropriate, fractions subsumed within
that range (e.g. 1 to 5 can include 1, 2, 3, 4 when referring to,
for example, a number of elements, and can also include 1.5, 2,
2.75 and 3.80, when referring to, for example, measurements). The
recitation of endpoints also includes the recited endpoint values
themselves (e.g. from 1.0 to 5.0 includes both 1.0 and 5.0). Any
numerical range recited herein is intended to include all
sub-ranges subsumed therein.
[0062] The particular features, structures, characteristics or
embodiments may be combined in any suitable manner, as would be
apparent to a person skilled in the art from this disclosure, in
one or more embodiments.
[0063] The process of the invention can be operated at high naphtha
feedstock conversion, reducing the need for a further cracking
process.
[0064] Reference is made to FIG. 1. The invention the invention
provides a process of catalytic conversion by hydrocracking of a
paraffinic and naphthenic hydrocarbons from a naphtha feedstock 1
to propylene, wherein the process comprises the following steps:
[0065] a) providing a naphtha feedstock 1 containing one or more
paraffins comprising 4 to 10 carbon atoms with preferably no
olefins i.e. an olefin content of less than 1 wt % preferably less
than 0.1 wt % even more preferably less than 0.01 wt % relating to
the total weight of said naphtha feedstock; and [0066] b)
submitting said naphtha feedstock 1 to a hydrocracking step 3 by
contacting said naphtha feedstock 1 with a catalyst composition in
the presence of hydrogen in a reaction zone under hydrocracking
conditions to produce an effluent 5; [0067] c) submitting the
effluent 5 to a separation step 7 to recover propane 15; and d)
submitting said propane 15 to a step of dehydrogenation 19 into
propylene 23 in a propane dehydrogenation reactor; [0068] e)
collecting hydrogen 21 produced in the step of dehydrogenation into
propylene 23, and recycling said hydrogen 21 back to the
hydrocracking reaction zone in order to perform the hydrocracking
step 3 and wherein the catalyst composition of the hydrocracking
step 3 comprises one or more zeolite catalysts comprising an acid
10-membered ring channels.
[0069] The effluent 5 resulting from the hydrocracking step 3 can
be submitted to a separation step 7, wherein: [0070] The C4
paraffins fraction 11 can be recycled back to the hydrocracking
reactor to be submitted to a further hydrocracking step 3. [0071]
The C5+ hydrocarbons fraction 17 can be valorized as gasoline or
aromatics. Indeed, the C5+ fraction shows a significant improvement
in Research Octane Number (RON) as compared to the naphtha
feedstock 1. [0072] The hydrogen 9 can be recycled back to the
hydrocracking reactor to be used in the hydrocracking step 3.
[0073] The C1 and C2 fraction 13 (methane and ethane fraction) can
be use as off-gas.
[0074] The process further comprises a step of collecting hydrogen
21 produced in the step of dehydrogenation into propylene 23, and a
step of recycling said hydrogen 21 back to the hydrocracking
reaction zone in order to perform the hydrocracking step 3.
[0075] With preference, the process comprises a step of recovering
the unconverted propane 15 after the step of dehydrogenation 19
into propylene 23 and recycling it to the propane dehydrogenation
reactor to be submitted to a further step of dehydrogenation 19
into propylene.
[0076] The naphtha feedstock 1 used in the invention comprises
paraffinic and naphthenic hydrocarbons, preferably the naphtha
feedstock 1 comprises one or more paraffins comprising 4 to 10
carbon atoms.
[0077] The naphtha feedstock 1 may comprise compounds other than
paraffins. Preferably, the naphtha feedstock comprises at least 10
wt % of paraffins comprising 4 to 10 carbon atoms as based on the
total weight of the naphtha feedstock, more preferably at least 50
wt %, more preferably at least 60 wt % of paraffins comprising 4 to
10 carbon atoms.
[0078] Preferably, the naphtha feedstock 1 comprises in the range
of from 10 wt % to 100 wt % of paraffins comprising 4 to 10 carbon
atoms as based on the total weight of the naphtha feedstock, more
preferably of from 50 wt % to 99.5 wt %, more preferably of from 60
wt % to 95 wt % of paraffins comprising 4 to 10 carbon atoms.
[0079] The naphtha feedstock 1 may comprise straight run naphtha or
naphtha fractions derived from natural gas, natural gas liquids or
associated gas. The naphtha feedstock 1 may comprise naphtha
fractions derived from pyrolysis gas. The naphtha feedstock 1 may
also comprise naphthas or naphtha fractions obtained from a
Fischer-Tropsch process for synthesising hydrocarbons from hydrogen
and carbon monoxide. For example, the naphtha feedstock 1 is or
comprises desalted crude oil.
[0080] The naphtha feedstock 1 may also comprise higher paraffins,
i.e. paraffins comprising more than 10 carbon atoms. Cracking such
higher paraffins typically requires the use of temperatures and
pressures which are at the higher end of the preferred temperature
and pressure ranges.
[0081] Preferably, the naphtha feedstock 1 comprises at least 10%
of naphthenes. More preferably, the naphtha feedstock 1 comprises
in the range of from 10 to 40 wt %, more preferably of from 50 to
90 wt % of naphthenes and paraffins C6+, based on the total weight
of the naphtha feedstock.
[0082] The naphtha feedstock 1 may comprise olefins. However, as
the olefins are hydrogenated during the hydrocracking process, the
presence of olefins results in an undesired increased hydrogen
consumption. Preferably, the naphtha feedstock 1 comprises in the
range of from 0 to 20 wt % of olefins, based on the total weight of
the naphtha feedstock, more preferably of from 0 to 10 wt % of
olefins. Optionally, the naphtha feedstock 1 is subjected to a
hydrogenation treatment prior to being supplied to a process
according to the present invention.
[0083] In an embodiment, the naphtha feedstock 1 comprises at least
2.0 wt % of aromatics of five or more carbon atoms as based on the
total weight of the naphtha feedstock, preferably at least 2.2 wt
%, more preferably at least 2.5 wt %.
[0084] In an embodiment, the naphtha feedstock 1 comprises at most
10.0 wt % of aromatics of five or more carbon atoms as based on the
total weight of the naphtha feedstock, preferably at most 9.0 wt
%.
[0085] In a preferred embodiment, the hydrocracking step 3 is
performed in a single reactor. Indeed, the invention provides a
one-stage hydrocracking process.
[0086] The one or more zeolites catalysts having acid 10-membered
ring channels that can be used for the invention can be selected
from: [0087] one-dimensional zeolites catalysts having 10-membered
ring channels in one direction, which are not intersected by others
channels from another direction; [0088] three-dimensional zeolites
catalysts having intersecting channels in at least two directions,
whereby the channels in one direction are 10-membered ring
channels, intersected by 8, 9 or 10-membered ring channels in
another direction.
[0089] Examples of zeolites catalysts having acid 10-membered ring
channels suitable for the process of the invention can be of, but
not limited to, the MFI-type, the MEL-type, the MSE-type, the
CON-type, the ZSM-8-type, the FER-type, the MTT-type, the TON-type,
the EUO-type, the MFS-type, the AEL-type and the ZSM-48-type
zeolites catalysts. Preferably, the catalyst is or comprises a
zeolite of the MFI-type. The zeolite are preferably under their
acidic form or H-form MFI-type have a three-dimensional structure,
preferably the zeolites catalysts of the MFI-type are selected from
ZSM-5, silicalite-1. The preferred MFI-type zeolite is ZSM-5.
MEL-type have a three-dimensional structure, preferably the zeolite
of the MEL-type is selected from ZSM-11, silicalite-2, and SSZ-46.
The preferred zeolite of the MSE-type is MCM-68. The zeolite of the
CON-type is selected from CIT-1 and SSZ-33. The zeolite of the
FER-type is selected from Ferrierite, FU-9 and ZSM-35. The
preferred zeolite of the MTT-type is ZSM-23. The zeolite of the
TON-type is selected from ZSM-22, Theta-1 and NU-10. The zeolite of
the EUO-type is selected from ZSM-50 and EU-1. The preferred
zeolite of the MFS-type is ZSM-57. The preferred zeolite of the
AEL-type is SAPO-11. ZSM-48 refers to the family of microporous
materials consisting of silicon, aluminium, oxygen and optionally
boron. All the zeolite are preferably under their acidic form or
H-form
[0090] Preferably, the catalyst composition comprises one or more
zeolites catalysts having an acid 10-membered ring channels
selected from the list comprising ZSM-5, silicalite-1, ZSM-11,
silicalite-2, SSZ-46, MCM-68, CIT-1, SSZ-33, ZSM-8, Ferrierite,
FU-9, ZSM-35, ZSM-23, ZSM-22, Theta-1, NU-10, ZSM-50, EU-1, ZSM-57,
SAPO-11 and ZSM-48. More preferably the catalyst is or comprises
ZSM-5 zeolite. All the zeolite are preferably under their acidic
form or H-form
[0091] In a preferred embodiment, the catalyst composition
comprises 3D zeolite without cages (cavities) and containing at
least one 10-member rings.
[0092] Preferably, the catalyst composition comprises at least 60
wt % of one or more zeolite catalysts having an acid 10-membered
ring channels, more preferably at least 70 wt %, even more
preferably at least 80 wt % and most preferably at least 90 wt
%.
[0093] Several mesoporisation approaches may be used to create
zeolites crystals that contain both mesopores and micropores,
including destructive approaches such as demetallation
(desilication and dealumination) and recrystallization; and
constructive approaches such as using hard templates,
supramolecular templates, and surface silanization.
[0094] To provide sufficient acidity for the hydrocracking
reaction, it is preferred that the zeolites catalysts are at least
partly in their hydrogen form, e.g. H-ZSM-5 or H-ZSM-11. Preferably
at least 50 wt % of the total amount of the zeolites catalysts used
are in their hydrogen form, preferably at least 80 wt %, more
preferably at least 90 wt %, and even more preferably 100 wt % of
the zeolites catalysts are in their hydrogen form.
[0095] When the zeolites catalysts are prepared in the presence of
an organic cation, they may be activated by heating them in an
inert or oxidative atmosphere to remove the organic cation. For
example, they may be activated at a temperature over 500.degree. C.
for at least 1 hour.
[0096] In a preferred embodiment, the one or more zeolite catalysts
have a Si/Al molar ratio ranging from 10 to 100, preferably from 20
to 80, more preferably from 30 to 60.
[0097] In an embodiment, the zeolite is shaped with a binder. The
binder is an inorganic material.
[0098] Preferably, the binder is selected from silica,
alpha-alumina, gamma-alumina, clays, alumina phosphates, mullite,
zirconia, titania, yttria, silicon nitride, silicon carbide, iron,
bronze and stainless steel, glass, and carbon, preferably the
binder is alumina.
[0099] The zeolite shaped with the binder forms a catalyst
composition, and the catalyst composition of the present invention
preferably comprises at least 10 wt % of a binder as based on the
total weight of the catalyst composition, most preferably at least
20 wt % of a binder and preferably comprises up to 40 wt % of a
binder.
[0100] Optionally, the zeolite may be modified by phosphorous,
steaming, leaching, mesoporization, dealumination or modification
with alkali-earth or rare-earth metals.
[0101] In a further aspect, the catalyst composition used in the
process of the present invention is prepared by the method
comprising the steps of: [0102] steaming acid zeolite catalyst at a
temperature between 500.degree. C.-750.degree. C. for a period from
0.1 to 24 h under steam pressure from 0.1 to 10 bars; [0103]
optionally, contacting the steamed zeolite with the one or more
source of phosphorus; [0104] optionally, introducing to the
phosphate sample at least 0.1 wt % of Mg, Ca, Sr, Ba, Ce, La, Fe,
Ga; [0105] drying and steaming of the one or more modified acid
zeolite catalysts at a temperature between 500.degree.
C.-750.degree. C. for a period from 0.1 to 24 h under steam
pressure from 0.1 to 10 bars.
[0106] When at least 0.1 wt % of Mg, Ca, Sr, Ba, Ce, La, Fe, Ga is
introduced to the phosphate sample, the said metals may be further
presented on catalyst in form of oxides, silicates, aluminates or
phosphates.
[0107] Accordingly, the one or more acid zeolite catalysts are
contacted with a solution in which one or more basic compounds are
dissolved, and wherein, with preference, one or more redox elements
are dissolved as well. Preferably, the solution is an aqueous
solution. The preferred source of phosphorous is the phosphoric
acid. The preferred soluble salts of the basic compounds and of the
redox elements are nitrate salts. The preferred soluble salts of
the basic compounds are selected from the list consisting of
Mg(NO.sub.3).sub.2, Ca(NO.sub.3).sub.2, Sr(NO.sub.3).sub.2,
La(NO.sub.3).sub.3, Ga(NO.sub.3).sub.3, Fe(NO.sub.3).sub.3.
[0108] In an embodiment, the phosphorous modified acid zeolite
catalyst is preferably obtained by the process described in
WO2009/016156, which is incorporated herein by reference. The
process comprises the following steps in this order: [0109]
selecting a zeolite with low Si/Al molar ratio (advantageously
lower than 30) among H.sup.+ or NH.sub.4.sup.+-form of MFI, MEL,
FER, MOR, clinoptilolite, said zeolite having been made preferably
without direct addition of organic template; [0110] steaming at a
temperature ranging from 400 to 870.degree. C. for 0.01 to 200 h;
[0111] leaching with an aqueous acid solution containing the source
of P at conditions effective to remove a substantial part of Al
from the zeolite and to introduce at least 0.3 wt % of P; [0112]
separation of the solid from the liquid; [0113] an optional washing
step or an optional drying step or an optional drying step followed
by a washing step; [0114] a calcination step.
[0115] The basic compounds and the optional redox element(s) and
phosphorus (P), may be deposited by contacting the one or more acid
zeolite catalysts with a single solution in which the soluble salts
of the basic compounds, soluble salts of the redox elements and
phosphoric acid are dissolved.
[0116] Alternatively, the basic compounds and the optional redox
element(s) and phosphorus (P) may be deposited by subsequently
contacting the one or more acid zeolite catalysts with the
different elements and/or phosphorus, whereby the composition is
dried to evaporate the solvent before contacting the composition
with the following element. After depositing all the required
elements, the resulting composition (catalyst precursor) is
dried.
[0117] In one embodiment of the present invention, the catalyst
precursor is air-dried, preferably for about 8 hours at a
temperature ranging from 60.degree. C. to 80.degree. C. while
stirring.
[0118] After drying, the zeolite-comprising composition, on which
the basic compounds and the optional redox element(s) and the
phosphorus (P) are deposited, is calcined in an oxygen-comprising
atmosphere, preferably in a moisture-free atmospheric air.
Preferably, the catalyst precursor is calcined at a temperature
ranging from 450.degree. C. to 550.degree. C. to remove the
residual amount of nitrates and carbons.
[0119] Most preferably, the catalyst precursor is calcined at about
500.degree. C. for about 4 hours. When a binder is present, it is
preferred that the one or more acid zeolite catalysts are mixed
with the binder prior to contacting the one or more acid zeolite
catalysts with one or more solutions comprising soluble salts of
basic compounds and the optional soluble salts of redox elements
and phosphoric acid.
[0120] In the process of the invention, the naphtha feedstock is
contacted with the catalyst composition at elevated temperatures
and elevated pressures in hydrocracking conditions to perform a
hydrocracking step.
[0121] In an embodiment, the hydrocracking conditions comprise the
naphtha feedstock being contacted with the catalyst composition at
a temperature ranging from 200.degree. C. to 600.degree. C.,
preferably ranging from 250.degree. C. to 550.degree. C., more
preferably ranging from 300.degree. C. to 450.degree. C.
Preferably, the hydrocracking conditions comprise the naphtha
feedstock being contacted with the catalyst composition at a
pressure ranging from 1 to 10 MPa, preferably in the range of 2 to
6 MPa, more preferably from 3 to 5 MPa.
[0122] In an embodiment, the hydrocracking conditions comprise the
naphtha feedstock being contacted with the catalyst composition at
a weight hourly space velocity of the naphtha feedstock (VVHSV) of
at least 0.1 h.sup.-1, preferably is ranging from 0.1 h.sup.-1 to
10.0 h.sup.-1, more preferably from 0.5 h.sup.-1 to 8.0 h.sup.-1,
even more preferably from 1.0 h.sup.-1 to 6.0 h.sup.-1, and most
preferably from 1.5 h.sup.-1 to 5.0 h.sup.-1.
[0123] Hydrogen may be provided at any suitable ratio to the
paraffins contained in the naphtha feedstock. Preferably, the
hydrogen is provided in a molar ratio hydrogen to the paraffins in
the naphtha feedstock ranging from 1000:1 to 1:1, more preferably
from 100:1 to 1:1, even more preferably from 20:1 to 1:1; wherein
the number of moles of the paraffins in the naphtha feedstock is
based on the average molecular weight of the naphtha feedstock. The
process according to the invention is to achieve a set conversion
of the hydrocarbon naphtha feedstock. Preferably, the ratio of
hydrogen to the paraffins in the naphtha feedstock is chosen such
that the process conditions enable to achieve the desired
conversion.
[0124] PDH is a catalytic dehydrogenation process that converts
paraffins (in this case propane) to their corresponding light
olefins (propylene) dating back from the 1930s. In the late 1980s,
the chromia-alumina catalyst was specifically applied in the
dehydrogenation of propane to propylene. The conversion process is
favored under high temperature and low partial pressure of propane.
The reaction is run optimally at 500.degree. C.-700.degree. C. to
minimize thermal cracking, while the reaction pressure is typically
atmospheric. The product mixture goes through a deethaniser to
remove light hydrocarbons and traces of hydrogen. The last
separation step involves a propane-propylene splitter to achieve
polymer-grade propylene. It features high yields of 90 wt %
propylene, low generation of by-products and relatively low
investment and operating costs compared to steam cracking. Lummus
CATOFIN and UOP OLEFLEX are the most commonly licensed and proven
technology. Less common, at the global level, are the ThyssenKrupp
Udhe STAR process and Linde/BASF's PDH technology. Recently, Dow
Chemicals introduced a new fluidized bed propane dehydrogenation
technology to the market.
[0125] The catalytic dehydrogenation of propane is an endothermic
reaction, which produces propane and hydrogen. Hydrogen can be
reused for naphtha cracking. The extent of conversion is limited by
the thermodynamics of the reaction with higher temperature
favouring higher conversion. The selectivity decreases as the
conversion increases.
[0126] In U.S. Pat. No. 6,392,113 B1 assigned to ABB Lummus, the
performance and economics of a catalytic dehydrogenation process
are improved by using two pre-reactors before two main
dehydrogenation reactors (one in operation and the other in
regeneration). In this process, propane is preheated and then
passed to the pre-reactor, wherein the catalyst bed is not heated,
and propane is partially dehydrogenated with conversion at about
10-25% with the effluent leaves the reactor at 100.degree. C. The
pre-reactor can be operated for hours before it is subjected to
regeneration. The partially dehydrogenated reactor effluent is then
reheated by passing through a fired unit, where the heat is
supplied by a portion of the effluent air from the regeneration of
the main dehydrogenation reactor. The remainder of the effluent air
is used to regenerate the catalyst in the pre-reactor. The heated
effluent from the pre-reactor is then dehydrogenated in the main
dehydrogenation reactor. The operation of the pre-reactor is
comparatively steady than that of the main reactor, which permits
an extended cycle of about 24 hours before any catalyst
regeneration is required. In WO95/23123, the overall performance of
the dehydrogenation process can be improved with respect to the
dehydrogenation cycle and heating cycle by the change from
co-current to counter-current flow through the catalyst bed. In the
process, the regeneration gas is introduced at the opposite of the
reactor from the feed hydrocarbon. This provides the highest
temperature at the end of the reactor, thus creating the most
favourable conditions for the dehydrogenation reaction with respect
to equilibrium. The reaction is carried at 590.degree. C. and 0.05
MPa in the presence of chromium catalyst supported on alumina with
a conversion of propane at 47% and selectivity to propylene at
87%.
[0127] In the UOP Oleflex process, the dehydrogenation reaction is
carried out at about 600.degree. C. and 0.1 MPa in four sequential
moving-bed reactors with interheaters to reheat the reactor
effluent to the desired reaction temperature before passing it to
the next catalyst bed (U.S. Pat. No. 5,321,192). Hydrogen and other
inert compounds may be added to the feed stream to the
dehydrogenation reactor. The catalyst comprises 0.7-0.75 wt % Pt,
0.5 wt % Sn, and 3.5-4.4 wt % alkali metal supported on
.gamma.-alumina (EP 0448858 B1, U.S. Pat. No. 5,457,256). It passes
through annular bed via gravity flow (U.S. Pat. Nos. 5,227,567,
5,177,293). The reactor effluent is compressed and dried before it
passes to a cryogenic operation system (a cold box) where
hydrogen-methane gas is removed from C2+ hydrocarbon compounds. The
gas is subjected to adsorption with an adsorbent, such as alumina.
Silica gel, active carbon, or molecular sieves are used to remove
methane from the hydrogen gas (U.S. Pat. No. 5,457,256). In U.S.
Pat. No. 6,293,999 assigned to UOP, a pressure swing adsorption
(PSA) process is used for the separation of a hydrocarbon feed gas
comprising propylene and propane into a fraction, which comprises
predominantly propylene and a fraction comprising propane. The
process uses a small pore aluminophosphate molecular sieve,
ALPO-14, to selectively adsorb propylene while essentially
excluding propane at an adsorption temperature of 70-120.degree. C.
and at a propylene partial pressure of 0.05-0.2 MPa. The desorption
conditions are conducted under a temperature of 70-120.degree. C.
and a propylene partial pressure of 0.001 and 0.05 MPa. Overall,
the process comprises the adsorption step, the series of connected
co-purge steps, and the counter-current depressurization and
repressurization step. In another UOP patent (U.S. Pat. No.
6,218,589), a selective hydrogenation is employed to treat a
mixture of the reactor effluent from a dehydrogenation and a
recycle stream from a downstream propane-propylene splitter, and to
convert the majority of highly unsaturated impurities, such as
methylacetylene and propadiene, to propane. The effluent from the
selective hydrogenation reactor is sent to a deethanizer before
being purified within a propylene-propane splitter.
[0128] In a preferred embodiment the step 19 of dehydrogenation of
propane 15 to propylene 23 (i.e. the PDH step) is performed with a
catalyst selected from: [0129] 0.7-0.75 wt % Pt, 0.5 wt % Sn, and
3.5-4.4 wt % alkali metal supported on .gamma.-alumina; [0130]
Cr.sub.2O.sub.3/Al.sub.2O.sub.3; [0131] Pt--Ga/Al.sub.2O.sub.3.
[0132] In a preferred embodiment, the PDH step 19 is performed at a
temperature ranging from 500 to 800.degree. C. and a partial
pressure of propylene below one atmosphere.
[0133] In a preferred embodiment, the PDH step is performed as
described in US2010/0236985. In an embodiment, the catalyst used in
the dehydrogenation step of propane to propylene comprises: [0134]
i. from 0.25 to 5.0 wt %, preferably 0.3 to 3.0 wt % of the first
component, preferably gallium, or a compound thereof; [0135] ii.
from 0.0005 to 0.05 wt %, preferably 0.0007 to 0.04 wt % of the
second component, preferably platinum, or a compound thereof;
[0136] iii. from 0.0 to 2.0 wt %, preferably 0.1 to 1.0 wt % of an
alkali metal or alkaline earth metal, preferably potassium; and
[0137] iv. a support comprising alumina in the gamma crystalline
form.
[0138] Methods
[0139] Gas chromatography was performed on Columns: DB1 (40 m, 0.1
mm, 0.4 .mu.m) and Al.sub.2O.sub.3 (50 m, 0.32 mm, 5 .mu.m) using
Agilent operated by ChemStation software.
[0140] The elemental composition (i.e. the metal content) of
catalyst composition can preferably be determined by ICP-OES
according for instance to the method UOP Method 961-12. The
platinum content can preferably be determined according to the
method ASTM D4642.
EXAMPLES
[0141] The following examples illustrate the invention.
Example 1
[0142] The process was conducted in a fixed bed reactor loaded with
a ZSM-5 (Si/Al-40, CBV8014 from Zeolyst INT) containing catalyst
extruded with an Al.sub.2O.sub.3 binder (80 wt % zeolite, 20 wt %
Al.sub.2O.sub.3) in form of cylinders 1.6 mm. After the extrusion,
the catalyst composition was dried at room temperature for 24 hours
followed by calcination at 550.degree. C. for 6 h. The
demonstration of the invention was performed in both micro
pilots.
[0143] A stainless-steel reactor tube having an internal diameter
of 10 mm is used. 10 ml of the catalyst composition, as pellets of
35-45 mesh, is loaded in the tubular reactor. The void spaces,
before and after the catalyst composition, are filled with SiC
granulates of 2 mm. The temperature profile is monitored with the
aid of a thermocouple placed inside the reactor at the top of the
catalyst bed. Prior to the reaction, the catalyst composition was
pretreated with an hydrogen flow at 400.degree. C. for 6 h (heating
rate 60.degree. C./h) followed by cooling down to the reaction
temperature. The performance test is performed down-flow at 4 MPa
of total pressure, molar H.sub.2/Naphtha of 13.2; WHSV (naphtha) of
2.5 h.sup.-1, temperature of 400.degree. C. Analysis of the
products is performed by using an on-line gas chromatography. The
results are provided in tables 1 to 3.
[0144] Feedstock Characteristics
TABLE-US-00001 TABLE 1 Naphtha feedstock, POINA analysis nPar iPar
Napht Arom Total 36.38 38.06 23.28 2.29 C1 0 0 0 0 C2 0 0 0 0 C3 0
0 0 0 C4 8.1 0.58 0 0 C5 15.51 14.38 3.9 0 C6 11.43 17.75 14.97
1.98 C7 1.34 5.35 4.41 0.3 C8 0 0 0 0 C9 0 0 0 0 C10 0 0 0 0 C11 0
0 0 0 >200 C. 0 0 0 0 >200 C. 0 0 0 0 >200 C. 0 0 0 0 P +
A
TABLE-US-00002 TABLE 2 The distillation cut (according to ASTM D
86) is given in below: DIST_86: Distillation ASTM D86 T.degree. C.
at IBP 28.9.degree. C. DIST_86: T.degree. C. at 5% vol 38.1.degree.
C. DIST_86: T.degree. C. at 10% vol 40.3.degree. C. DIST_86:
T.degree. C. at 20% vol 44.2.degree. C. DIST_86: T.degree. C. at
30% vol 48.4.degree. C. DIST_86: T.degree. C. at 40% vol
52.6.degree. C. DIST_86: T.degree. C. at 50% vol 57.1.degree. C.
DIST_86: T.degree. C. at 60% vol 61.6.degree. C. DIST_86: T.degree.
C. at 70% vol 66.6.degree. C. DIST_86: T.degree. C. at 80% vol
72.1.degree. C. DIST_86: T.degree. C. at 90% vol 79.4.degree. C.
DIST_86: T.degree. C. at 95% vol 89.9.degree. C. DIST_86: T.degree.
C. at FBP 91.3.degree. C. DIST_86: % Recovered at 76.2% vol
70.degree. C. (*) DIST_86: Recovered % vol 97.0% vol DIST_86:
Residue % vol 0.5% vol DIST_86: Loss % vol 2.5% vol DIST_86: (*) %
Loss included N Yes/No
[0145] The density was determined according to ISO 12185 at
15.degree. C. and was found to be 0.6702 g/ml.
TABLE-US-00003 TABLE 3 FEED EFFLUENT Effluent with C4 (in wt %) (in
wt %) recycling (in wt %) Paraffins 74.4 92.8 91.4 Cy-Paraffins
23.3 0.3 0.1 Olefins 0.00 0.28 0.4 Aromatics 2.3 6.6 8.1 Breakdown
Methane 0 1.7 2.3 Ethane 0 4.7 6.3 Propane 0 38.5 51.6 i-Butane
0.58 13.0 0.2 n-Butane 8.1 12.1 0.2 C5+ 91.3 30.0 37.1
[0146] The process allows valorizing about 50% naphtha as propane
and about 37 wt % as gasoline, which shows a very high carbon
efficiency.
[0147] Taking into account an overall propylene yield in propane
dehydrogenation of about 82%, the process allows converting of
about 41 wt % of naphtha to propylene as opposed to 17 wt % in
steam cracking. This offers about 2.4 times higher propylene yield
from the same feedstock.
TABLE-US-00004 TABLE 4 RON of the C5+ fractions FEED Effluent RON
74 (71-77) 92 (89-93) Aromatic content, wt % 2.5 21.8 Olefins
content, wt % <0.1 <1.0
[0148] The C5+ produced in the reaction demonstrates very valuable
properties for the use of gasoline.
[0149] The catalyst composition was two times regenerated in the
air at 550.degree. C. and demonstrated full recovering of the
activity.
[0150] The process shows a low H.sub.2 consumption. Indeed, the
naphtha feedstock contained 15.7 wt % of hydrogen as based on the
total weight of the feed, whereas the weight content of hydrogen in
the effluent was 16.4 wt %.
Example 2
[0151] The experiment was repeated with compositions comprising
zeolites catalysts comprising 12-membered ring channels. They
showed fast deactivation and significantly lower yield of propane,
as shown by the results given in table 5.
[0152] Conditions: pressure: 4 MPa [0153] Molar H.sub.2/Naphtha:
13.2 [0154] WHSV (naphtha): 2.5 h-1 [0155] Temperature: 450.degree.
C.
TABLE-US-00005 [0155] TABLE 5 FEED Composition Catalyst of the feed
MOR ZSM-5 BETA ZSM-12 SI/AI 10 40 40 40 Structure 1D 3D 3D 1D
Members ring 8, 12 10 12 12 Composition of the effluent Paraffins
84 78 92.5 79 83 Cy-paraffins 13.23 1.15 0.05 6.76 0.30 Olefins
0.00 8.61 0.92 2.69 0.93 aromatics 2.9 11.8 6.5 11.2 16.0
Conversion to C5- Total -- 59.8 83.9 15.6 60.0 Propane 0.0 27.4
43.9 5.8 32.0 Ethane 0.0 7.4 12.7 0.6 3.4 Iso-butane 0.1 11.8 11.6
5.3 12.4 N-butane 2.3 7.6 9.0 2.8 9.4 Conversion to C5+ Total 100
40.2 16.1 84.4 40.0 % aromatics 2.9 26.9 40.6 13.3 39.1 in C5+ BTX
(in arom) 100 81.3 89.0 75.9 76.5
[0156] Many zeolites catalysts deactivate in the presence of
aromatics in the feed. ZSM-5 shows stable performances.
[0157] The results demonstrate a higher selectivity for propane
than for ethane for all the zeolite catalysts tested. However, best
results were achieved with ZSM-5.
* * * * *