Molecular sieve abstract
Hydrocarbon conversion process in which a feedstock is contacted
with a non zeolitic molecular sieve to remove most, if not all,
of the halogen contained in the catalyst. The halogen may be removed
by one of several methods. One method includes heating the catalyst
in a low moisture environment, followed by contacting the heated
catalyst with air and/or steam. Another method includes steam-treating
the catalyst at a temperature from 400.degree. C. to 1000.degree.
C. The hydrocarbon conversion processes include the conversion of
oxygenates to olefins, the conversion of oxygenates and ammonia
to alkylamines, the conversion of oxygenates and aromatic compounds
to alkylated aromatic compounds, cracking and dewaxing.
Molecular sieve claims
1. A hydrocarbon conversion process comprising the steps of: (a)
introducing a feedstock to a reactor system in the presence of a
catalyst comprising a non zeolitic molecular sieve, inorganic oxide
matrix, and matrix material, wherein the catalyst contains from
10 ppm to 600 ppm by weight halogen; (b) withdrawing from the reactor
system an effluent stream; and (c) passing the effluent gas through
a recovery system recovering at least the one or more conversion
products.
2. The process of claim 1 wherein the feedstock comprises one or
more oxygenates.
3. The process of claim 2 wherein the one or more oxygenates comprises
methanol.
4. The process of claim 1 wherein the one or more conversion products
comprises one or more olefins.
5. The process of claim 4 wherein the one or more olefins comprises
ethylene, propylene and mixtures thereof.
6. The process of claim 1 wherein the feedstock comprises one or
more oxygenates and ammonia.
7. The process of claim 6 wherein the one or more conversion products
comprises one or more alkylamines.
8. The process of claim 7 wherein the one or more alkylamines comprises
one or more methylamines.
9. The process of claim 6 wherein the one or more oxygenates comprises
methanol.
10. The process of claim 1 wherein the feedstock comprises one
or more oxygenates and one or more aromatic compound.
11. The process of claim 10 wherein the one or more conversion
products comprises one or more alkylated aromatic compound.
12. The process of claim 1 wherein the conversion process is cracking.
13. The process of claim 1 wherein the conversion process is dewaxing.
14. The process of claim 1 wherein the halogen is chlorine.
15. The process of claim 14 wherein the catalyst contains from
10 ppmw to 400 ppmw chlorine.
16. The process of claim 14 wherein the catalyst contains from
10 ppmw to 200 ppmw chlorine.
17. The process of claim 15 wherein the catalyst contains from
10 ppmw to 80 ppmw chlorine.
18. The process of claim 1 wherein the catalyst has a GAL Index
of less than 5.
19. The process of claim 18 wherein the GAL Index is less than
3.
20. The process of claim 1 wherein the non-zeolitic molecular sieve
is selected from SAPO-5 SAPO-8 SAPO-11 SAPO-16 SAPO-17 SAPO-18
SAPO-20 SAPO-31 SAPO-34 SAPO-35 SAPO-36 SAPO-37 SAPO-40 SAPO-41
SAPO-42 SAPO-44 SAPO-47 SAPO-56 ALPO-5 ALPO-11 ALPO-18 ALPO-31
ALPO-34 ALPO-36 ALPO-37 ALPO-46 the metal containing forms of
each thereof, and mixtures thereof.
21. The process of claim 1 wherein the catalyst comprises 20% to
45% by weight non-zeolitic molecular sieve, 5% to 20% by weight
of inorganic oxide matrix, and 20% to 70% by weight matrix material.
22. The process of claim 21 wherein the catalyst comprises 25%
to 42% by weight non-zeolitic molecular sieve, 8% to 15% by weight
of inorganic oxide matrix, and 40% to 60% by weight matrix material.
23. The process of claim 19 wherein the GAL Index is less than
2.
24. The process of claim 1 wherein the inorganic oxide matrix comprises
an aluminum oxide matrix.
25. The process of claim 2 wherein the source of chlorine is aluminum
chlohydrol.
26. A hydrocarbon conversion process comprising the steps of: (a)
providing a halogen-containing non-zeolitic molecular sieve catalyst;
(b) heating the catalyst in a low moisture environment at a temperature
from 400.degree. C. to 1000.degree. C.; (c) contacting the heated
catalyst with steam at a temperature from 400.degree. C. to 1000.degree.
C. to produce a steam-treated catalyst; (d) introducing a feedstock
to a reactor system in the presence of the steam-treated catalyst
obtained at step (c); (e) withdrawing from the reactor system an
effluent stream; and (f) passing the effluent gas through a recovery
system recovering at least the one or more conversion products.
27. The process of claim 26 wherein the halogen is chlorine.
28. The process of claim 26 wherein the feedstock comprises one
or more oxygenates.
29. The process of claim 28 wherein the one or more oxygenates
comprises methanol.
30. The process of claim 26 wherein the one or more conversion
products comprises one or more olefins.
31. The process of claim 30 wherein the one or more olefins comprises
ethylene, propylene and mixtures thereof.
32. The process of claim 26 wherein the feedstock comprises one
or more oxygenates and ammonia.
33. The process of claim 32 wherein the one or more conversion
products comprises one or more alkylamines.
34. The process of claim 33 wherein the one or more alkylamines
comprises one or more methylamines.
35. The process of claim 32 wherein the one or more oxygenates
comprises methanol.
36. The process of claim 26 wherein the feedstock comprises one
or more oxygenates and one or more aromatic compound.
37. The process of claim 36 wherein the one or more conversion
products comprises one or more alkylated aromatic compound.
38. The process of claim 26 wherein the conversion process is cracking.
39. The process of claim 26 wherein the conversion process is dewaxing.
40. A hydrocarbon conversion process comprising the steps of: (a)
providing a halogen-containing non-zeolitic molecular sieve catalyst;
(b) heating the catalyst in an oxygen environment at a temperature
from 400.degree. C. to 1000.degree. C.; (c) contacting the heated
catalyst with steam at a temperature from 400.degree. C. to 1000.degree.
C. to produce a steam-treated catalyst; (d) introducing a feedstock
to a reactor system in the presence of the steam-treated catalyst
obtained at step (c); (e) withdrawing from the reactor system an
effluent stream; and (f) passing the effluent gas through a recovery
system recovering at least the one or more conversion products.
41. The process of claim 40 wherein the halogen is chlorine.
42. The process of claim 40 wherein the feedstock comprises one
or more oxygenates.
43. The process of claim 42 wherein the one or more oxygenates
comprises methanol.
44. The process of claim 40 wherein the one or more conversion
products comprises one or more olefins.
45. The process of claim 44 wherein the one or more olefins comprises
ethylene, propylene and mixtures thereof.
46. The process of claim 40 wherein the feedstock comprises one
or more oxygenates and ammonia.
47. The process of claim 46 wherein the one or more conversion
products comprises one or more alkylamines.
48. The process of claim 47 wherein the one or more alkylamines
comprises one or more methylamines.
49. The process of claim 46 wherein the one or more oxygenates
comprises methanol.
50. The process of claim 40 wherein the feedstock comprises one
or more oxygenates and one or more aromatic compound.
51. The process of claim 50 wherein the one or more conversion
products comprises one or more alkylated aromatic compound.
52. The process of claim 40 wherein the conversion process is cracking.
53. The process of claim 40 wherein the conversion process is dewaxing.
54. A hydrocarbon conversion process comprising the steps of: (a)
providing a halogen-containing non-zeolitic molecular sieve catalyst;
(b) heating the catalyst in an environment containing steam at a
temperature from 400.degree. C. to 1000.degree. C.; (c) removing
from 70% to 99.99% by weight of the halogen from the catalyst, thereby
producing a steam-treated catalyst; (d) introducing a feedstock
to a reactor system in the presence of the steam-treated catalyst
obtained at step (c); (e) withdrawing from the reactor system an
effluent stream; and (f) passing the effluent gas through a recovery
system recovering at least the one or more conversion products.
55. The process of claim 54 wherein the halogen is chlorine.
56. The process of claim 54 wherein the feedstock comprises one
or more oxygenates.
57. The process of claim 56 wherein the one or more oxygenates
comprises methanol.
58. The process of claim 54 wherein the one or more conversion
products comprises one or more olefins.
59. The process of claim 58 wherein the one or more olefins comprises
ethylene, propylene and mixtures thereof.
60. The process of claim 54 wherein the feedstock comprises one
or more oxygenates and ammonia.
61. The process of claim 60 wherein the one or more conversion
products comprises one or more alkylamines.
62. The process of claim 61 wherein the one or more alkylamines
comprises one or more methylamines.
63. The process of claim 60 wherein the one or more oxygenates
comprises methanol.
64. The process of claim 54 wherein the feedstock comprises one
or more oxygenates and one or more aromatic compound.
65. The process of claim 64 wherein the one or more conversion
products comprises one or more alkylated aromatic compound.
66. The process of claim 54 wherein the conversion process is cracking.
67. The process of claim 54 wherein the conversion process is dewaxing.
Molecular sieve description
[0001] This application claims priority to U.S. patent application
Ser. No. 09/891674 filed June 2001 the entire disclosure of which
is incorporated herein by reference.
FIELD OF THE INVENTION
[0002] The present invention relates to methods of removing halogen
from non-zeolitic, molecular sieve catalysts, the catalysts produced
from such methods, and the use of such catalysts in hydrocarbon
conversion processes.
BACKGROUND OF THE INVENTION
[0003] Molecular sieve catalysts used in a fluidized-bed reactor
or a riser reactor will typically have an average particle diameter
from 40 .mu.m to 300 .mu.m. Catalyst particle size within this range
is needed for proper fluidization as well as to efficiently separate
the catalyst from the gaseous products in a cyclone separator. To
maintain the desired catalyst diameter the molecular sieve is formulated
with other materials. Dilution of the molecular sieve with these
materials is also used to control the rate of reaction, control
the temperature of the reactor and regenerator, and to stabilize
and protect the molecular sieve.
[0004] Formulated molecular sieve catalysts present a problem not
found in other types of industrial catalysts, that is, how to maintain
the physical integrity of the molecular sieve catalyst during the
fluidized cyclic process of reaction, separation, and regeneration.
The cycles of reaction, separation, and regeneration are carried
out at high temperatures and high flow rates. Collisions and abrasions
between catalyst particles, between the catalyst particles and reactor
walls and between the catalyst particles and other parts of the
unit tend to cause physical breakdown of the original catalyst into
smaller catalyst particles known as fines. This physical breakdown
is referred to as catalyst attrition. The fines usually have particle
diameters smaller than 20 microns--much smaller than the original
catalyst particles. Catalysts with higher attrition resistance are
desirable because, among other reasons, fewer fines are generated
for disposal, less environmental impact is caused by unrecoverable
airborne particulates, optimal fluidized conditions are maintained,
operating costs are lower, and less replacement catalyst is required.
[0005] Molecular sieve catalysts are formed by various methods,
for example, by spray drying or extruding a slurry containing the
molecular sieve and the other catalyst components. The catalysts
are formed by mixing the zeolitic molecular sieve with one or more
binding agents such as one or more types of alumina and/or silica.
Matrix materials, typically clays, are also added and serve as diluents
to control the rate of the catalytic reaction, and to facilitate
heat transfer during many stages of the process. In U.S. Pat. No.
5346875 to Wachter et al. zeolite-Y (21.8 wt %) is mixed with
Kaolin clay (14.5 wt %), silica sol (48.3 wt %), and Reheis chlorhydrol
(15.4 wt %) to form a slurry which is then spray dried and calcined.
A conventional calcination procedure was used; heating at 550.degree.
C. in air for 2 hours.
[0006] Non-zeolitic, molecular sieve catalysts are known to convert
oxygenates, particularly methanol, to light olefins. The oxygenate
to olefin process includes separate processing zones for conducting
the catalytic reaction, product-catalyst separation, and catalyst
regeneration. The produced olefin and other hydrocarbon products
are separated from the catalyst particles in a separator, suitably
a cyclone separator. A portion of the catalyst is recovered from
the separator and passed to a regenerator. In the regenerator the
non-zeolitic molecular sieve catalyst contacts a combusting gas,
e.g. air, at a temperature sufficient to burn off carbon deposits,
commonly referred to a coke, that accumulate on the surface and
in the pores of the catalyst. The regenerated catalyst is then returned
to the oxygenate conversion reactor.
[0007] In this process, the non-zeolitic molecular sieve catalyst
is subjected to great mechanical stresses. As the catalyst is transferred
from the reaction zone to cyclone separators, to regenerators, and
finally back to the reaction zone the catalyst will tend to disintegrate
into catalyst fines. These catalyst fines must be removed from the
reactor process and discarded. No matter how resistant the catalyst
is to attrition, eventually the oxygenate to olefin process will
break down the non-zeolitic molecular sieve catalyst because the
catalyst moves through the system at such high speeds. The resistance
of the catalyst to attrition is an important property of the catalyst.
[0008] In PCT Publication No. WO 99/21651 to Wachter et al. and
U.S. Pat. No. 4973792 to Lewis et al., silicoaluminophosphate
(SAPO) molecular sieve catalysts were produced by preparing a slurry
containing SAPO-34 Kaolin clay, and Reheis chlorhydrol. The slurry
was then directed to a spray dryer to form catalyst particles with
the desired size. The spray dried catalysts were calcined, however
the conditions of the calcination were stated to be not critical.
[0009] In U.S. Pat. Nos. 5248647 and 5095163 to Barger et al.
SAPO molecular sieve is mixed with an aqueous silica sol and spray
dried. The spray dried catalyst is mixed with an aqueous solution
of ammonium sulfate at 60.degree. C. three times, then washed with
water and dried at 100.degree. C. The dried, ion-exchanged catalyst
is then calcined in air at 550.degree. C. for over 3.3 hours and
then the temperature is lowered to ambient room temperature over
a period of 2 hours. A portion of this catalyst is then contacted
with steam at 725.degree. C. or 750.degree. C. for 10 hours. Steam
treatment following calcination is shown to increase catalyst life,
increase selectivity to ethylene and propylene, and decrease selectivity
to propane.
[0010] If SAPO molecular sieve catalysts are ever going to be used
commercially to convert oxygenates to olefins, catalysts with greater
attrition properties are needed. For this reason, the Applicants'
sought to develop SAPO catalysts with a relatively high resistance
to attrition.
SUMMARY OF THE INVENTION
[0011] The present invention is directed to methods of removing
a portion of the halogen present in non-zeolitic molecular sieve
catalysts. One embodiment of removing halogen includes heating the
catalyst in a low moisture environment at a temperature from about
400.degree. C. to about 1000.degree., and contacting the heated
catalyst with steam at a temperature from about 400.degree. C. to
about 1000.degree. C. to produce a steam-treated catalyst. Preferably,
the low moisture environment contains less than 5% by volume, more
preferably less than 1% by volume, water. The steam treatment can
take place in an oxygen environment. Also, it is preferred that
the steam treatment take place in an environment containing at least
10% by volume water. In the preferred embodiment, the steam treatment
can remove from about 50% to about 99% by weight, more preferably
from about 90% to about 99% by weight, of halogen from the heated
catalyst. The method can be used to remove halogen from silicoaluminophosphate
and/or aluminophosphate molecular sieve selected from the group
consisting of SAPO-5 SAPO-8 SAPO-11 SAPO-16 SAPO-17
[0012] SAPO-18 SAPO-20 SAPO-31 SAPO-34 SAPO-35 SAPO-36 SAPO-37
SAPO-40 SAPO-41 SAPO-42 SAPO-44 SAPO-47 SAPO-56 ALPO-5 ALPO-11
ALPO-18 ALPO-31 ALPO-34 ALPO-36 ALPO-37 ALPO-46 the metal
containing forms of each thereof, or mixtures thereof.
[0013] In another embodiment, a portion of the halogen can be removed
from a non zeolitic molecular sieve catalyst by heating the catalyst
in an oxygen environment at a temperature from about 400.degree.
C. to about 1000.degree. C. to produce a heated catalyst, and contacting
the heated catalyst with steam at a temperature from about 400.degree.
C. to about 1000.degree. C. Preferably, the oxygen environment contains
greater than about 10% by volume oxygen. It is also preferred, that
the steam treatment take place in an environment containing at least
about 10% by volume water. In many cases, the halogen to be removed
will be chlorine, and preferably from about 70% to about 99% by
weight, more preferably from about 90% to about 99% by weight, of
the chlorine will be removed from the heated catalyst.
[0014] In another embodiment, a portion of the halogen can be removed
from a non zeolitic molecular sieve catalyst by calcining the catalyst
in an environment containing steam at a temperature from about 400.degree.
C. to about 1000.degree. C., preferably from about 500.degree. C.
to about 800.degree. C., and more preferably from about 550.degree.
C. to about 700.degree. C., to remove from about 70% to about 99.99%
by weight of the halogen from the catalyst. If the halogen to be
removed from the catalyst is chlorine, the likely source of the
chlorine is aluminum chlorhydrol that is used to produce the catalyst.
The environment can contain from 5% to about 10% by volume water,
or at least 10% by volume, water. The environment can further contain
air, nitrogen, helium, flue gas, or any combination thereof.
[0015] In one embodiment, the catalyst is heated in a low-moisture
environment at a temperature of from about 400.degree. C. to about
1000.degree. C. to remove at least about 50% by weight of the halogen
in the catalyst prior to steam treatment. Preferably, the low moisture
environment contains less than about 5% by volume, more preferably
less then about 1% by volume, water. Also, the steam-treated catalyst
can optionally be heated in an oxygen environment that contains
greater than about 5% by volume oxygen.
[0016] In another embodiment, a portion of the halogen can be removed
from a silicoaluminophosphate molecular sieve catalyst by heating
the catalyst in a low moisture environment at a temperature from
400.degree. C. to about 1000.degree. C. to remove at least about
50% by weight of the chlorine from the catalyst, followed by contacting
the heated catalyst in a second calcination environment containing
about 5% to about 10% by volume water at a temperature from 400.degree.
C. to about 1000.degree. C. Preferably, the low moisture environment
contains less than about 1% by volume water.
[0017] The invention is also directed to a catalyst containing
a non zeolitic molecular sieve, inorganic oxide matrix, and matrix
material, wherein the catalyst contains from about 10 ppmw to about
600 ppmw by weight halogen. Generally, the halogen is chlorine,
and the catalyst will contain from about 10 ppmw to about 200 ppmw,
preferably from about 10 ppmw to about 80 ppmw, chlorine. It is
also preferred that the catalyst have a GAL Index of less than about
5 more preferably less than about 3 most preferably less than
about 2. The non-zeolitic molecular sieve in the catalyst is preferably
selected from SAPO-5 SAPO-8 SAPO-11 SAPO-16 SAPO-17 SAPO-18
SAPO-20 SAPO-31 SAPO-34 SAPO-35 SAPO-36 SAPO-37 SAPO-40 SAPO-41
SAPO-42 SAPO-44 SAPO-47 SAPO-56 ALPO-5 ALPO-11 ALPO-18 ALPO-31
ALPO-34 ALPO-36 ALPO-37 ALPO-46 the metal containing forms of
each thereof, or mixtures thereof. Preferably, the catalyst contains
about 20% to about 45% by weight, more preferably from about 25%
to about 42% by weight, non-zeolitic molecular sieve, about 5% to
about 20% by weight, more preferably about 8% to about 15% by weight,
of inorganic oxide matrix, and about 20% to about 70% by weight,
more preferably from about 40% to about 60% by weight, matrix material.
In the preferred embodiment, the inorganic oxide matrix contains
an aluminum oxide matrix that is formed from the heat treatment
of aluminum chlorhydrol.
[0018] The present invention also relates to hydrocarbon conversion
processes in which a feedstock is contacted with a non zeolitic
catalyst from which halogen has been removed. More specifically,
the present invention relates to a hydrocarbon conversion process
comprising the steps of (a) introducing a feedstock to a reactor
system in the presence of a catalyst comprising a non zeolitic molecular
sieve, inorganic oxide matrix, and matrix material, wherein the
catalyst contains from 10 ppm to 600 ppm by weight halogen; (b)
withdrawing from the reactor system an effluent stream; and (c)
passing the effluent gas through a recovery system recovering at
least the one or more conversion products.
[0019] In another embodiment, the present invention relates to
a hydrocarbon conversion process comprising the steps of: (a) providing
a halogen-containing non-zeolitic molecular sieve catalyst; (b)
heating the catalyst in a low moisture environment at a temperature
from 400.degree. C. to 1000.degree. C.; (c) contacting the heated
catalyst with steam at a temperature from 400.degree. C. to 1000.degree.
C. to produce a steam-treated catalyst; (d) introducing a feedstock
to a reactor system in the presence of the steam-treated catalyst
obtained at step (c); (e) withdrawing from the reactor system an
effluent stream; and (f) passing the effluent gas through a recovery
system recovering at least the one or more conversion products.
[0020] In another embodiment, the present invention relates to
a hydrocarbon conversion process comprising the steps of (a) providing
a halogen-containing non-zeolitic molecular sieve catalyst; (b)
heating the catalyst in an oxygen environment at a temperature from
400.degree. C. to 1000.degree. C.; (c) contacting the heated catalyst
with steam at a temperature from 400.degree. C. to 1000.degree.
C. to produce a steam-treated catalyst; (d) introducing a feedstock
to a reactor system in the presence of the steam-treated catalyst
obtained at step (c); (e) withdrawing from the reactor system an
effluent stream; and (f) passing the effluent gas through a recovery
system recovering at least the one or more conversion products.
[0021] In a further embodiment, the present invention relates to
a hydrocarbon conversion process comprising the steps of (a) providing
a halogen-containing non-zeolitic molecular sieve catalyst; (b)
heating the catalyst in an environment containing steam at a temperature
from 400.degree. C. to 1000.degree. C.; (c) removing from 70% to
99.99% by weight of the halogen from the catalyst, thereby producing
a steam-treated catalyst; (d) introducing a feedstock to a reactor
system in the presence of the steam-treated catalyst obtained at
step (c); (e) withdrawing from the reactor system an effluent stream;
and (f) passing the effluent gas through a recovery system recovering
at least the one or more conversion products.
[0022] In each of the preceding embodiments, the present invention
is applicable to a wide range of processes including those in which
the feedstock comprises one or more oxygenates, ammonia, aromatic
compounds, or mixtures thereof, which are converted to olefins,
alkylamines or alkylated aromatic compounds. The invention is also
applicable for feedstock cracking and dewaxing.
BRIEF DESCRIPTION OF THE DRAWING
[0023] The present invention will be better understood by reference
to the Detailed Description of the Invention when taken together
with the attached drawing, wherein FIG. 1 is a schematic representation
of one embodiment for removing chlorine from a formed catalyst.
DETAILED DESCRIPTION OF THE INVENTION
[0024] To produce non-zeolitic molecular sieve catalyst with a
relatively high resistance to attrition, an inorganic oxide sol
that contains halogen can be used. A preferred route to produce
non-zeolitic molecular sieve catalyst is to use an alumina sol that
contains chlorine, more preferably aluminum chlorhydrol, as a binder.
The inorganic oxide sol functions as a "glue" which binds
the catalyst components together. However, using an inorganic oxide
sol that contains halogen presents a problem not associated with
the use of halogen-free binders. A portion of the halogen from the
inorganic oxide sol remains in the formed catalyst. It is desirable
to remove most, if not nearly all, of the halogen from the catalyst
before the catalyst is used in the oxygenate to olefin process.
If most of the halogen is not removed from the catalyst, halogen-containing
acids will form in the oxygenate to olefin reactor. Over time, the
released acid will corrode the oxygenate to olefin reactor and other
process units. While the invention will be further illustrated for
the case where the halogen is chlorine, it should be understood
that the invention applies to other halogens as well, such as fluorine,
bromine and iodine. In the case of a catalyst containing chlorine,
hydrochloric acid will form in the oxygenate to olefin reactor.
HCl may be in the gas or condensed form, usually in a hydrated form,
hereinafter referred to as HCl.sub.(aq). All forms of acids are
potentially corrosive, the hydrated form being the most corrosive.
[0025] The invention addresses the problem associated with the
use of inorganic oxide sols that contains halogen by removing much
of the halogen from the catalyst during calcination of the catalyst.
The invention addresses these problems by providing methods of heat
treating or calcining a formed non-zeolitic molecular sieve catalyst
prepared with an inorganic oxide sol that contains halogen. The
methods of the invention minimize the production of halogen-containing
acids, or at least confines much of the produced halogen-containing
acids to a single heating or calcination unit that can be designed
to accommodate the corrosive effects of halogen-containing acids.
The methods of the invention also reduce the amount of halogen remaining
in the catalyst over that of conventional procedures.
[0026] The catalyst is made by preparing a slurry containing non-zeolitic
molecular sieve, an inorganic oxide binder, and a matrix material.
The slurry is then dried and shaped in a forming unit. Preferably,
the slurry is spray dried, and a dry powder catalyst with an average
catalyst particle size is obtained. The formed catalyst is then
heat treated, i.e., calcined.
[0027] Calcination is used to remove the template molecule from
the cage structure of the framework. During calcination all or part
of the template molecule exits the cage structure. Calcination is
also used to harden the formed catalyst particle. The relatively
high temperatures used during calcination transform the inorganic
oxide sol to an inorganic oxide matrix. It is this inorganic oxide
matrix that increases the attrition resistance of the catalyst particle.
[0028] If a conventional calcination procedure is used on a catalyst
containing chlorine, that is, heating in air at temperatures greater
than 400.degree. C., large amounts of HCl are produced over time
in the calcination unit. The formation of HCl.sub.(aq) is the result
of small amounts of water or water vapor contained in the air and
the water generated from the oxidative combustion of the organic
template during calcination. The released HCl, if not accounted
for, will eventually corrode the heating or calcination unit. Therefore,
it is desirable to control the removal of chlorine from the catalyst
in a manner that will either minimize the amount of HCl produced
during the calcination process or limit the evolution of HCl to
a single calcination unit.
[0029] A conventional calcination procedure also does not remove
enough of the halogen from the catalyst. In the case of chlorine,
the remaining chlorine in the catalyst is then released into the
oxygenate to olefin reactor and other oxygenate to olefin process
units as HCl.sub.(aq) due to the hydrothermal conditions of the
oxygenate to olefin process. If not accounted for, the release of
this HCl.sub.(aq) will corrode the oxygenate to olefin process units.
The presence of HCl.sub.(aq) in the olefin monomer feed used for
polymerization might also damage or poison expensive polymerization
catalysts. Therefore, it is desirable to remove as much chlorine
from the catalyst during the calcination process so as to minimize
the amount of HCl.sub.(aq) released into the oxygenate to olefin
process units.
[0030] As a result of using the calcination methods of the invention,
a non-zeolitic molecular sieve catalyst with low amounts of halogen
is obtained. A preferred catalyst of the invention contains a SAPO
molecular sieve, an aluminum oxide matrix, and clay, most preferably
Kaolin. The catalyst will also contain some halogen resulting from
the use of a binder that contains halogen. Although the invention
is directed to removing as much halogen from the catalyst as efficiently
possible, some of the halogen is not removed during the calcination
process. Following the calcination procedures of the invention,
the catalyst will contain from about 10 ppmw to 600 ppmw halogen,
preferably from about 10 ppmw to 200 ppmw halogen, more preferably
from about 10 ppmw to 60 ppmw halogen. The catalyst will also have
a Gross Attrition Loss (GAL) Index of less than 5 preferably a
GAL Index less than 3 more preferably a GAL Index less than 2.
The smaller the GAL Index, the more resistant to attrition is the
catalyst.
[0031] Non Zeolitic Molecular Sieve
[0032] The catalyst used according to the present invention contains
a non zeolitic molecular sieve. Examples of suitable non-zeolitic
molecular sieves are silicoaluminophosphates (SAPOs) and aluminophosphates
(ALPOs). In general, SAPO molecular sieves comprise a molecular
framework of corner-sharing [SiO.sub.4], [AlO.sub.4], and [PO.sub.4]
tetrahedral units. The [PO.sub.4] tetrahedral units are provided
by a variety of compositions. Examples of these phosphorus-containing
compositions include phosphoric acid, organic phosphates such as
triethyl phosphate, and aluminophosphates. The [AlO.sub.4] tetrahedral
units are provided by a variety of compositions. Examples of these
aluminum-containing compositions include aluminum alkoxides such
as aluminum isopropoxide, aluminum phosphates, aluminum hydroxide,
sodium aluminate, and pseudoboehmite. The [SiO.sub.4] tetrahedral
units are provided by a variety of compositions. Examples of these
silicon-containing compositions include silica sols and silicium
alkoxides such as tetra ethyl orthosilicate. The phosphorus-, aluminum-,
and silicon-containing compositions are mixed with water and a template
molecule and heated under appropriate conditions to form the molecular
sieve.
[0033] SAPO molecular sieves are generally classified as being
microporous materials having 8 10 or 12 membered ring structures.
These ring structures can have an average pore size ranging from
about 3.5-15 angstroms. Preferred are the small pore SAPO molecular
sieves having an average pore size of less than about 5 angstroms,
preferably an average pore size ranging from about 3.5 to 5 angstroms,
more preferably from 3.5 to 4.2 angstroms. These pore sizes are
typical of molecular sieves having 8 membered rings.
[0034] An aluminophosphate (ALPO) molecular sieve can also be included
in the catalyst composition. Aluminophosphate molecular sieves are
crystalline microporous oxides which can have an AlPO.sub.4 framework.
They can have additional elements within the framework, typically
have uniform pore dimensions ranging from about 3 Angstroms to about
10 Angstroms, and are capable of molecular size selective separations
of molecular species. More than two dozen structure types have been
reported, including zeolite topological analogues.
[0035] For a catalyst used in the conversion of oxygenate to light
olefin the non-zeolitic molecular sieve will have a relatively low
Si/Al.sub.2 ratio. In general, for SAPOs, a Si/Al.sub.2 ratio of
less than 0.65 is desirable, with a Si/Al.sub.2 ratio of not greater
than 0.40 being preferred, and a Si/Al.sub.2 ratio of not greater
than 0.32 being particularly preferred. A Si/Al.sub.2 ratio of not
greater than 0.20 is most preferred.
[0036] Substituted SAPOs and ALPOs can also be used in this invention.
These compounds are generally known as MeAPSOs, MeAPOs, metal-containing
silicoaluminophosphates or metal-containing aluminophosphates. The
metal can be alkali metal ions (Group IA), alkaline earth metal
ions (Group IIA), rare earth ions (Group IIIB, including the lanthanide
elements, and the additional transition cations of Groups IB, IIB,
IVB, VB, VIB, VIIB, and VIIIB. Preferably, the Me represents atoms
such as Zn, Ni, and Cu. These atoms can be inserted into the tetrahedral
framework through a [MeO.sub.2] tetrahedral unit. Incorporation
of the metal component is typically accomplished by adding the metal
component during synthesis of the molecular sieve. However, post-synthesis
metal incorporation can also be used.
[0037] SAPO and ALPO molecular sieves that can be used include
SAPO-5 SAPO-8 SAPO-11 SAPO-16 SAPO-17 SAPO-18 SAPO-20 SAPO-31
SAPO-34 SAPO-35 SAPO-36 SAPO-37 SAPO-40 SAPO-41 SAPO-42 SAPO-44
SAPO-47 SAPO-56 ALPO-5 ALPO-11 ALPO-18 ALPO-31 ALPO-34 ALPO-36
ALPO-37 ALPO-46 the metal containing forms thereof, and mixtures
thereof. Preferred are SAPO-18 SAPO-34 SAPO-35 SAPO-44 SAPO-56
ALPO-18 and ALPO-34 particularly SAPO-18 SAPO-34 ALPO-34 and
ALPO-18 including the metal containing forms thereof, and mixtures
thereof. As used herein, the term mixture is synonymous with combination
and is considered a composition of matter having two or more components
in varying proportions, regardless of their physical state.
[0038] SAPO and ALPO molecular sieves are synthesized by hydrothermal
crystallization methods generally known in the art. See, for example,
U.S. Pat. Nos. 4440871; 4861743; 5096684; and 5126308 the
disclosures of which are fully incorporated herein by reference.
A reaction mixture is formed by mixing together reactive silicon,
aluminum and phosphorus components, along with at least one template.
Generally the mixture is sealed and heated, preferably under autogenous
pressure, to a temperature of at least 100.degree. C., preferably
from 100-250.degree. C., until a crystalline product is formed.
[0039] Formation of the crystalline product can take anywhere from
around 2 hours to as much as 2 weeks. In some cases, stirring or
seeding with crystalline material will facilitate the formation
of the product. Typically, the molecular sieve product is formed
in solution. It can be recovered by standard means, such as by centrifugation
or filtration. The product can also be washed, recovered by the
standard means, and dried. In one method, the molecular sieve is
washed and collected by a filtration process that maintains the
molecular sieve in slurry form. This process includes adding wash
fluid as the molecular sieve is concentrated from the synthesis
solution.
[0040] Additional molecular sieve materials can be included as
a part of the non zeolitic catalyst or they can be used as separate
molecular sieve catalysts in admixture with the non zeolitic molecular
sieve catalyst if desired. Structural types of small pore molecular
sieves that are suitable for use in this invention include AEI,
AFT, APC, ATN, ATT, ATV, AWW, BIK, CAS, CHA, CHI, DAC, DDR, EDI,
ERI, GOO, KFI, LEV, LOV, LTA, MON, PAU, PHI, RHO, ROG, THO, and
substituted forms thereof. Structural types of medium pore molecular
sieves that are suitable for use in this invention include MFI,
MEL, MTW, EUO, MTT, HEU, FER, AFO, AEL, TON, and substituted forms
thereof. These small and medium pore molecular sieves are described
in greater detail in the Atlas of Zeolite Structural Types, W. M.
Meier and D. H. Olsen, Butterworth Heineman, 3rd ed., 1997 the
detailed description of which is explicitly incorporated herein
by reference. Preferred molecular sieves which can be combined with
a silicoaluminophosphate and/or an aluminophosphate catalyst include
ZSM-5 ZSM-34 erionite, and chabazite.
[0041] Binders
[0042] Once the desired type or types of non-zeolitic molecular
sieve is selected based upon the desired activity and selectivity
of the catalyst, other materials are blended with the non-zeolitic
molecular sieve. One of these materials includes one or more binders,
such as a type of hydrated alumina, and/or an inorganic oxide sol
such as aluminum chlorhydrol. The inorganic oxide sol is essentially
a "glue" which binds the catalyst components together
upon thermal treatment. After the formed catalyst particle is formed
and heated, the inorganic oxide sol is converted to an inorganic
oxide matrix component. For example, an alumina sol will convert
to an aluminum oxide matrix following a heat treatment of the formed
catalyst. Aluminum chlorhydrol is a hydroxylated aluminum based
sol containing chloride as the counter ion. Aluminum chlorhydrol
has the general formula of Al.sub.mO.sub.n(OH).sub.oCl.sub.p.xH.sub.2O
wherein m is 1 to 20 n is 1 to 8 o is 5 to 40 p is 2 to 15 and
x is 0 to 30. Although the equilibria of alumina species in the
sol is complex, the predominant species is believed to be [Al.sub.13O.sub.4(OH).sub.24Cl.sub.-
7(H.sub.2O).sub.12]. In addition, other alumina materials may be
added with the aluminum chlorhydrol. Materials that can be used
include, but are not necessarily limited to aluminum oxyhydroxide,
.gamma.-alumina, boehmite, diaspore, and transitional aluminas such
as .alpha.-alumina, .beta.-alumina, .gamma.-alumina, .delta.-alumina,
.epsilon.-alumina, .kappa.-alumina, and .rho.-alumina. Aluminum
trihydroxide, such as gibbsite, bayerite, nordstrandite, doyelite,
and mixtures thereof, also can be used. A sufficient amount of the
binder is added to the slurry mixture so that the amount of the
resultant inorganic oxide matrix in the catalyst, not including
the inorganic oxide framework of the non-zeolitic molecular sieve,
is from about 2% to about 30% by weight, preferably from about 5%
to about 20% by weight, and more preferably from about 7% to about
12% by weight.
[0043] Matrix Materials
[0044] The non zeolitic molecular sieve catalysts will also contain
clay, preferably Kaolin. Matrix materials may also include compositions
such as various forms of rare earth metals, metal oxides, titania,
zirconia, magnesia, thoria, beryllia, quartz, silica or silica or
silica sol, and mixtures thereof. The added matrix materials components
are effective in reducing, inter alia, overall catalyst cost, acting
as a thermal sink to assist in heat shielding the catalyst during
regeneration, densifying the catalyst and increasing catalyst strength.
The use of matrix materials such as naturally occurring clays, e.g.,
bentonite and kaolin, improves the crush strength of the catalyst
under commercial operating conditions. Thus, the addition of clays
improve upon the attrition resistance of the catalyst. The inactive
materials also serve as diluents to control the rate of conversion
in a given process so that more expensive means for controlling
the rate of reaction is eliminated or minimized. Naturally occurring
clays which can be used in the present invention include the montmorillonite
and kaolin families which include the sabbentonites, and the kaolins,
commonly known as Dixie, McNamee, Georgia and Florida clays, or
other in which the main mineral constituent is haloysite, kaolinite,
dickite, nacrite, or anauxite.
[0045] As with most catalysts clay is used in the invention as
an inert densifier, and for the most part the clay has no effect
on catalytic activity or selectivity. Kaolin's ability to form pumpable,
high solid content slurries, low fresh surface area, and ease of
packing because of its platelet structure makes it particularly
suitable for catalyst processing. The preferred average particle
size of the kaolin is 0.1 .mu.m to 0.6 .mu.m with a D90 particle
size of about 1 .mu.m. Because of environmental concerns, the crystalline
silica content of the clay has also become an important parameter.
[0046] Mixing and Spray Drying.
[0047] Rigorous mixing of the catalyst components is necessary
to produce a hard, dense, homogeneous catalyst particle. The primary
consequence of poor mixing are poor attrition and poor catalyst
density. Stratification of the components caused by incomplete mixing
can also effect the activity and selectivity of the catalyst. Generally,
the mixers are of a high shear type because of the thixotropic nature
of the slurries. The resultant slurry may be colloid-milled for
a period sufficient to obtain a desired sub-particle texture, sub-particle
size, and/or sub-particle size distribution.
[0048] The catalyst particle contains a plurality of catalyst sub-particles.
The average diameter of the catalyst particle is from 40 .mu.m to
300 .mu.m, preferably from 50 .mu.m to 200 .mu.m. The catalyst sub-particles
contain non-zeolite molecular sieve, typically SAPO molecular sieve,
an aluminum oxide matrix, and a matrix material, typically clay.
Preparation of the catalyst begins with mixing one or more non-zeolite
molecular sieve, one or more inorganic oxide sols, one or more matrix
materials, and a fluid, typically water, to form a slurry. Other
fluids, e.g., alcohol, can be used along with the water.
[0049] The preferred slurry is prepared by mixing the non-zeolitic
molecular sieve with aluminum chlorhydrol and Kaolin clay, together
or in sequence, in dry form or as slurries. If the solids are added
together as dry solids, a limited and controlled amount of water
is added. The slurry may also contain other materials including
other forms of molecular sieve, other binders, and other matrix
materials. The mesoporosity of the catalyst and the mechanical strength
of the catalyst is dependent on the amount of water contained in
the slurry. In general, it has been found that the weight percent
of solids in the slurry can range from 20% to 70% by weight, preferably
from 40% to 60% by weight. When the weight percent of solids in
the slurry is greater than 70% by weight, the viscosity of the slurry
is too high to spray dry, and when the weight percent of solids
in the slurry is less than 20% by weight the attrition resistance
of the catalyst is poor. It is also desirable that the density of
the slurry be greater than 1.1 g/cc, and preferably greater than
1.18 g/cc to form the catalysts of this invention.
[0050] The solid content of the slurry will contain about 10% to
about 50%, preferably from about 20% to about 45% by weight, non-zeolitic
molecular sieve, about 5% to about 20%, preferably from about 8%
to about 15% by weight, binder, and about 30% to about 80%, preferably
about 40% to about 60% by weight, matrix material. The slurry is
mixed or milled to achieve a sufficiently uniform slurry of catalyst
sub-particles. The slurry is then fed to a forming unit to produce
catalyst particles. The forming unit is maintained at a temperature
sufficient to remove most of the water from the formed catalyst
particles. Preferably, the forming unit is a spray dryer. The formed
catalyst particles typically take the form of microspheres. Typically,
the slurry is fed to a spray drier at an average inlet temperature
ranging from 200.degree. C. to 450.degree. C., and an outlet temperature
ranging from 100.degree. C. to about 225.degree. C.
[0051] During spray drying, the slurry is passed through a nozzle
which distributes the slurry into small droplets, resembling an
aerosol. A single nozzle unit or multiple nozzle unit may be used
to disperse an inlet stream of slurry (single-fluid nozzle) into
the atomization chamber. Alternatively, a multiple nozzles may be
used to co-feed the slurry into the atomization chamber. Alternatively,
the slurry is directed to the perimeter of a spinning wheel which
also distributes the slurry into small droplets. The size of the
distributed slurry droplets is controlled by many factors including
flow rate, pressure, and temperature of the slurry, the shape and
dimension of the nozzle(s), or the spinning rate of the wheel. The
droplets are then dried in a co-current or counter-current flow
of air passing through the spray drier. Dry catalyst particles in
the form of a powder are recovered from each droplet.
[0052] Catalyst particle size to some extent is controlled by the
solids content of the slurry and its viscosity. All else being equal,
the catalyst particle size is directly proportional to the solids
content of the slurry. However, control of the catalyst particle
size and spherical characteristics also depend on the size and shape
of the drying chamber as well as the atomization procedure used.
A Boltzmann distribution of catalyst particle size is invariably
obtained around a mean, which is usually set at approximately 70
.mu.m average catalyst particle size. The average catalyst particle
size is controlled by a variation in the slurry feed properties
to the dryer and by the conditions of atomization. It is preferred
that the formulated catalyst composition have a catalyst size from
40 .mu.m to 300 .mu.m, more preferably 50 .mu.m to 200 .mu.m, most
preferably 50 .mu.m to 150 .mu.m.
[0053] Calcination.
[0054] To harden and/or activate the formed catalysts a heat treatment,
i.e., calcination, at an elevated temperature is usually necessary.
Ordinarily, catalysts with alumina or silica binders are heated
in a calcination environment at a temperature between 500.degree.
C. and 800.degree. C. The conventional calcination environment is
air, which may include small amounts of water vapor.
[0055] The invention provides methods of heat treating a formed
non-zeolitic molecular sieve catalyst prepared with an inorganic
oxide sol that contains halogen. The methods of the invention minimize
the production of halogen-containing acids or at least confines
much of the produced halogen-containing acids to a single heating
unit. The schematic diagram in FIG. 1 depicts one embodiment of
the invention, in which a chlorine-containing SAPO catalyst 12 is
used by way of example. Catalyst 12 is supplied from a forming unit,
preferably a spray dryer, and is directed to a heat treatment unit
14. The catalyst is heated at a temperature from about 400.degree.
C. to about 1000.degree. C., preferably from about 500.degree. C.
to about 800.degree. C., most preferably from about 550.degree.
C. to about 700.degree. C. in a low moisture calcination environment
containing less than 5% by volume water, preferably less than 1%
by volume water. The low moisture calcination environment can be
provided by using a dry gas 18 e.g., air that has been adequately
dried, nitrogen, helium, flue gas, or any combination thereof. In
the preferred embodiment, the catalyst is heated in nitrogen off
gas at a temperature from about 600.degree. C. to about 700.degree.
C. Nitrogen off gas is the gas produced from the boil-off gas of
a liquid nitrogen source. Heating is carried out for a period of
time sufficient to remove chlorides, typically for a period of from
0.5 to 10 hours, preferably of from 1 to 5 hours, most preferably
from 2 to 4 hours.
[0056] As the catalyst is heated in the low moisture calcination
environment most of the chlorine is removed as chlorine gas or as
a non-hydrated form of hydrochloric acid (HCl.sub.(g)). HCl.sub.(g)
is not as corrosive as HCl.sub.(aq). Approximately 60% to 98% by
weight, preferably 85% to 98% by weight, of the chlorine in the
formed catalyst 12 is removed during the heat treatment in the low
moisture calcination environment. Following this heat treatment
the catalyst 16 contains less than about 6000 ppmw chlorine, preferably
less than about 3000 ppmw chlorine.
[0057] Following the low moisture heat treatment, the catalyst
16 is directed to heating unit 20. The catalyst in heating unit
20 is heated in a second calcination environment 24. This second
calcination environment 24 contains from about 5% to about 10% by
volume water. The remaining volume of gas in the calcination environment
24 may include air, nitrogen, helium, flue gas, or any combination
thereof. The second heat treatment of the catalyst 16 will take
place at a temperature from about 400.degree. C. to about 1000.degree.
C., preferably from about 500.degree. C. to about 800.degree. C.,
more preferably from about 600.degree. C. to about 700.degree. C.
The period during which the catalysts is heated in unit 20 ranges
from 0.1 to 5 hours, preferably from 0.25 to 4 hours. This second
heat treatment results in a loss of about 2% to about 95% of the
chlorine remaining in catalyst 16. Catalyst 22 will contain less
than about 600 ppmw chlorine, preferably less than about 200 ppmw
chlorine, more preferably less than about 80 ppmw chlorine.
[0058] In another embodiment, the second calcination environment
contains at least 10% by volume water. The remaining volume of gas
in the second calcination environment may include air, nitrogen,
helium, flue gas, or any combination thereof. Preferably, the second
calcination environment contains air. A catalyst that is contacted
with a calcination environment containing at least 10% by volume
water is said to be steam-treated. Steam treatment results in a
loss of about 50% to 99%, preferably in a loss of about 90% to about
99% of the chlorine remaining in the catalyst following the low
moisture heat treatment. The steam-treated catalyst will contain
about 10 ppmw to about 400 ppmw chlorine, preferably about 10 ppmw
to about 200 ppmw chlorine, more preferably about 10 ppmw to about
80 ppmw chlorine.
[0059] Steam treatment of the catalyst will take place at a temperature
from about 400.degree. C. to about 1000.degree. C., preferably from
about 500.degree. C. to about 800.degree. C., more preferably from
about 600.degree. C. to about 700.degree. C. The period during which
the catalysts is heated in unit 20 ranges from 0.1 to 5 hours, preferably
from 0.25 to 4 hours. Although temperatures of about 400.degree.
C. are sufficient to adequately remove most of the chlorine from
the catalyst, the rate at which the additional chlorine is removed
will be lower than if a higher temperature, e.g., 600.degree. C.,
is used during steam treatment. On the other hand, if the temperature
of the steam treatment is too high, e.g., greater than 1000.degree.
C., degradation of the catalyst may occur. The temperature at which
degradation of the catalyst will occur will vary for different catalyst
formulations and various non-zeolitic molecular sieve.
[0060] The low moisture heat treatment followed by steam treatment
can remove about 70% to about 99.99% by weight, preferably about
95% to about 99.99% by weight, more preferably about 98% to about
99.99% by weight, of the chlorine in the formed catalyst. The steam
treatment will produce HCl.sub.(aq), but the amount of HCl.sub.(aq)
produced is significantly reduced because most of the chlorine is
removed during the initial heat treatment in the low moisture calcination
environment. As a result, the production of the HCl.sub.(aq) is
minimized. Also, if separate heating units are used the production
of HCl.sub.(aq) will be confined to the steam treatment unit, which
can be designed to accommodate the HCl.sub.(aq) produced.
[0061] If air is not used in the steam treatment, the catalyst
may be calcined in a calcination environment containing at least
3% by volume, preferably at least 10% by volume, oxygen to remove
template material that may have remained in the pores of the sieve.
A catalyst that has been calcined in an environment that contains
at least 3% by volume oxygen is said to be oxygen treated. The oxygen
environment may be provided by air or a mixture of air and nitrogen.
The calcination temperature of this oxygen environment may be the
same or different than the temperature of the steam treatment.
[0062] It is to be understood that although FIG. 1 depicts more
than one heating unit for each type of heat treatment, a single
heating unit may be used. In this case, the heating environment
is changed by alternating the type of gas flow, e.g., from nitrogen
off gas to steam, or from air to steam. Alternatively, different
heating zones in a singular heating unit may be used according to
the invention. Each heating zone will contain a different calcination
environment with a transition zone disposed between the heating
zones. The temperature and gas flow for each heating zone or heating
unit can be programmed to minimize the time required to remove the
desired amount of chlorine, while minimizing the amount of HCl.sub.(aq)
produced. The heat and steam treatments may be done in any of a
number of heating units well known to those skilled in the art including
moving bed reactors, rotary kilns, rotary calciners, fluidized beds
and packed-bed batch reactors.
[0063] In another embodiment the steam treatment is used to remove
most of the halogen from the formed catalyst. Prior heating in a
low moisture environment is not necessary. The formed catalyst is
steam-treated at a temperature from about 400.degree. C. to about
1000.degree. C., preferably from about 500.degree. C. to about 800.degree.
C., more preferably from about 600.degree. C. to about 700.degree.
C. The period during which the catalysts is heated in unit 20 ranges
from 0.1 to 5 hours, preferably from 0.25 to 4 hours. The steam
treatment may remove from about 70% to about 99.99% by weight, preferably
from about 95% to about 99.99% by weight, more preferably from about
98% to about 99.99% by weight, of the chlorine in the formed catalyst.
[0064] Following the steam treatment, the catalyst may be oxygen
treated to remove template material that may have remained in the
pores of the sieve. The calcination temperature of this oxygen environment
may be the same or different than the temperature of the steam treatment.
[0065] In another embodiment, steam treatment of the catalyst may
take place after an oxygen heat treatment. The catalyst is heated
in an oxygen environment at a temperature from 400.degree. C. to
1000.degree. C., preferably from about 500.degree. C. to about 800.degree.
C., more preferably from about 600.degree. C. to about 700.degree.
C. The period during which the catalysts is heated in unit 20 ranges
from 0.1 to 5 hours, preferably from 0.25 to 4 hours. Approximately
50% to 95% by weight, preferably 75% to 95% by weight, of the chlorine
in the formed catalyst is removed during the oxygen heat treatment.
The oxygen treated catalyst is then contacted with steam to remove
additional amounts of chlorine from the catalyst. This steam contacted
catalyst will contain about 10 ppmw to about 600 ppmw chlorine,
preferably 10 ppmw to about 200 ppmw chlorine, more preferably 10
ppmw to about 80 ppmw chlorine.
[0066] The oxygen heat treatment and the steam treatment of the
catalyst may take place in separate heating units or in the same
heating unit though in different regions of that unit. For example,
the oxygen environment may be introduced near the entrance to the
heating unit and steam added near the middle of the heating unit.
In this way partial calcination of the catalyst occurs prior to
the catalyst contacting the steam.
[0067] The catalysts of the present invention are useful in a variety
of processes including: cracking, hydrocracking, isomerization,
polymerisation, reforming, hydrogenation, dehydrogenation, dewaxing,
hydrodewaxing, absorption, alkylation, transalkylation, dealkylation,
hydrodecylization, disproportionation, oligomerization, dehydrocyclization
and combinations thereof. Due to the low level, or even absence
of halogen in the non zeolitic catalyst, less corrosive acid is
created during the catalytic process. Reactors may thus be used
for longer periods of time before repair or replacement needs to
take place.
[0068] The preferred processes of the present invention include
a process directed to the conversion of a feedstock comprising one
or more oxygenates to one or more olefin(s) and a process directed
to the conversion of ammonia and one or more oxygenates to alkyl
amines and in particular methylamines.
[0069] In a preferred embodiment of the process of the invention,
the feedstock contains one or more oxygenates, more specifically,
one or more organic compound(s) containing at least one oxygen atom.
In the most preferred embodiment of the process of the invention,
the oxygenate in the feedstock is one or more alcohol(s), preferably
aliphatic alcohol(s) where the aliphatic moiety of the alcohol(s)
has from 1 to 20 carbon atoms, preferably from 1 to 10 carbon atoms,
and most preferably from 1 to 4 carbon atoms. The alcohols useful
as feedstock in the process of the invention include lower straight
and branched chain aliphatic alcohols and their unsaturated counterparts.
[0070] Non-limiting examples of oxygenates include methanol, ethanol,
n-propanol, isopropanol, methyl ethyl ether, dimethyl ether, diethyl
ether, di-isopropyl ether, formaldehyde, dimethyl carbonate, dimethyl
ketone, acetic acid, and mixtures thereof.
[0071] In the most preferred embodiment, the feedstock is selected
from one or more of methanol, ethanol, dimethyl ether, diethyl ether
or a combination thereof, more preferably methanol and dimethyl
ether, and most preferably methanol.
[0072] In the most preferred embodiment, the feedstock, preferably
of one or more oxygenates, is converted in the presence of a catalyst
into olefin(s) having 2 to 6 carbons atoms, preferably 2 to 4 carbon
atoms. Most preferably, the olefin(s), alone or combination, are
converted from a feedstock containing an oxygenate, preferably an
alcohol, most preferably methanol, to the preferred olefin(s) ethylene
and/or propylene.
[0073] The most preferred process is generally referred to as gas-to-olefins
(GTO) or alternatively, methanol-to-olefins (MTO). In a MTO process,
typically an oxygenated feedstock, most preferably a methanol containing
feedstock, is converted in the presence of a catalyst into one or
more olefin(s), preferably and predominantly, ethylene and/or propylene,
often referred to as light olefin(s).
[0074] In one embodiment of the process for conversion of a feedstock,
preferably a feedstock containing one or more oxygenates, the amount
of olefin(s) produced based on the total weight of hydrocarbon produced
is greater than 50 weight percent, preferably greater than 60 weight
percent, more preferably greater than 70 weight percent.
[0075] The feedstock, in one embodiment, contains one or more diluent(s),
typically used to reduce the concentration of the feedstock, and
are generally non-reactive to the feedstock or catalyst. Non-limiting
examples of diluents include helium, argon, nitrogen, carbon monoxide,
carbon dioxide, water, essentially non-reactive paraffins (especially
alkanes such as methane, ethane, and propane), essentially non-reactive
aromatic compounds, and mixtures thereof. The most preferred diluents
are water and nitrogen, with water being particularly preferred.
[0076] The diluent, water, is used either in a liquid or a vapour
form, or a combination thereof. The diluent is either added directly
to a feedstock entering into a reactor or added directly into a
reactor, or added with a molecular sieve catalyst composition. In
one embodiment, the amount of diluent in the feedstock is in the
range of from about 1 to about 99 mole percent based on the total
number of moles of the feedstock and diluent, preferably from about
1 to 80 mole percent, more preferably from about 5 to about 50
most preferably from about 5 to about 25. In one embodiment, other
hydrocarbons are added to a feedstock either directly or indirectly,
and include olefin(s), paraffin(s), aromatic(s) (see for example
U.S. Pat. No. 4677242 addition of aromatics) or mixtures thereof,
preferably propylene, butylene, pentylene, and other hydrocarbons
having 4 or more carbon atoms, or mixtures thereof.
[0077] The process for converting a feedstock, especially a feedstock
containing one or more oxygenates, in the presence of a catalyst
of the invention, is carried out in a reaction process in a reactor,
where the process is a fixed bed process, a fluidised bed process
(includes a turbulent bed process), preferably a continuous fluidised
bed process, and most preferably a continuous high velocity fluidised
bed process.
[0078] The reaction processes can take place in a variety of catalytic
reactors such as hybrid reactors that have a dense bed or fixed
bed reaction zones and/or fast fluidised bed reaction zones coupled
together, circulating fluidised bed reactors, riser reactors, and
the like. Suitable conventional reactor types are described in for
example U.S. Pat. No. 4076796 U.S. Pat. No. 6287522 (dual riser),
and Fluidization Engineering, D. Kunii and O. Levenspiel, Robert
E. Krieger Publishing Company, New York, N.Y. 1977 which are all
herein fully incorporated by reference.
[0079] The preferred reactor type are riser reactors generally
described in Riser Reactor, Fluidization and Fluid-Particle Systems,
pages 48 to 59 F. A. Zenz and D. F. Othmo, Reinhold Publishing
Corporation, New York, 1960 and U.S. Pat. No. 6166282 (fast-fluidised
bed reactor), and U.S. patent application Ser. No. 09/564613 filed
May 4 2000 (multiple riser reactor), which are all herein fully
incorporated by reference.
[0080] In the preferred embodiment, a fluidised bed process or
high velocity fluidised bed process includes a reactor system, a
regeneration system and a recovery system.
[0081] The reactor system preferably is a fluid bed reactor system
having a first reaction zone within one or more riser reactor(s)
and a second reaction zone within at least one disengaging vessel,
preferably comprising one or more cyclones. In one embodiment, the
one or more riser reactor(s) and disengaging vessel is contained
within a single reactor vessel. Fresh feedstock, preferably containing
one or more oxygenates, optionally with one or more diluent(s),
is fed to the one or more riser reactor(s) in which a catalyst or
coked version thereof is introduced. In one embodiment, the catalyst
or coked version thereof is contacted with a liquid or gas, or combination
thereof, prior to being introduced to the riser reactor(s), preferably
the liquid is water or methanol, and the gas is an inert gas such
as nitrogen.
[0082] In an embodiment, the amount of fresh feedstock fed separately
or jointly with a vapour feedstock, to a reactor system is in the
range of from 0.1 weight percent to about 85 weight percent, preferably
from about 1 weight percent to about 75 weight percent, more preferably
from about 5 weight percent to about 65 weight percent based on
the total weight of the feedstock including any diluent contained
therein. The liquid and vapour feedstocks are preferably the same
composition, or contain varying proportions of the same or different
feedstock with the same or different diluent.
[0083] The feedstock entering the reactor system is preferably
converted, partially or fully, in the first reactor zone into a
gaseous effluent that enters the disengaging vessel along with a
coked catalyst. In the preferred embodiment, cyclone(s) within the
disengaging vessel are designed to separate the catalyst, preferably
a coked catalyst, from the gaseous effluent containing one or more
olefin(s) within the disengaging zone. Cyclones are preferred, however,
gravity effects within the disengaging vessel will also separate
the catalyst compositions from the gaseous effluent. Other methods
for separating the catalyst compositions from the gaseous effluent
include the use of plates, caps, elbows, and the like.
[0084] In one embodiment of the disengaging system, the disengaging
system includes a disengaging vessel; typically a lower portion
of the disengaging vessel is a stripping zone. In the stripping
zone the coked catalyst is contacted with a gas, preferably one
or a combination of steam, methane, carbon dioxide, carbon monoxide,
hydrogen, or an inert gas such as argon, preferably steam, to recover
adsorbed hydrocarbons from the coked catalyst that is then introduced
to the regeneration system. In another embodiment, the stripping
zone is in a separate vessel from the disengaging vessel and the
gas is passed at a gas hourly superficial velocity (GHSV) of from
1 hr.sup.-1 to about 20000 hr.sup.-1 based on the volume of gas
to volume of coked catalyst, preferably at an elevated temperature
from 250.degree. C. to about 750.degree. C., preferably from about
350.degree. C. to 650.degree. C., over the coked catalyst.
[0085] The conversion temperature employed in the conversion process,
specifically within the reactor system, is in the range of from
about 200.degree. C. to about 1000.degree. C., preferably from about
250.degree. C. to about 800.degree. C., more preferably from about
250.degree. C. to about 750 .degree. C., yet more preferably from
about 300.degree. C. to about 650.degree. C., yet even more preferably
from about 350.degree. C. to about 600.degree. C. most preferably
from about 350.degree. C. to about 550.degree. C.
[0086] The conversion pressure employed in the conversion process,
specifically within the reactor system, varies over a wide range
including autogenous pressure. The conversion pressure is based
on the partial pressure of the feedstock exclusive of any diluent
therein. Typically the conversion pressure employed in the process
is in the range of from about 0.1 kPaa to about 5 MPaa, preferably
from about 5 kPaa to about 1 MPaa, and most preferably from about
20 kPaa to about 500 kpaa.
[0087] The weight hourly space velocity (WHSV), particularly in
a process for converting a feedstock containing one or more oxygenates
in the presence of a catalyst within a reaction zone, is defined
as the total weight of the feedstock excluding any diluents to the
reaction zone per hour per weight of molecular sieve in the catalyst
in the reaction zone. The WHSV is maintained at a level sufficient
to keep the catalyst composition in a fluidised state within a reactor.
[0088] Typically, the WHSV ranges from about 1 hr.sup.-1 to about
5000 hr.sup.-1 preferably from about 2 hr.sup.-1 to about 3000
hr.sup.-1 more preferably from about 5 hr.sup.-1 to about 1500
hr.sup.-1 and most preferably from about 10 hr.sup.-1 to about
1000 hr.sup.-1. In one preferred embodiment, the WHSV is greater
than 20 hr.sup.-1; preferably the WHSV for conversion of a feedstock
containing methanol and dimethyl ether is in the range of from about
20 hr.sup.-1 to about 300 hr.sup.-1.
[0089] The superficial gas velocity (SGV) of the feedstock including
diluent and reaction products within the reactor system is preferably
sufficient to fluidise the catalyst within a reaction zone in the
reactor. The SGV in the process, particularly within the reactor
system, more particularly within the riser reactor(s), is at least
0.1 meter per second (m/sec), preferably greater than 0.5 m/sec,
more preferably greater than 1 m/sec, even more preferably greater
than 2 m/sec, yet even more preferably greater than 3 m/sec, and
most preferably greater than 4 m/sec. See for example U.S. patent
application Ser. No. 09/708753 filed Nov. 8 2000 which is herein
incorporated by reference.
[0090] In one preferred embodiment of the process for converting
an oxygenate to olefin(s) using a silicoaluminophosphate catalyst,
the process is operated at a WHSV of at least 20 hr.sup.-1 and a
Temperature Corrected Normalized Methane Selectivity (TCNMS) of
less than 0.016 preferably less than or equal to 0.01. See for
example U.S. Pat. No. 5952538 which is herein fully incorporated
by reference.
[0091] In another embodiment of the processes for converting an
oxygenate such as methanol to one or more olefin(s) using a catalyst,
the WHSV is from 0.01 hr.sup.-1 to about 100 hr.sup.-1 at a temperature
of from about 350.degree. C. to 550.degree. C., and silica to Me.sub.2O.sub.3
(Me is a Group IIIA or VIII element from the Periodic Table of Elements)
molar ratio of from 300 to 2500. See for example EP-0 642 485 B1
which is herein fully incorporated by reference.
[0092] Other processes for converting an oxygenate such as methanol
to one or more olefin(s) using a catalyst are described in PCT WO
01/23500 published Apr. 5 2001 (propane reduction at an average
catalyst feedstock exposure of at least 1.0), which is herein incorporated
by reference.
[0093] The coked catalyst is withdrawn from the disengaging vessel,
preferably by one or more cyclones(s), and introduced to the regeneration
system. The regeneration system comprises a regenerator where the
coked catalyst composition is contacted with a regeneration medium,
preferably a gas containing oxygen, under general regeneration conditions
of temperature, pressure and residence time.
[0094] Non-limiting examples of the regeneration medium include
one or more of oxygen, O.sub.3 SO.sub.3 N.sub.2O, NO, NO.sub.2
N.sub.2O.sub.5 air, air diluted with nitrogen or carbon dioxide,
oxygen and water (U.S. Pat. No. 6245703), carbon monoxide and/or
hydrogen. The regeneration conditions are those capable of burning
coke from the coked catalyst composition, preferably to a level
less than 0.5 weight percent based on the total weight of the coked
catalyst entering the regeneration system. The coked catalyst withdrawn
from the regenerator forms a regenerated catalyst.
[0095] The regeneration temperature is in the range of from about
200.degree. C. to about 1500.degree. C., preferably from about 300.degree.
C. to about 1000.degree. C., more preferably from about 450.degree.
C. to about 750.degree. C., and most preferably from about 550.degree.
C. to 700.degree. C. The regeneration pressure is in the range of
from about 15 psia (103 kPaa) to about 500 psia (3448 kPaa), preferably
from about 20 psia (138 kPaa) to about 250 psia (1724 kpaa), more
preferably from about 25 psia (172 kPaa) to about 150 psia (1034
kPaa), and most preferably from about 30 psia (207 kPaa) to about
60 psia (414 kPaa).
[0096] The preferred residence time of the catalyst in the regenerator
is in the range of from about one minute to several hours, most
preferably about one minute to 100 minutes, and the preferred volume
of oxygen in the gas is in the range of from about 0.01 mole percent
to about 5 mole percent based on the total volume of the gas.
[0097] In one embodiment, regeneration promoters, typically metal
containing compounds such as platinum, palladium and the like, are
added to the regenerator directly, or indirectly, for example with
the coked catalyst composition. Also, in another embodiment, a fresh
catalyst is added to the regenerator containing a regeneration medium
of oxygen and water as described in U.S. Pat. No. 6245703 which
is herein fully incorporated by reference.
[0098] In an embodiment, a portion of the coked catalyst from the
regenerator is returned directly to the one or more riser reactor(s),
or indirectly, by pre-contacting with the feedstock, or contacting
with fresh catalyst, or contacting with a regenerated catalyst or
a cooled regenerated catalyst described below.
[0099] The burning of coke is an exothermic reaction, and in an
embodiment, the temperature within the regeneration system is controlled
by various techniques in the art including feeding a cooled gas
to the regenerator vessel, operated either in a batch, continuous,
or semi-continuous mode, or a combination thereof. A preferred technique
involves withdrawing the regenerated catalyst from the regeneration
system and passing the regenerated catalyst through a catalyst cooler
that forms a cooled regenerated catalyst. The catalyst cooler, in
an embodiment, is a heat exchanger that is located either internal
or external to the regeneration system.
[0100] In one embodiment, the cooler regenerated catalyst is returned
to the regenerator in a continuous cycle, alternatively, (see U.S.
patent application Ser. No. 09/587766 filed Jun. 6 2000) a portion
of the cooled regenerated catalyst is returned to the regenerator
vessel in a continuous cycle, and another portion of the cooled
molecular sieve regenerated catalyst is returned to the riser reactor(s),
directly or indirectly, or a portion of the regenerated catalyst
or cooled regenerated catalyst is contacted with by-products within
the gaseous effluent (PCT WO 00/49106 published Aug. 24 2000),
which are all herein fully incorporated by reference. In another
embodiment, a regenerated catalyst contacted with an alcohol, preferably
ethanol, 1-propnaol, 1-butanol or mixture thereof, is introduced
to the reactor system, as described in U.S. patent application Ser.
No. 09/785122 filed Feb. 16 2001 which is herein fully incorporated
by reference.
[0101] Other methods for operating a regeneration system are in
disclosed U.S. Pat. No. 6290916 (controlling moisture), which
is herein fully incorporated by reference.
[0102] The regenerated catalyst withdrawn from the regeneration
system, preferably from the catalyst cooler, is combined with a
fresh catalyst and/or re-circulated catalyst and/or feedstock and/or
fresh gas or liquids, and returned to the riser reactor(s). In another
embodiment, the regenerated catalyst withdrawn from the regeneration
system is returned to the riser reactor(s) directly, preferably
after passing through a catalyst cooler. In one embodiment, a carrier,
such as an inert gas, feedstock vapour, steam or the like, semi-continuously
or continuously, facilitates the introduction of the regenerated
catalyst to the reactor system, preferably to the one or more riser
reactor(s).
[0103] By controlling the flow of the regenerated catalyst or cooled
regenerated catalyst from the regeneration system to the reactor
system, the optimum level of coke on the catalyst entering the reactor
is maintained. There are many techniques for controlling the flow
of a catalyst described in Michael Louge, Experimental Techniques,
Circulating Fluidised Beds, Grace, Avidan and Knowlton, eds. Blackie,
1997 (336-337), which is herein incorporated by reference.
[0104] Coke levels on the catalyst are measured by withdrawing
from the conversion process the catalyst at a point in the process
and determining its carbon content. Typical levels of coke on the
catalyst, after regeneration is in the range of from 0.01 weight
percent to about 15 weight percent, preferably from about 0.1 weight
percent to about 10 weight percent, more preferably from about 0.2
weight percent to about 5 weight percent, and most preferably from
about 0.3 weight percent to about 2 weight percent based on the
total weight of the molecular sieve and not the total weight of
the catalyst.
[0105] In one preferred embodiment, the mixture of fresh catalyst
and regenerated catalyst and/or cooled regenerated catalyst contains
in the range of from about 1 to 50 weight percent, preferably from
about 2 to 30 weight percent, more preferably from about 2 to about
20 weight percent, and most preferably from about 2 to about 10
coke or carbonaceous deposit based on the total weight of the mixture
of catalysts. See for example U.S. Pat. No. 6023005 which is
herein fully incorporated by reference.
[0106] The gaseous effluent is withdrawn from the disengaging system
and is passed through a recovery system. There are many well-known
recovery systems, techniques and sequences that are useful in separating
olefin(s) and purifying olefin(s) from the gaseous effluent. Recovery
systems generally comprise one or more or a combination of a various
separation, fractionation and/or distillation towers, columns, splitters,
or trains, reaction systems such as ethylbenzene manufacture (U.S.
Pat. No. 5476978) and other derivative processes such as aldehydes,
ketones and ester manufacture (U.S. Pat. No. 5675041), and other
associated equipment for example various condensers, heat exchangers,
refrigeration systems or chill trains, compressors, knock-out drums
or pots, pumps, and the like.
[0107] The metalloaluminophosphate molecular sieve materials and
catalyst compositions of the present invention may be used in the
manufacture of alkylamines, using ammonia. Examples of suitable
processes are as described in published European patent application
EP 0 993 867 A1 and in U.S. Pat. No. 6153798 to Hidaka et. al,
which are herein fully incorporated by reference.
[0108] This invention will be better understood with reference
to the following examples, which are intended to illustrate specific
embodiments within the overall scope of the invention as claimed.
EXAMPLE 1
[0109] SAPO-34 molecular sieve, 50% by weight, aluminum chlorhydrol,
10% by weight, and UF grade kaolin clay, 40% by weight, was mixed
with sufficient water to produce a slurry with approximately 40%
by weight solids. The slurry was fed into a spray drier to form
spray dried catalyst. The spray dried catalyst was analyzed by XRF
(X-ray Fluorescence) spectroscopy. The amount of chlorine in the
spray dried catalyst was 33800 ppmw. The GAL Index of the un-calcined
catalyst was greater than 50.
EXAMPLES 2-4
[0110] Spray dried catalyst of Example 1 was heated in a nitrogen
stream at temperatures of 600.degree. C., 650.degree. C. and 700.degree.
C. for one hour. The heat treated catalyst was then analyzed by
XRF to determine the amount of residual chlorine remaining in the
catalyst. Table 1 lists the residual chlorine content of each catalyst.
EXAMPLES 5-7
[0111] Spray dried catalyst of Example 1 was heated in a nitrogen
stream at temperatures of 600.degree. C., 650.degree. C. and 700.degree.
C. for nine hours. The heat treated catalysts were then analyzed
by XRF to determine the amount of residual chlorine remaining in
each catalyst. Table 1 lists the residual chlorine content of each
catalyst.
EXAMPLES 8
[0112] Spray dried catalyst of Example 1 was heated in a nitrogen
stream at temperatures of 650.degree. C. for five hours followed
by heating in air at 650.degree. C. for two hours. The heat treated
catalyst was then analyzed by XRF to determine the amount of residual
chlorine remaining in the catalyst. The chlorine content of the
catalyst was 390 ppm by weight.
EXAMPLES 9-11
[0113] The spray dried catalyst of Example 1 was heated in a nitrogen
stream at temperatures of 600.degree. C., 650.degree. C. and 700.degree.
C. for one hour followed by heating in air at 600.degree. C., 650.degree.
C. and 700.degree. C. for one hour, respectively. The heat treated
catalysts then analyzed by XRF to determine the amount of residual
chlorine remaining catalyst. Table 1 lists the residual chlorine
content of each catalyst.
1TABLE 1 Example Temperature, Sweep gas Sweep gas Chlorine, No.
.degree. C. time, hrs time, hrs ppmw 1 n/a n/a n/a 33800 2 600
N2/1 N/A 620 3 650 N2/1 N/A 520 4 700 N2/1 N/A 480 5 600 N2/9 N/A
440 6 650 N2/9 N/A 430 7 700 N2/9 N/A 350 8 650 N2/5 air/2 390 9
600 N2/1 air/1 510 10 650 N2/1 air/1 470 11 700 N2/1 air/1 430
EXAMPLE 12
[0114] The spray dried catalyst was heated at 600.degree. C. in
air for 120 minutes in an open container placed in an electrically
heated muffle furnace. The calcined catalyst contained 1090 ppm
chlorine (see Table 2).
EXAMPLE 13
[0115] The spray dried catalyst was heated at 650.degree. C. in
air for 120 minutes in an open container placed in an electrically
heated muffle furnace. The calcined catalyst contained 730 ppm chlorine,
and the GAL Index was 1.85 (see Table 2).
EXAMPLE 14
[0116] The spray dried catalyst was heated at 700.degree. C. in
air for 120 minutes in an open container placed in an electrically
heated muffle furnace. The calcined catalyst contained 350 ppm chlorine
(see Table 2).
EXAMPLE 15
[0117] The spray dried catalyst was heated at 600.degree. C. in
air for 120 minutes in an open container placed in an electrically
heated muffle furnace. The calcined catalyst, 12 g, was placed in
a 3/4" OD stainless steel, packed bed tubular reactor that
was electrically heated. About 1 g/min of steam was fed to the reactor
maintained at a temperature of 600.degree. C. The catalyst was heated
in the presence of steam for 120 minutes. The chlorine content of
the treated catalyst was 250 ppm (see Table 2).
EXAMPLE 16
[0118] The same procedure as in Example 15 was used except that
the temperature was maintained at 650.degree. C. for both the heating
in air and heating in steam. The chlorine content of the treated
catalyst was 140 ppm, and the GAL Index was 1.48 (see Table 2).
EXAMPLE 17
[0119] The same procedure as in Example 15 was used except that
the temperature was maintained at 700.degree. C. for both the heating
in air and heating in steam. The chlorine content of the treated
catalyst was 30 ppm (see Table 2).
EXAMPLE 18
[0120] The same procedure as in Examples 12 was used except that
the catalyst was heated in air for 240 minutes. The chlorine content
of the treated catalyst was 830 ppm (see Table 2).
EXAMPLE 19
[0121] The same procedure as in Examples 13 was used except that
the catalyst was heated in air for 240 minutes. The chlorine content
of the treated catalyst was 590 ppm (see Table 2).
EXAMPLE 20
[0122] The same procedure as in Examples 14 was used except that
the catalyst was heated in air for 240 minutes. The chlorine content
of the treated catalyst was 290 ppm (see Table 2).
EXAMPLES 21-26
[0123] The same procedure as in Examples 15 were used except the
times and temperatures of heating in air and the times and temperatures
of heating in steam as indicated in Table 2.
EXAMPLE 27
[0124] The spray dried catalyst was heated at 600.degree. C. in
air for 120 minutes in an open container placed in an electrically
heated muffle furnace. The calcined catalyst then placed in a 3/4"
OD stainless steel, packed bed tubular reactor that was electrically
heated. About 1 g/min of steam at about 1 atm was fed to the reactor
maintained at a temperature of 600.degree. C. The catalyst was heated
in the presence of steam for 240 minutes. The chlorine content of
the treated catalyst was 150 ppmw, and the GAL Index was 2.24 (see
Table 2).
EXAMPLE 28
[0125] The same procedure as in Example 27 was used except the
temperatures of heating in air and the steam treatment was 650.degree.
C. The chlorine content of the treated catalyst was 40 ppmw, and
the GAL Index was 1.62 (see Table 2).
[0126] As summarized in Table 2 heating in air for 120 minutes
at 600.degree. C., 650.degree. C. and 700.degree. C. without a subsequent
steam treatment reduces the chlorine content to 1090 730 or 350
ppm respectively. Increasing the heating time to 240 minutes at
600.degree. C., 650.degree. C. and 700.degree. C., results in the
further reduction in chlorine content to 830 590 or 290 ppm, respectively.
As indicated only small amounts of additional chlorine is removed
by a doubling of the heating time. For example, heating at 650.degree.
C. during the first 120 minutes reduces the chlorine content in
the catalyst by about 98%, i.e., from 33800 ppm to 730 ppm. Heating
for a second 120 minutes reduces the remaining chlorine content
by an additional 19%, i.e., from 730 ppm to 590 ppm.
[0127] Heating in air for 120 minutes followed by heating in the
presence of steam for 120 minutes at temperatures of 600.degree.
C., 650.degree. C., and 700.degree. C. reduces the chlorine content
to 250 140 and 30 ppm, respectively. Increasing the time the catalyst
is heated in air and steam to 240 minutes, respectively, has little
affect on further reducing the chlorine content as shown by a comparison
of Examples 15-17 with Examples 21-23 respectively.
[0128] Examples 24-26 indicate that increasing the time the catalyst
is steam treated at a given temperature (650.degree. C. in these
examples) following the heat treatment in air for 120 minutes results
in a yet greater reduction in chlorine content. The most dramatic
reduction in chlorine content is made during the first 15 minutes
of contacting the heat treated catalyst with steam. For example,
comparison of Example 13 with Example 24 suggests that the chlorine
content is reduced from 730 ppm to 230 ppm after an additional 15
minute steam treatment at 650.degree. C. This amounts to an additional
chlorine reduction of about 68%. Also, as indicated in Table 2 greater
than 99% of the chlorine may be removed from the catalyst following
the steam treatment of the catalyst.
2TABLE 2 Time Time Example (air) (steam) GAL Chlorine No. Temp.
.degree. C. min. min. Index Ppmw 1 N/A N/A N/A >50 33800 12 600
120 0 1090 13 650 120 0 1.85 730 14 700 120 0 350 15 600 120 120
250 16 650 120 120 1.48 140 17 700 120 120 30 18 600 240 0 830 19
650 240 0 590 20 700 240 0 290 21 600 240 240 220 22 650 240 240
160 23 700 240 240 40 24 650 120 15 230 25 650 120 30 120 26 650
120 60 70 27 600 120 240 2.24 150 28 650 120 240 1.62 40
[0129] The attrition properties of Examples 1 13 16 27 and
28 are listed in Table 2. Attrition properties of catalysts can
be defined by the Gross Attrition Loss (GAL) Index. The smaller
the GAL Index the more resistant to attrition is the catalyst. The
GAL Index is measured in the following manner. About 6.0.+-.0.1
g of SAPO catalyst was added to an attrition cup of an attrition
apparatus known in the art. 23700 scc/min of nitrogen gas was bubbled
through a water-containing bubbler to humidify the N.sub.2. The
wet nitrogen passed through the attrition cup and exited the attrition
apparatus through a porous fiber thimble. This thimble separates
the fine catalyst particles resulting from the attrition of the
catalyst particles in the attrition cup as the catalyst particles
are circulated in the attrition cup by the fast flowing nitrogen
gas. The pore size of the thimble determines the size of the fine
particles that are separated from the catalyst. The pore size of
the thimble used to measure the GAL Index was less than about 2
.mu.m.
[0130] The nitrogen flow passing through the attrition cup was
maintained for 60 minutes. The contents of the attrition cup were
transferred to an elutriation cup. The elutriation cup is designed
not to cause further attrition of the catalyst particles, but to
remove any fine particles remaining in the attrition cup so that
the fine particles may be included in the GAL Index. 23700 scc/min
of nitrogen gas was passed through the elutriation cup for 30 minutes.
Additional fine particles were separated by the thimble. The collection
of fine SAPO particles separated by the thimble were weighed. The
amount in grams of fine particles divided by the original amount
of catalyst added to the attrition cup is the GAL Index.
GAL Index=C/(B+C).times.100
[0131] wherein
[0132] B=weight of catalyst in elutriation cup
[0133] C=weight of collected fine catalyst particles
[0134] Having now fully described this invention, it will be appreciated
by those skilled in the art that the invention can be performed
within a wide range of parameters within what is claimed, without
departing from the spirit and scope of the invention. |