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Improved Hydrocracking Performance by Combining Conventional Hydrotreating and Zeolitic Catalysts in Stacked Bed Reactors
M.1,. Occelli and R.G. Anthony (Editors), Advances in Hydrotreating Catalysts 0 1989 Elsevier Science Puhlishers R.V.. Amsterdam - Printed in The Netherlands
263
IMPROVED HYDROCRACKING PERFORMANCE BY COMBINING CONVENTIONAL HYDROTREATING AND ZEOLITIC CATALYSTS IN STACKED BED REACTORS
A.A. ESENER and I.E. MAXWELL Koninklijke/Shell-Laboratorium, Amsterdam (Shell Research B.V.), P.O. Box 3003, 1003 AA Amsterdam, The Netherlands
ABSTRACT A stacked bed hydrocracker reactor configuration composed of conventional hydrotreating and zeolite-based catalysts is shown to offer marked improvements in performance compared to single bed systems. Significant gains in overall hydrocracking activity are achieved, together with good catalyst stability, which is characteristic of zeolitic catalysts. The overall hydrocracking and hydrodenitrogenation kinetics can be described using Langmuir-Hinshelwood type kinetics. Inter-catalyst organic nitrogen levels are shown to play an important role due to their strong inhibiting effect on the activity of the zeolite catalyst. A newly developed zeolitic catalyst (S703) is shown to exhibit a marked improvement in middle distillate selectivity compared to previous, more conventional zeolite-based systems (S753). The product qualities obtained are shown to be quite acceptable, particularly at high conversion levels. INTRODUCTION Hydrocracking
s
an oil conversion process of growing importance in view of
the trend towards increasing the middle distillate/gasoline ratio in refineries. This trend is particularly evident in the rapidly developing countries, for example, in the Pacific Basin and the Indian continent. Even in North America and particularly in the United States, where gasoline continues to be a dominant refining product, hydrocracking is expected to become a major
conversion-upgrading process. Furthermore, hydrocracking is complementary to catalytic cracking, particularly in view of the envisaged future fuel specifications and restrictions on total aromatics in diesel and sulphur specifications. The hydrocracking process is typified by complex chemistry and normally consists of two separate stages (refs. 1-3). The first stage is primarily a hydrotreatment step involving hetero-atom removal reactions
(S &
N) and
hydrogenation of aromatic structures with only a limited amount of cracking. The actual hydrocracking reactions are carried out primarily in the second stage over a bifunctional catalyst containing both hydrogenation and acidic
264
components. In the usual two-stage configuration the first-stage products are sent to an inter-stage separation unit and the second stage therefore receives relatively clean feedstocks. A more cost effective process configuration is series flow in which all the first-stage products are sent directly to the second reactor stage. This type of operation only became possible with the advent of zeolitic catalysts, which show high activity and stability in the presence of NH3. Zeolitic catalysts have also been used (by replacing part of the bottom fraction of the hydrotreating catalyst: stacked bed) in the first-stage reactors or in single-stage hydrocracking to improve the cracking activity, particularly in mild hydrocracking applications (refs. 4 - 6 ) . Single-stage hydrocracking is the simplest process configuration in which once-through flow of the feed (typically straight-run or processed flashed distillates and deasphalted oils (DAO))
results in a conversion to distillate
products (e.g. <370 OC b.p.) of 40 to 80 %w on feed. The heavier, partially converted product fraction is deeply hydrogenated and is particularly suitable as a feedstock for an ethylene cracker or a catalytic cracker (refs. 6-7). The implementation of these stacked bed systems does, however, requires a good understanding of the catalyst systems (both amorphous and zeolitic), process chemistry and technology to ensure optimal application and process integration. The present study is intended to provide more insight into the interaction of the various factors which need to be assessed, in particular as they relate to highly active zeolite-based cracking catalysts applied in stacked beds. EXPERIMENTAL The experimental data were mainly collected from bench-scale experiments (120 g catalyst) in the high-pressure (120-150 bar) range. In the stacked bed the top and bottom catalysts were operated at the same temperature (Fig. 1). The feedstocks used were typically Arabian Heavy type straight-run flashed distillates with end points above 600
OC.
Organic nitrogen contents of these
feedstocks were typically around 1000 ppmw, with a sulphur content of 2-3 %w. Catalysts used were proprietary Shell catalysts, which are commercially available. The hydrotreatingfiydrocracking catalyst used in the top fraction o f the stacked bed was S324 (NiMo amorphous), while zeolitic hydrocracking catalysts i.e. S753 and S703 (Ni/W/zeolite Y), were used at the bottom of the bed.
265 FEED
i
NF
N I
-
(INTERSTAGE)
-
X
T EFFLUENT
Fig. 1. Stacked bed reactor configuration. A = amorphous hydrotreating catalyst (S324) B zeolitic hydrocracking catalyst (S753 or 5703) NF = nitrogen in feed interstage nitrogen NI NE = effluent nitrogen
-
The net conversion to <370 OC products and the selectivity to a certain cut were calculated according to the following formulae (1) and ( 2 ) :
-
Net conversion to <370 OC (%w)
(1 - (>370 OC in products / >370 OC in feed)) Selectivity to fraction A(%w)
*
100
(amount of A in products / products below 370 OC)
*
(1) 100
(2)
All the boiling points refer to TBP-GLC results. ACTIVITY It is well known that the intrinsic cracking activity k-HC of the zeolitic catalysts is very high compared to that of amorphous catalysts. However, under certain hydrocracking conditions zeolites, like any other solid acid catalyst, are sensitive to the presence of ammonia and, more importantly, organic nitrogen. Nitrogen effectively reduces the activity by neutralising the active sites (ref. 3).
266
Therefore, for use in stacked bed service the influence of the inter-stage organic nitrogen (Fig. 1) needs to be understood before the optimal top and bottom catalyst volume ratios can be selected. On the basis of previous studies (ref. 8 ) we have therefore analysed the performance of a stacked bed of
S324/S753. Both the hydrodenitrogenation (HDN) and hydrocracking (HC) activities were described by a simple Langmuir-Hinshelwood type model relating the observed (apparent) activity to the interstage catalyst poison (organic nitrogen) concentration (Fig. 2). The effect of ammonia was not considered since under the relevant operating conditions it can be assumed constant. On
,(q k-HDN
- 1.2 - 1.0 - 0.8
- 0.6 - 0.4
0
50
150
100 [N]
wmw
Fig. 2. Hydrocracking (HC) and hydrodenitrogenation (HDN) activity for the S 7 5 3 concentration at constant catalyst as a function of organic nitrogen (“1) temperature (K normalized first order reaction rate constant: ton m-3h-1)
-
the basis of the model description and kinetic measurements both the HC and HDN activities of a stacked bed can now be calculated as a function of the catalyst volume ratios. As shown in Fig. 3 , the first order rate constant for hydrocracking shows an optimum at about 20-30 %v of S753 catalyst at the given constant temperature.
267
K-HC (g/(L.h) C 370 OC
I50
I00
50
O/o
50 1 1 ZEOLlTlC CAT IN STACKED BED
0
Fig. 3 . Hydrocracking activity and interstage/effluent organic nitrogen concentrations as a function of the volumetric stacked bed catalyst ratio. normalized hydrocracking reaction rate constant) (K-HC
-
Similarly, the HDN efficiency shows an optimum, i.e. minimum total nitrogen in the stacked bed effluent, at the same volume ratio. At a higher zeolite ratio the interstage nitrogen becomes more than desired and the relative activity advantage of the zeolitic catalyst decreases and eventually becomes even less than that of the amorphous catalyst because of the nitrogen poisoning. The optimum volume ratio is obviously also dependent on the temperature, e.g. the influence of the nitrogen adsorption phenomenon will be relatively smaller at higher temperatures. Because of the nitrogen effects discussed above the stacked beds display very high apparent activation energies (Table 1). This means that the reactor temperature needs to be carefully controlled by means of the appropriate quench systems.
268 TABLE 1
1
1
Apparent energy of activation for hydrocracking (HC) and hydrodenitrogenation (HDN) for amorphous and stacked bed systems EA (kJ/mol) HC <370 OC HDN
I
Catalyst system
SIlI
S32411153
155
S324/S703 440
240
255
SELECTIVITY ASPECTS For heavy feedstocks, as used i n this study, amorphous catalysts are, in general, the most selective to middle distillates, while zeolitic systems tend to produce more light products. However, modification of the zeolite properties enables the product yield structure to be influenced advantageously, as is illustrated in Fig. 4. The S324 system gives the highest middle distillate (MD)
MIDDLE DISTILLATE ( 1 5 0 - 37OoC1 SELECT IV ITYI '10w
-loot
401-
IHYDROTREATING (S324)
A S324/S703 STACKED BED S324/S753 STACKED BED
0
0
20
40
80 100 CONVERSION TO < 370 o C l o / o ~ 60
Fig. 4 . Middle distillate selectivity (150-370 OC b.p.) as a function of conversion per pass (to <370 OC b.p.) for single and stacked bed catalyst systerns.
269
selectivity and the S324/S753 gives the lowest. The stacked bed with the improved zeolite, 5 7 0 3 , is substantially more MD selective than the S753 system. The loss in MD selectivity for both zeolitic systems at high conversion levels relative to S324 is due to the increasing contribution of secondary cracking mechanisms, which is also accompanied by an increase in gas make. However, at modest conversion levels of, for example 4 0 to 7 0 %w per pass, attractive product yields and quality can be obtained with the stacked bed systems with significant activity advantages. This is shown in Table 2 , where the S324 catalyst is the most selective for kerosine and gas oil fractions. TABLE 2 Comparison of product selectivities Catalyst Conversion to
<370 OC, %w
I
S324
50
S324/S703
70
3 24/S7 5 3
70
-
-
4
4
5
Naphtha/gasoline
23
45
56
Keros ine
21
25
20
26
19
T-30
T-32
Gas
Gas oil
52
-
43
Activity*
T req
*
OC
T
-
T required for the given conversion at a space velocity of 1 kg/(l.h)
However, it does exhibit a slightly higher gas make than the stacked beds, which can mcst likely be attributed to the actual operating temperature required for a given conversion level being much higher with the S-324 catalyst than with the stacked bed systems. It is also shown that by appropriate choice of the stacked bed system the product package can be adjusted to some degree. The substantial activity gains of the stacked beds, particularly at high temperatures (high conversion loads), are due to the high apparent (measured) energies of activation for these systems. PRODUCT QUALITY ASPECTS In general, at moderate conversion loads, the S324 hydrotreating catalyst gives better product quality (particularly with kerosine and gas oil) than stacked beds. However, at high conversion loads, because of its relatively low activity the S324 system has to be operated at very high temperatures as a
270
result of which polyaromatics hydrogenation can become incomplete due to the less favourable thermodynamics. Table 3 shows some product property data collected with the S703 stacked bed system. In general, the product quality improves with increasing conversion. The partially converted residue (>370
OC
fraction) is deeply hydrogenated at or above 6 0 %w conversion and is considered to be an excellent catalytic cracker or ethylene cracker feedstock. The tops fractions from stacked beds are expected to be good gasoline blending components because of their high iso/normal ratio. TABLE 3 Product Properties S324/S703 stacked bed
Conversion to O C (%w)
<370
45
61
91
14.65 0.735
14.93 0.726
13.36 0.830 17
13.66 0.818 20
14.04 0.798 24
H (%W) 1 3 . 4 3 Density (g/ml) 0 . 8 6 0 -1 2 Pour pt. (OC)
13.85 0.840 -9
0.817
14.07 0.852 42
14.36 0.838 39
Nauh tha/Gas01ine
H (%W) 1 4 . 3 1 Density* (g/ml) 0 . 7 4 0 Kerosine
H (%W) Density (g/ml) Smoke pt. (mm) Gas oil
14.33 -9
Residue
*
H (%W) Density (g/ml) Pour pt. (OC)
13.96 0.861 45
(density 2 0 / 4
basis)
OC
DISCUSSION AND CONCLUSIONS It has been demonstrated that stacked bed systems composed of conventional hydrotreating and zeolite-based catalysts can offer significant improvements in performance compared to single-bed systems. Substantial gains in overall activity for hydrocracking can be achieved. Further, the low coke forming characteristics of the zeolitic component offer significant improvements in overall catalyst stability.
211
The use of a newly developed zeolite-based catalyst (S703) in a stacked bed has resulted in a marked improvement in middle distillate selectivity compared to previous, more conventional zeolite-based catalyst systems (S753). The product qualities obtained are quite acceptable particularly at high conversion levels. Further, it has been shown that both the hydrocracking and the hydrodenitrogenation reactions can be quite well described by means of Langmuir-Hinshelwood type expressions. On the basis of such a relatively simple model the optimal ratio of hydrotreating to zeolite catalyst beds can be calculated for a given desired overall performance. The implementation of these stacked bed systems does, however, require a thorough understanding of the catalysts, kinetics, product quality and process technology for optimal application and integration in the refinery. REFERENCES I.E. Maxwell, Catalysis Today, 1 (1987) 385. ACS Symposium Series 20, April 1975, J.W. Ward and S.A. Qadar (Eds.). P.J. Nat, Paper presented at the NPRA Annual Meeting, AM-88-75, March 20-22, 1988. U . S . Patent 4,534,852. J.W. Gosselink, A. van de Paverd and W.H.J. Stork, “Mild Hydrocracking: Optimization of multiple catalyst systems for increased vacuum gas oil conversion”, paper to be presented to at the “Catalyst in Petroleum Refining Conference”, Kuwait, 4-8 March 1989. C.T. Adams, D.M. Washcheck, R.H. Stade and W.J. Daniels, Hydrocarbon Processing, May 1986, p. 46. P.H. Desai, M.Y. Asim, F.W. van Houtert and P.J. Nat, Oil and Gas Journal, July 22, 1985, p. 106. I.E. Maxwell and J.A. van de Griend, “New Developments in Zeolite Science and Technology” Proceedings of the 7th International Zeolite Conference, Y. Murakami et al. (Eds.) (1986), p. 795.
263
IMPROVED HYDROCRACKING PERFORMANCE BY COMBINING CONVENTIONAL HYDROTREATING AND ZEOLITIC CATALYSTS IN STACKED BED REACTORS
A.A. ESENER and I.E. MAXWELL Koninklijke/Shell-Laboratorium, Amsterdam (Shell Research B.V.), P.O. Box 3003, 1003 AA Amsterdam, The Netherlands
ABSTRACT A stacked bed hydrocracker reactor configuration composed of conventional hydrotreating and zeolite-based catalysts is shown to offer marked improvements in performance compared to single bed systems. Significant gains in overall hydrocracking activity are achieved, together with good catalyst stability, which is characteristic of zeolitic catalysts. The overall hydrocracking and hydrodenitrogenation kinetics can be described using Langmuir-Hinshelwood type kinetics. Inter-catalyst organic nitrogen levels are shown to play an important role due to their strong inhibiting effect on the activity of the zeolite catalyst. A newly developed zeolitic catalyst (S703) is shown to exhibit a marked improvement in middle distillate selectivity compared to previous, more conventional zeolite-based systems (S753). The product qualities obtained are shown to be quite acceptable, particularly at high conversion levels. INTRODUCTION Hydrocracking
s
an oil conversion process of growing importance in view of
the trend towards increasing the middle distillate/gasoline ratio in refineries. This trend is particularly evident in the rapidly developing countries, for example, in the Pacific Basin and the Indian continent. Even in North America and particularly in the United States, where gasoline continues to be a dominant refining product, hydrocracking is expected to become a major
conversion-upgrading process. Furthermore, hydrocracking is complementary to catalytic cracking, particularly in view of the envisaged future fuel specifications and restrictions on total aromatics in diesel and sulphur specifications. The hydrocracking process is typified by complex chemistry and normally consists of two separate stages (refs. 1-3). The first stage is primarily a hydrotreatment step involving hetero-atom removal reactions
(S &
N) and
hydrogenation of aromatic structures with only a limited amount of cracking. The actual hydrocracking reactions are carried out primarily in the second stage over a bifunctional catalyst containing both hydrogenation and acidic
264
components. In the usual two-stage configuration the first-stage products are sent to an inter-stage separation unit and the second stage therefore receives relatively clean feedstocks. A more cost effective process configuration is series flow in which all the first-stage products are sent directly to the second reactor stage. This type of operation only became possible with the advent of zeolitic catalysts, which show high activity and stability in the presence of NH3. Zeolitic catalysts have also been used (by replacing part of the bottom fraction of the hydrotreating catalyst: stacked bed) in the first-stage reactors or in single-stage hydrocracking to improve the cracking activity, particularly in mild hydrocracking applications (refs. 4 - 6 ) . Single-stage hydrocracking is the simplest process configuration in which once-through flow of the feed (typically straight-run or processed flashed distillates and deasphalted oils (DAO))
results in a conversion to distillate
products (e.g. <370 OC b.p.) of 40 to 80 %w on feed. The heavier, partially converted product fraction is deeply hydrogenated and is particularly suitable as a feedstock for an ethylene cracker or a catalytic cracker (refs. 6-7). The implementation of these stacked bed systems does, however, requires a good understanding of the catalyst systems (both amorphous and zeolitic), process chemistry and technology to ensure optimal application and process integration. The present study is intended to provide more insight into the interaction of the various factors which need to be assessed, in particular as they relate to highly active zeolite-based cracking catalysts applied in stacked beds. EXPERIMENTAL The experimental data were mainly collected from bench-scale experiments (120 g catalyst) in the high-pressure (120-150 bar) range. In the stacked bed the top and bottom catalysts were operated at the same temperature (Fig. 1). The feedstocks used were typically Arabian Heavy type straight-run flashed distillates with end points above 600
OC.
Organic nitrogen contents of these
feedstocks were typically around 1000 ppmw, with a sulphur content of 2-3 %w. Catalysts used were proprietary Shell catalysts, which are commercially available. The hydrotreatingfiydrocracking catalyst used in the top fraction o f the stacked bed was S324 (NiMo amorphous), while zeolitic hydrocracking catalysts i.e. S753 and S703 (Ni/W/zeolite Y), were used at the bottom of the bed.
265 FEED
i
NF
N I
-
(INTERSTAGE)
-
X
T EFFLUENT
Fig. 1. Stacked bed reactor configuration. A = amorphous hydrotreating catalyst (S324) B zeolitic hydrocracking catalyst (S753 or 5703) NF = nitrogen in feed interstage nitrogen NI NE = effluent nitrogen
-
The net conversion to <370 OC products and the selectivity to a certain cut were calculated according to the following formulae (1) and ( 2 ) :
-
Net conversion to <370 OC (%w)
(1 - (>370 OC in products / >370 OC in feed)) Selectivity to fraction A(%w)
*
100
(amount of A in products / products below 370 OC)
*
(1) 100
(2)
All the boiling points refer to TBP-GLC results. ACTIVITY It is well known that the intrinsic cracking activity k-HC of the zeolitic catalysts is very high compared to that of amorphous catalysts. However, under certain hydrocracking conditions zeolites, like any other solid acid catalyst, are sensitive to the presence of ammonia and, more importantly, organic nitrogen. Nitrogen effectively reduces the activity by neutralising the active sites (ref. 3).
266
Therefore, for use in stacked bed service the influence of the inter-stage organic nitrogen (Fig. 1) needs to be understood before the optimal top and bottom catalyst volume ratios can be selected. On the basis of previous studies (ref. 8 ) we have therefore analysed the performance of a stacked bed of
S324/S753. Both the hydrodenitrogenation (HDN) and hydrocracking (HC) activities were described by a simple Langmuir-Hinshelwood type model relating the observed (apparent) activity to the interstage catalyst poison (organic nitrogen) concentration (Fig. 2). The effect of ammonia was not considered since under the relevant operating conditions it can be assumed constant. On
,(q k-HDN
- 1.2 - 1.0 - 0.8
- 0.6 - 0.4
0
50
150
100 [N]
wmw
Fig. 2. Hydrocracking (HC) and hydrodenitrogenation (HDN) activity for the S 7 5 3 concentration at constant catalyst as a function of organic nitrogen (“1) temperature (K normalized first order reaction rate constant: ton m-3h-1)
-
the basis of the model description and kinetic measurements both the HC and HDN activities of a stacked bed can now be calculated as a function of the catalyst volume ratios. As shown in Fig. 3 , the first order rate constant for hydrocracking shows an optimum at about 20-30 %v of S753 catalyst at the given constant temperature.
267
K-HC (g/(L.h) C 370 OC
I50
I00
50
O/o
50 1 1 ZEOLlTlC CAT IN STACKED BED
0
Fig. 3 . Hydrocracking activity and interstage/effluent organic nitrogen concentrations as a function of the volumetric stacked bed catalyst ratio. normalized hydrocracking reaction rate constant) (K-HC
-
Similarly, the HDN efficiency shows an optimum, i.e. minimum total nitrogen in the stacked bed effluent, at the same volume ratio. At a higher zeolite ratio the interstage nitrogen becomes more than desired and the relative activity advantage of the zeolitic catalyst decreases and eventually becomes even less than that of the amorphous catalyst because of the nitrogen poisoning. The optimum volume ratio is obviously also dependent on the temperature, e.g. the influence of the nitrogen adsorption phenomenon will be relatively smaller at higher temperatures. Because of the nitrogen effects discussed above the stacked beds display very high apparent activation energies (Table 1). This means that the reactor temperature needs to be carefully controlled by means of the appropriate quench systems.
268 TABLE 1
1
1
Apparent energy of activation for hydrocracking (HC) and hydrodenitrogenation (HDN) for amorphous and stacked bed systems EA (kJ/mol) HC <370 OC HDN
I
Catalyst system
SIlI
S32411153
155
S324/S703 440
240
255
SELECTIVITY ASPECTS For heavy feedstocks, as used i n this study, amorphous catalysts are, in general, the most selective to middle distillates, while zeolitic systems tend to produce more light products. However, modification of the zeolite properties enables the product yield structure to be influenced advantageously, as is illustrated in Fig. 4. The S324 system gives the highest middle distillate (MD)
MIDDLE DISTILLATE ( 1 5 0 - 37OoC1 SELECT IV ITYI '10w
-loot
401-
IHYDROTREATING (S324)
A S324/S703 STACKED BED S324/S753 STACKED BED
0
0
20
40
80 100 CONVERSION TO < 370 o C l o / o ~ 60
Fig. 4 . Middle distillate selectivity (150-370 OC b.p.) as a function of conversion per pass (to <370 OC b.p.) for single and stacked bed catalyst systerns.
269
selectivity and the S324/S753 gives the lowest. The stacked bed with the improved zeolite, 5 7 0 3 , is substantially more MD selective than the S753 system. The loss in MD selectivity for both zeolitic systems at high conversion levels relative to S324 is due to the increasing contribution of secondary cracking mechanisms, which is also accompanied by an increase in gas make. However, at modest conversion levels of, for example 4 0 to 7 0 %w per pass, attractive product yields and quality can be obtained with the stacked bed systems with significant activity advantages. This is shown in Table 2 , where the S324 catalyst is the most selective for kerosine and gas oil fractions. TABLE 2 Comparison of product selectivities Catalyst Conversion to
<370 OC, %w
I
S324
50
S324/S703
70
3 24/S7 5 3
70
-
-
4
4
5
Naphtha/gasoline
23
45
56
Keros ine
21
25
20
26
19
T-30
T-32
Gas
Gas oil
52
-
43
Activity*
T req
*
OC
T
-
T required for the given conversion at a space velocity of 1 kg/(l.h)
However, it does exhibit a slightly higher gas make than the stacked beds, which can mcst likely be attributed to the actual operating temperature required for a given conversion level being much higher with the S-324 catalyst than with the stacked bed systems. It is also shown that by appropriate choice of the stacked bed system the product package can be adjusted to some degree. The substantial activity gains of the stacked beds, particularly at high temperatures (high conversion loads), are due to the high apparent (measured) energies of activation for these systems. PRODUCT QUALITY ASPECTS In general, at moderate conversion loads, the S324 hydrotreating catalyst gives better product quality (particularly with kerosine and gas oil) than stacked beds. However, at high conversion loads, because of its relatively low activity the S324 system has to be operated at very high temperatures as a
270
result of which polyaromatics hydrogenation can become incomplete due to the less favourable thermodynamics. Table 3 shows some product property data collected with the S703 stacked bed system. In general, the product quality improves with increasing conversion. The partially converted residue (>370
OC
fraction) is deeply hydrogenated at or above 6 0 %w conversion and is considered to be an excellent catalytic cracker or ethylene cracker feedstock. The tops fractions from stacked beds are expected to be good gasoline blending components because of their high iso/normal ratio. TABLE 3 Product Properties S324/S703 stacked bed
Conversion to O C (%w)
<370
45
61
91
14.65 0.735
14.93 0.726
13.36 0.830 17
13.66 0.818 20
14.04 0.798 24
H (%W) 1 3 . 4 3 Density (g/ml) 0 . 8 6 0 -1 2 Pour pt. (OC)
13.85 0.840 -9
0.817
14.07 0.852 42
14.36 0.838 39
Nauh tha/Gas01ine
H (%W) 1 4 . 3 1 Density* (g/ml) 0 . 7 4 0 Kerosine
H (%W) Density (g/ml) Smoke pt. (mm) Gas oil
14.33 -9
Residue
*
H (%W) Density (g/ml) Pour pt. (OC)
13.96 0.861 45
(density 2 0 / 4
basis)
OC
DISCUSSION AND CONCLUSIONS It has been demonstrated that stacked bed systems composed of conventional hydrotreating and zeolite-based catalysts can offer significant improvements in performance compared to single-bed systems. Substantial gains in overall activity for hydrocracking can be achieved. Further, the low coke forming characteristics of the zeolitic component offer significant improvements in overall catalyst stability.
211
The use of a newly developed zeolite-based catalyst (S703) in a stacked bed has resulted in a marked improvement in middle distillate selectivity compared to previous, more conventional zeolite-based catalyst systems (S753). The product qualities obtained are quite acceptable particularly at high conversion levels. Further, it has been shown that both the hydrocracking and the hydrodenitrogenation reactions can be quite well described by means of Langmuir-Hinshelwood type expressions. On the basis of such a relatively simple model the optimal ratio of hydrotreating to zeolite catalyst beds can be calculated for a given desired overall performance. The implementation of these stacked bed systems does, however, require a thorough understanding of the catalysts, kinetics, product quality and process technology for optimal application and integration in the refinery. REFERENCES I.E. Maxwell, Catalysis Today, 1 (1987) 385. ACS Symposium Series 20, April 1975, J.W. Ward and S.A. Qadar (Eds.). P.J. Nat, Paper presented at the NPRA Annual Meeting, AM-88-75, March 20-22, 1988. U . S . Patent 4,534,852. J.W. Gosselink, A. van de Paverd and W.H.J. Stork, “Mild Hydrocracking: Optimization of multiple catalyst systems for increased vacuum gas oil conversion”, paper to be presented to at the “Catalyst in Petroleum Refining Conference”, Kuwait, 4-8 March 1989. C.T. Adams, D.M. Washcheck, R.H. Stade and W.J. Daniels, Hydrocarbon Processing, May 1986, p. 46. P.H. Desai, M.Y. Asim, F.W. van Houtert and P.J. Nat, Oil and Gas Journal, July 22, 1985, p. 106. I.E. Maxwell and J.A. van de Griend, “New Developments in Zeolite Science and Technology” Proceedings of the 7th International Zeolite Conference, Y. Murakami et al. (Eds.) (1986), p. 795.