1278
I n d . Eng. C h e m . Res. 1989, 28, 1278-1284
Mild Hydrocracking of Bitumen-Derived Coker and Hydrocracker Heavy Gas Oils: Kinetics, Product Yields, and Product Propertied Sok M. Yui* and E m e r s o n C. S a n f o r d Research Department, Syncrude Canada Ltd., P.O. Box 5790, Edmonton, Alberta, Canada T6C 4G3
Bitumen-derived coker and hydrocracker heavy gas oils were hydrotreated at 350-400 "C, 7-11 MPa, 0.7-1.5 h-' LHSV, and 600 S m3 of Hz/m3 of feed in a pilot-scale trickle-bed reactor, over presulfided commercial NiMo/A1203 catalysts. T h e conversion of HGO materials (343+ "C) in feed to naphtha (195- "C) and LGO (195/343 "C) was determined by gas chromatographic simulated distillation. T h e degree of conversion was analyzed with modified first-order kinetics, which incorporate power terms for LHSV and hydrogen partial pressure. The equations were based on three cracking schemes; parallel, consecutive, and combined parallel-consecutive conversion. Conversion could be described reasonably well by either the parallel scheme or the consecutive scheme. Total liquid product yields could be estimated by power forms of operating parameters. Hydrodesulfurization and hydrodenitrogenation obey 1.5th- and first-order kinetics, respectively. A good relationship between the sulfur and nitrogen contents of LGO and HGO cuts versus those of total liquid products was obtained. Syncrude Canada Ltd. operates a surface mining oil sand plant at the Athabasca oil sands deposit in northern Alberta and produces synthetic crude oil from extracted bitumen. The bitumen is currently upgraded in two fluid cokers and an ebullated-bed hydrocracker, followed by hydrotreating of the naphtha and light and heavy gas oils. In the mid 1980s, a long-term middle-distillate shortage was projected, and we conducted extensive studies on mild hydrocracking (MHC) of heavy gas oil with a view to shifting the product slate toward increased middle-distillate production. The results from the tests were encouraging; however, a revised study showed that the shortage was less urgent than projected, and the plans for MHC were shelved. Nevertheless, the pilot studies led to useful findings, some of which are explored in this report. MHC is defined as once-through catalytic hydroconversion of heavier materials in a feed into lighter materials at milder operating conditions than conventional hydrocracking. Compared to hydrotreating, however, the MHC mode of operation requires higher reactor temperatures. It produces a greater yield of middle distillates as well as more naphtha and gases. The catalysts may be specially developed silica-alumina-based catalyst or the same as the conventional alumina-based hydrotreating catalyst. There is extensive literature on MHC, reported from the viewpoint of petroleum refiners, process developers, and catalyst suppliers (e.g., Desai et al. (1985), Elkes et al. (1987), Gembicki et al. (1983), Kalnes et al. (1984), Plantenga and Sonnemans (1983)). Wilson et al. (1987) conducted pilot-plant hydrocracking experiments using the 385/525 "C fraction of synthetic crude oil (hydrotreated fluid coker products from Athabasca bitumen) with NiW/Si02-A1203catalysts. The authors reported product yields and properties and discussed an optimum condition to produce a maximum distillate yield. Reports on the kinetics are scarce. In the present study, experiments were conducted in a pilot-scale trickle-bed reactor to investigate MHC kinetics, as well as product yields and properties. We used Athabasca bitumen-derived coker and hydrocracker heavy gas
* Author
to whom correspondence should be addressed. This work was presented a t t h e Division of Industrial and Engineering Chemistry, American Chemical Society Meeting, Toronto, J u n e 1988.
0888-588518912628-1278$01.50/0
oils (HGO) as the feed and commercial NiMo/A1203hydrotreating catalysts. The degree of conversion of heavier materials in the feed into light gas oil (LGO) and naphtha was analyzed with modified first-order kinetic equations, which incorporate power terms for liquid hourly space velocity (LHSV) and hydrogen partial pressure. The equations were based on three cracking schemes: (1)parallel conversion from HGO to LGO and naphtha, (2) consecutive conversion from HGO to LGO and then LGO to naphtha, and (3) combined parallel and consecutive conversion. Cracking to gases was not studied in the present report. Correlations to predict total liquid product (TLP) yields were developed. Kinetics of hydrodesulfurization (HDS) and hydrodenitrogenation (HDN) were also studied. The feeds and TLPs were distilled by spinning band distillation into naphtha, LGO, and HGO. Density, sulfur and nitrogen contents, aniline point, and simulated distillation were determined. Correlations to predict sulfur and nitrogen contents in the distilled products were developed as a function of corresponding properties of TLP. Employing a newly developed correlation (Yui and Sanford, 1988), cetane numbers of the LGO cut were estimated. Experimental Section Two commercial NiMo/A1,03 catalysts (K and S ) were used in this study. Catalysts K and S have similar metal loading (about 4 wt 70 NiO and 20 wt % Moo3), pore volume (about 0.4 cm3/g), and surface area (about 160 m2/g). The pilot reactor system (1.7-cm i.d. and 122-cm overall length) has three independent and identical downflow catalyst beds of 82 cm each. The 120 cm3 of catalysts was diluted with 50 vol % of 45-mesh silicon carbide. Each reactor was heated by three separately controlled electrical furnaces. The temperature was measured by movable thermocouples on the outside reactor skin. In a separate test using movable thermocouples a t the reactor center and skin, it was shown that the radial temperature gradient was within 1 "C. The heaters were controlled to give a uniform temperature throughout the reactor length, and a weight-average number was used as the isothermal temperature. Figure 1illustrates a typical temperature profile. The detailed reactor system and experimental methods are described elsewhere (Yui, 1989). Feedstocks were bitumen-derived HGOs from a com1989 American Chemical Society
Ind. Eng. Chem. Res., Vol. 28, No. 9, 1989 1279 Table I. Properties of Feeds and Products at Typical Operating Conditions (A) feed, coker HGO; catalyst, NiMo/A1203-K run no. operating conditions temp, "C pressure, MPa LHSV, v/h/v yields, m3/m3 of feed density a t 20 "C, g/cm3 sulfur, wppm nitrogen, wppm carbon, wt % hydrogen, wt % simulated distillation, "C IBP 5% 10% 30 % 50 % 70% 90 70 95 %
FBP 195- "C, % 343- "C, % Inlet
r--r---
~'
,
EUll 0
feed
1
5
8
10
11
0.9983 45100 3282 85.01 9.96
379 8.8 0.7 1.036 0.9193 1240 565 87.82 11.68
389 8.8 0.7 1.045 0.9113 421 246 87.82 11.89
399 9.5 0.7 1.053 0.9018 213 103 87.65 11.83
399 8.8 0.7 1.047 0.9072 323 246 87.92 11.84
232 302 325 374 413 450 496 517 551 0 15.9
96 215 261 332 375 418 473 495 532 3.8 34.8
91 190 239 318 363 407 467 492 542 5.4 40.7
78 157 210 301 350 396 458 485 536 8.3 46.8
76 167 221 309 356 402 463 490 542 7.2 43.8
*
.. *
)?
li
p
j-Cotoiyst
+c
Bed
i
r,
1 9,
I
8
10
12
14
400 8.8 1.0 1.038 0.9149 716 560 87.88 11.65
0.9402 15200 2850 86.70 11.37
360 8.9 0.7 1.025 0.9091 936 501 87.63 12.32
370 8.8 1.0 1.024 0.9089 696 504 87.58 12.20
380 7.0 1.5 1.019 0.9125 916 889 87.83 12.11
379 8.9 0.7 1.035 0.9002 122 100 87.68 12.56
390 8.8 1.0 1.031 0.9010 196 201 88.07 12.37
82 185 237 319 365 409 470 497 550 5.7 40.1
291 319 330 359 382 408 448 469 522 0 18.0
132 282 307 345 371 399 441 463 524 1.3 28.9
130 280 306 345 371 398 440 461 515 1.4 28.6
127 279 307 346 372 399 441 461 514 1.5 28.1
104 250 290 339 366 395 437 457 511 2.6 33.1
100 240 286 338 366 394 436 457 511 3.1 33.4
Yields of Total Liquid Products
'Y. 16t
Mesh SIC
7-16
I*
I
3
Lgmm ~ ~ - ~ -*..-... ---.____.
I?
- _ Wire
feed
spinning band distillation of feeds and TLPs, all at typical operating conditions.
.._ -*..
G l o s s Beods
-f
(B) feed, hydrocracker HGO; catalyst, NiMo/A1203-S
.....';( 0
,e Avg Bed
/I6 Mesh SIC d v_ / _
.
300
1
'
320
Temperature
,*/'
'
..-;' ?
340
,
!I
1
360
359'C I
380
Since measurement of TLPs from our pilot unit was not sufficiently accurate, the needed data were obtained from a forced carbon balance by analyzing feeds, gases, and liquid products (ignoring any dissolved gases) and by measuring feed and gas rates. The T L P volume yields (m3/m3of feed) so obtained are shown in Table I. Yields by weight were typically 95% from coker HGO and 99% from hydrocracker HGO. The lower value for coker HGO is mainly due to the higher content of sulfur material in the feed, which converts to gaseous products. It was found that the T L P volume yields can be correlated by power forms of temperature ( t , "C), hydrogen
Figure 1. Typical reactor temperature profile.
mercial fluid coker and a pilot ebullated-bed hydrocracker. Feed properties are summarized as part of Table I. It is observed that, compared to coker HGO, the hydrocracker HGO used in the present study has a lower density and viscosity, lower sulfur and nitrogen contents, a lower carbonlhydrogen ratio, and a smaller amount of heavy materials. Experiments were undertaken by varying the reactor temperature (350-400 "C), LHSV (0.7-1.5 h-l), and pressure (7-10 MPa). The system was once-through, and the gas was not recycled. The hydrogen was of 100% purity, and the measured pressure was the reactor total pressure. The hydrogenlfeed ratio was maintained constant (600 S m3/m3) throughout the study. The feeds and total liquid products were distilled by spinning band distillation into naphtha (195- "C), LGO (1951343 "C), and HGO (343+ "C), and the properties of each cut were determined. Measurements of the properties are described in detail elsewhere (Yui, 1989).
Experimental Results Table I summarizes the properties of the feeds and TLPs, and Table I1 gives the typical results from the
where Y = yield in m3/m3 of feed. In the present study, hydrogen partial pressure was assumed to be the reactor total pressure, because the hydrogen used was of 100% purity. This follows the industry practice in which all effects other than purity of recycled hydrogen are ignored. Table I11 summarizes the resulting coefficients. It is noticed that the effect of temperature is the most significant.
Kinetic Equations of MHC Because this study was mostly concerned with liquid products, our pilot unit was not equipped to measure dissolved gas production to an accuracy sufficient for analysis of gas formation kinetics. Overall material balance closure including measured gas products (again ignoring gases dissolved in the liquid products) was typically 97.5% for both coker and hydrocracking HGOs. The extent of conversion of HGO materials (343+ "C) to naphtha (195"C) and LGO (195/353 "C) was determined by gas chromatographic simulated distillation (ASTM D2887) of feeds and TLPs. In the parallel-consecutive scheme, we assumed that HGO materials (A) in the feed convert first
1280 Ind. Eng. Chem. Res., Vol. 28, No. 9, 1989 Table 11. Yields and Properties of Various Cuts from Feeds and Hydrotreated Products (Catalyst: NiMo/A1108-S) (A) feed: coker HGO (B)feed: hvdrocracker HGO run no. operating conditions temp, "C pressure, MPa LHSV, v/h/v naphtha cut (IBP/195 "C) recovery, wt % density a t 20 "C, g/cm3 sulfur, wppm nitrogen, wppm simulated distillation, "C IBP 5% 10% 50 % 90 % 95 7@ FBP LGO cut (195/343 "C) recovery, wt % density a t 20 "C, g/cm3 sulfur, wppm nitrogen, wppm aniline pt, "C cetane index" simulated distillation, "C IBP 5 7@ 10% 50 % 90 % 95 % FBP HGO cut (343+ "C) recovery, wt % density a t 20 "C, g/cm3 sulfur, wppm nitrogen, wppm simulated distillation, "C IBP 5% 10% 50 % 90 7 e 95 % FBP hangup, wt %
feed
1
2
3
4
351 8.8 0.8
360 8.8 0.8
380 8.8 0.8
400 8.8 0.8
1.20 0.7953 88 171
1.66 0.7933 52 78
3.46 0.7866 30 16
7.24 0.7834 54 5
91 101 133 159 197 205 242
89 100 130 153 194 202 221
80 93 120 145 189 197 217
77 91 119 144 190 199 217
7.87 0.9455 40250 1118 room temp 25.6*
21.62 0.9067 1800 579 29.9 29.5
23.01 0.9044 1050 314 30.6 29.7
27.91 0.9003 279 80 31.3 30.2
34.63 0.8980 41 18 32.2 30.1
183 243 264 315 34 1 347 371
192 227 246 311 343 348 366
187 223 242 309 342 348 365
188 217 235 305 341 347 364
91.91 1.0115 45100 4073
76.71 0.9554 8000 2501
74.81 0.9504 5410 1934
325 347 36 1 435 517 536 579 0.12
334 352 362 425 507 528 575 0.25
330 350 360 423 505 527 575 0.33
0
3
8
10
12
14
360 8.9 0.7
370 8.8 1.0
380 7.0 1.5
379 8.9 0.7
390 8.8 1.0
1.69 0.8136 68 23
2.04 0.8162 17 22
0.61 nae na na
1.37 0.8023 20 5
1.59 0.8008 58 5
70 99 103 174 231 238 253
66 98 102 177 232 240 259
109 118 127 157 183 190 213
99 115 122 155 183 189 207
97 113 119 154 183 188 207
11.02 0.9186 11200 1743 45.4 29.4
21.44 0.9060 254 87 42.8 31.1
21.08 0.9058 251 80 42.3 31.1
20.75 0.9041 227 277 39.2 31.0
25.24 0.8968 27 10 45.5 33.0
25.12 0.8975 81 19 42.4 32.3
182 210 226 297 339 346 364
211 300 307 331 351 357 380
233 262 276 320 342 347 366
234 262 276 320 342 348 366
177 228 261 319 343 349 370
180 225 251 314 341 347 367
180 221 247 314 340 346 365
68.24 0.9405 1790 1113
58.98 0.9333 323 67
88.43 0.9428 15900 2945
76.43 0.9130 1080 637
76.23 0.9128 893 651
77.76 0.9166 1020 1074
72.22 0.9037 147 131
71.98 0.9074 227 245
333 350 360 419 502 525 572 0.28
331 348 357 413 495 519 567 0.29
316 336 346 390 452 470 515 0.45
323 339 346 388 449 469 516 0.27
323 339 346 388 449 468 515 0.39
321 337 345 387 449 469 519 0.23
320 337 345 386 447 468 523 0.28
322 338 345 386 446 466 513 0.27
feed
Cetane index determined by eq 19. *Room temperature assumed to be 25 "C.
0
e Not
analyzed.
Table 111. Correlations To Predict TLP Yields (Equation 1 ) r2 Y, a b C coker HGO 98.28 1.0506 0.2517 0.0414 4.0163 hydrocracker HGO 93.95 1.0371 0.1133 0.0206 -0.0134
centage of A, B, and C in the feed. Assuming plug flow, the solutions to eq 2 and 3 are CA = C, exp[-(kl + kJ/LHSV] (5)
into LGO (B) and naphtha (C), and LGO (B) converts further into naphtha (C) as shown below:
CB = CBo exp(-k,/LHSV)
k
A-B-C
k,
CC =
If we assume that each conversion obeys first-order kinetics, then the rate equations are dCA/dt = -(ki 4- k,)cA (2) dCB/dt = k,CA - ksCB (3)
+ CB + CC
(6)
where LHSV = l/t. From mass balance eq 4,
w
and the mass balance is CAo + CBo + CC, = CA
kl X k3 - (hi + k,) Cb,(exp[-(kl + k,)/LHSV] - exp[-k,/LHSV]j
+
(4)
where k l , k2, and k3 are rate constants for each step; CA, CB, and Cc represent the percentage of A, B, and C at t ; t is the residence time; and C, CBo,and CC, are the per-
(c, + c B o + CC,) - (c, + CB)
(7)
In eq 5 and 6, LHSV (liquid hourly space velocity) is defined as the volume of feed/h/volume of catalyst. In the petroleum industry, the feed rate in commercial units is customarily expressed as LHSV, and the rate to a pilot unit is selected to fall within the same LHSV range. Flash calculations were conducted based on the Peng-Robinson equation of state, using a commercial computer program and varying temperatures (360,380, and 400 "C) at constant pressure (8.8 MPa) and H2/feed ratio (600 S m3/m3). Table IV summarizes the results. It is observed that, at the same operating conditions, hydrocracker HGO va-
Ind. Eng. Chem. Res., Vol. 28, No. 9, 1989 1281 Table IV. Results of Flash Calculation'
a t 15 "C
hydrogen feed total vapor liquid total vapor liquid total vapor liquid total
a t 360 "C
a t 380 "C
a t 400 "C
m3/h 60 000 100 60 100 1627 119 1756 1685 117 1802 1747 114 1861
coker vol % 99.8 0.2 100.0 93.2 6.8 100.0 93.5 6.5 100.0 93.9 6.1 100.0
HGO wt %
5.1 94.9 100.0 14.0 86.0 100.0 17.3 82.7 100.0 21.6 78.4 100.0
wt % on feed
100.0
90.6
87.1
82.6
hydrocracker HGO vol % w t 70 wt % on feed 99.8 5.4 0.2 94.6 100.0 100.0 100.0 93.3 17.2 6.7 82.8 87.6 100.0 100.0 93.7 22.1 6.3 77.9 82.3 100.0 100.0 94.2 28.8 5.8 71.2 75.3 100.0 100.0
m3/h 60 000 100 60 100 1627 118 1745 1689 114 1803 1756 108 1864
"Basis: Peng-Robinson thermodynamic method; feed rate = 100 m3/h; H2/oil = 600 S m3/m3; pressure = 0.1 MPa for feeds and 8.8 MPa for products.
porizes more readily than coker HGO and that the percent vapor from the liquid feed increases as the temperature increases. Accordingly, l/LHSV is not exactly the same as the residence time under the conditions tested; however, for simplicity, in applying the results in industry, LHSV was used in the present study. In the previous studies on the kinetics of HDS, HDN, and MHC (Yui, 1989) and aromatics hydrogenation (Yui and Sanford, 1989), we observed that plug flow was not attained in the pilot-scale trickle-bed catalytic reactor even with catalyst dilution and proposed that the plug-flow model used for reaction kinetics should be modified by including a power term for space velocity to compensate for nonplug flow. We also proposed that, as with LHSV, the effect of hydrogen partial pressure ( p H ) may be described by use of a power term. As a resuft, eq 5 and 6 may be modified to, respectively,
Table V. Kinetic Parameters coker HGO
hydrocracker HGO
kl k3
k0 EIR, K ko Mild Hydrocracking (a = 0.5, p = 0.4) a. Parallel Scheme (k3 = 0 ) 8932 4.274 X lo4 8.754 X lo4 3.775 X lo3 8.544 X lo3 7558 14987 6.847 X lo8 1.780 X lo8 b. Consecutive Scheme ( k , = 0 ) 8.754 X lo4 8932 4.274 X lo4 6.206 X lo7 13566 2.711 X lo5
kS
9.375
kN
H ydrodenitrogenation 1.524 X lo6 11747 1.485 X lo5 9108 (a = 1.0, p = 1.8) (a = 1.0, p = 1.3)
kl kl k2
+ k2
E/R, K
8674 7208 16189 8674 10295
Hydrodesulfurization 15 756 8.283 X 1O'O 15 848 (a = 1.0, p = 1.1) (a = 1.0, fl = 0.6) X
lo9
If we assume that the conversion occurs only consecutively (Le., k2 = 0), eq 8 and 9 can be simplified to k3
- (ki
+ k2)
C,(exp[-(k1
+ ~z)PH,B/LHSV"Iexp[-k#~t/LHSv"]) (9)
Rate constants kl, kz, and k3 may be expressed by the Arrhenius equations
CBo
(16)
(10)
Each conversion scheme was evaluated with pilot data.
k2 = k, exp(-E2/RT)
(11)
k3 = k3, exp(-E3/RT)
(12)
Analysis of the MHC Data For the parallel-consecutive conversion scheme, eq 8-12, the number of unknown parameters is eight: six kinetic parameters (k,,, k , k,,, El, E2, and E3)and two power terms (a and 0). +he model is complicated, and a good guess of the initial values is essential to obtain solutions to this nonlinear problem. The parallel scheme, on the other hand, can be solved by linear regression analysis. The consecutive scheme, eq 15 and 16, may require a combination of the linear and nonlinear regression analyses. Therefore, we analyzed the MHC data based on the relatively simple parallel scheme first, and obtained power terms and kinetic parameters. The values from the parallel scheme were then used for solving the other two schemes. Parallel Scheme. Equation 8 was linearized by putting k = kl + k2 and taking logarithms. A multiple linear regression technique was applied to obtain power terms for LHSV ( a ) and p H p (@),and kinetic parameters (ko and E ) for the combined rate constant k. The values of kl,, k,, El and E2were obtained from the relationship of kl and k2, eq 14. The results are shown as part of Table V.
where kl,,k,, and k3, are frequency factors; El, E2, and E3are activation energies (kJ/mol); R is the gas constant (8.314 kJ/mol/K); and T i s absolute temperature (K).It should be noted that units of rate constants include power terms for LHSV and pH2.Hence, the dimensions for k and k, are expressed as h-" MPa-B. If we assume that the conversion occurs only in parallel (i.e., k3 = 0), eq 8 can be used for C A and eq 9 for CB can be simplified to CB =
exp(-k3pH:/LHSV")]
kl = kl, exp(-El/RT)
kl +CA,,{~ - exp[-(kl + k z ) P ~ ~ / L H s v " l l k l + k2
(13) and the following relationship is maintained C B - CB,
cc - cc,
k,
=-
k2
(14)
1282 Ind. Eng. Chem. Res., Vol. 28, No. 9, 1989 80
410 CATALYSTS K AND S
I-
390
380
370
360
350
3409:
I
: 0
a
I
I
02
53 41
400
04
@@
m'
COKER HGO c A o = 84 I % C B ~ =15 9 % ct0= 0%
70
01501 008
-
-
006 -
0
004 -
S 1
t 2
; i v, 2
40
c
I
/
0 0 2 -
0015W
5
O O l -
0008 PHTHA I0
0006
0 PARALLEL SCHEME 0
CONSECUTIVE SCHEME
ooo4
ItCl
0 0
IC
20
30
40
50
60
OBSERVED YIELDS
70
1
80
(%)
Figure 2. Observed yields (simulated distillation) versus predicted yields using kinetic parameters in Table V.
Power terms for LHSV and p H pwere found to be 0.5 and 0.4, respectively, for both coker and hydrocracking HGOs. Consecutive Scheme. Using the power terms so obtained, we then analyzed the data for the consecutive scheme. It should be noted that k , in eq 15 for the consecutive scheme is identical with k , k , (or k ) in eq 8 for the parallel scheme. Since the kinetic parameters of kl in eq 16 are known, k , and E, for k , could be easily obtained by a nonlinear regression analysis. The results are shown in Table V. Parallel-Consecutive Scheme. Three methods were attempted to solve this problem. Method 1. The value of k (or k l + k,) for the parallel scheme is the same as k l + k , in eq 8 and 9 for the parallel-consecutive scheme. By use of the values of k , and E for k and a and 0,four parameters (k,,, k30,E,, and E,) were obtained by nonlinear regression analysis. By this method, we obtained negative values for k , both for coker and hydrocracker HGOs. Values for E 3 / R were 1.2 X lo5 K for hydrocracker HGO and 9.6 X lo6 K for coker HGO, both of which were large enough to lead to a value of zero for the rate constant k,. Method 2. In eq 9, we lumped k , / [ k , - ( k , + k,)] as one unknown constant k123, obtained three parameters k123, k,, and E, by nonlinear regression analysis, and then calculated k , (kloand E,). By this method, we obtained negative values for klZ3and k , for both coker and hydrocracker HGOs. Method 3. From the above two methods, we could not obtain satisfactory solutions. Nevertheless, we attempted to obtain six kinetic parameters a t once, employing the values from analysis of the parallel and consecutive schemes as the initial values for the nonlinear regression of eq 9. By this method, we obtained a negative value of k , for hydrocracking HGO and a negative value of k , for coRer HGO. Although the parallel-consecutive scheme would appear to be the most plausible conversion mechanism, none of the three methods in this scheme work for the present data. Chanda and Mukherjee (1987) observed a similar phenomenon when analyzing experimental data on oxidation of 2-propanol. The authors studied the kinetics with a parallel-consecutive scheme, where carbon dioxide is formed both by direct oxidation of 2-propanol and by oxidation of acetone formed in an intermediate step from
+
-
0 002
0 COKER HGO/CATALYST S
HYDROCRACKER HGO/CATALYST S
k
146
148
I50
I52
154
I56
158
160
162
164
IOOO/T ( K - ' )
Figure 3. Arrhenius plots of MHC (first order) of coker and hydrocracker heavy gas oils.
2-propanol, and obtained a negative rate constant for the oxidation of acetone. They concluded that the further oxidation of acetone was not a significant mechanism under the conditions employed and that the oxidation of 2-propanol to products may be assumed to proceed via parallel routes. Figure 2 compares the observed and calculated yields with the resulting kinetic parameters for parallel and consecutive schemes. Good agreement is observed for both schemes. Figure 3 shows Arrhenius plots for the parallel scheme. The results indicate that coker HGO cracks more easily than hydrocracker HGO, that both coker and hydrocracker HGOs have similar activation energies, and that activation energy for naphtha production is much higher than that for LGO production. In other words, cracking to naphtha (and presumably to gases as well) increases strongly as the temperature increases.
Kinetics and HDS and HDN Our experience shows that all gas oils (LGO, HGO, and combined LGO and HGO) from Syncrude bitumen obey 1.5th-order kinetics for HDS and first-order kinetics for HDN. The sulfur and nitrogen removal data were analyzed, employing modified rate equations that include power terms for LHSV and p H 2(Yui, 1989)
where S and N are sulfur (wt % ) and nitrogen contents (wppm), k , and k N are rate constants, and subscripts f and p represent feed and product, respectively. The results are summarized in Table V. The power term for LHSV was unity for both HDS and HDN but was 0.5 for MHC. The power term for pH2ranged from 0.6 to 1.8 for HDS and HDN but was 0.4 for MHC. Arrhenius plots are shown in Figures 4 and 5. As between coker HGO and hydrocracker HGO, hydrocracker HGO is more amenable to HDS and HDN.
Ind. Eng. Chem. Res., Vol. 28, No. 9, 1989 1283
'1
HDS t
*,*'
0 9 -
$
\ *
3
c
a
0
VYDROCRACKER HGO
0
B L E N D E D HGO
0-
3
=
W
2-
15-
-& +Z
E 5 w
$
I -
08-
06
-
04-
---
03:
o:i,:i":::I"'"";', 1 02-
LL
LGO CUI
HYDROCRACKER/CATALYST
,
S
I
004
148
146
I50
I52
154
IOOO/T
156
158
I60
162
164
(K-')
Figure 4. Arrhenius plots of HDS (1.5th order) of coker and hydrocracker heavy gas oils (value of p, see Table V).
-
? -
-IC
02015-
-
01
5
008:
$
006:
cn
z
8
5000
6000
7000 8 )O
Figure 6. Sulfur and nitrogen contents in total liquid products versus the corresponding values in LGO (195/343 "C) and HGO (343+ "C) cuts.
pressly for synthetic distillates from northern Alberta bitumen. Calculated CNs for the LGO cut from hydrotreated products were found to be about 30 for coker HGO and 31 for hydrocracker HGO. These values are well below the diesel fuel specification (CN = 40). Although we are interested in the sulfur and nitrogen contents of individual cuts, the plant operation is controlled by the quality of TLP. Therefore, it is important to know the relationship between quality of cuts and TLP. By analyzing the sulfur and nitrogen contents of LGO and HGO cuts, the TLPs from hydrotreated HGOs of coker, hydrocracker, and blended feed, the following correlations were obtained YHGO = 1.211yTLp for HGO cut (20)
i
0004 0006I46
3000 4000
for LGO Cut
(21)
where YHGo, YLGO, and YTLP are the sulfur or nitrogen content in HGO, LGO, and TLP, respectively. Figure 6 illustrates the goodness of fit of these equations. A
0008
2000
YLGO = 0.253yTLp
oo2[ 0015
IO00
S OR N CONTENT IN TLP (ppm)
ow003-
k!
0
;o\,
OC6
f
6/
015 -
0 0
COKER H G O / C A T A L Y S T K HGO/CATALYST S H Y D R O C R A C K E R /CATALYST S
COKER
148
150
152
154
IOOO/T
156
158
160
16'2
164
(K-')
Figure 5. Arrhenius plots of HDN (first order) of coker and hydrocracker heavy gas oils (value of @,see Table V).
Product Quality of LGO and HGO Cuts As for HGO hydrotreating, we were interested in the product quality of individual cuts such as naphtha, LGO, and HGO rather than the quality of TLP. We were particularly interested in the cetane number (CN) of the LGO cut and the nitrogen content of the HGO cut. CNs for LGO cuts were estimated with a newly developed correlation (Yui and Sanford, 1988) CN = -6979.40 + 7040.55DE2 - 13997.1DE(ln DE) 7.91843AP(ln DE) + 0.00771926AP(MP) 0.546587AP(ln MP) - 0.000241340MP2 (19) where DE is the density at 20 "C (g/cm3),AP is the aniline point ("C), and M P is the mid-boiling point ("C) determined by ASTM D2887. Equation 19 was developed ex-
Conclusions Mild hydrocracking of HGO into LGO and naphtha could not be described by a parallel-consecutive scheme for the data sets from this study. It could be described reasonably well by either a first-order parallel or consecutive model. The rate of conversion for coker HGO is higher than that for hydrocracker HGO. Both coker and hydrocracker HGOs have similar activation energies. The activation energy for naphtha production is higher than that for LGO production. In other words, the naphtha/LGO ratio increases as the temperature increases. The product yield can be estimated by power forms of reactor temperature, hydrogen partial pressure, and LHSV. HDS and HDN of HGO obey 15th- and first-order kinetics, respectively. The rate constants for HDS and HDN of hydrocracker HGO are higher than those of coker HGO. Power terms for LHSV are unity for HDS and HDN and 0.5 for MHC for this study. Calculated cetane numbers of the LGO cut from hydrotreated products are about 30 for coker HGO and 31 for hydrocracker HGO.
1284 Ind. Eng. Chem. Res., Vol. 28, No. 9, 1989
A good relationship between sulfur and nitrogen contents of LGO and HGO cuts versus those of TLP was obtained.
Acknowledgment Pilot-plant operation was performed by Betty Andrichuk, Sherry Court, A1 Maskwa, and Murray Noble. The authors thank D. Heaton for the flash calculations, R. Ashworth, J. Cooley, R. Kirchen, and V. Nowlan for their valuable discussions, and Syncrude Canada Ltd. for permission to publish this paper.
Abbreviations
AP = aniline point, "C CN = cetane number DE = density at 20 "C,g/cm3 HDN = hydrodenitrogenation HDS = hydrodesulfurization HGO = heavy gas oil LGO = light gas oil LHSV = liquid hourly space velocity, volume of feed/h/ volume of catalyst MHC = mild hydrocracking MP = mid-boiling point determined by ASTM D2887, "C TLP = total liquid product
Literature Cited Nomenclature a , b, c = power terms for temperature, hydrogen partial pressure, and LHSV, respectively (eq 1) CA,CB, CC = 90 A (HGO), % B (LGO), and % C (naphtha) in product C,, CBo,Cco = 90 A (HGO), % B (LGO), and % C (naphtha) in feed E,, El, E3 = activation energies (eq 10-12), kJ/mol k , , k,, k3 = reaction rate constants for step 1 (HGO to LGO), step 2 (HGO to naphtha), and step 3 (LGO to naphtha), respectively hi,, kZo,k3, = frequency factors (eq 10-12) k 1 2 3 = defined as k l / [ k 3- ( k , + k2)l (eq 9) k N , k s = reaction rate constants for HDN and HDS Nf, N , = nitrogen contents of feed and product, wppm pH2= hydrogen partial pressure, MPa R = gas constant, 8.324 kJ/mol/K Sf,S , = sulfur contents in feed and product, wt YC t = temperature, O C t = residence time, h T = absolute temperature, K JJHGO, JJLGO,yTLp = sulfur or nitrogen content in HGO, LGO, and TLP (eq 20 and 21), wppm Y = TLP yield (eq l),m3/m3of feed Yo = constant (eq I), m3/m3of feed Greek Symbols a = power term for LHSV (eq 8 and following) p = power term for hydrogen partial pressure (eq 8 and fol-
lowing)
Chanda, M.; Mukherjee, A. K. Kinetics of Vapor-Phase Oxidation of 2-Propanol to Acetone over a Copper Catalyst. Ind. Eng. Chem. Res. 1987, 26(12), 2429-2437. Desai, P. H.; Asim, M. Y.; van Houtert, F. W.; Nat, P. J. Mild Hydrocracking of FCC Feeds Produces Yield Benefits in Mid-Distillates, Gasoline. Oil Gas J. 1985 (July 22), 106-117. Elkes, G. J.; Page, T. H.; Thomas, M. E. Mild Hydrocracking-A Flexible Option to Produce High Quality Middle Distillate. Prepr.-AIChE (Spring Meeting, Houston, TX) 1987, paper 62a. Gembicki, V. A.; Andermann, R. E.; Tajbl, D. G. Mild Hydrocracking Fills Processing Gap. Oil Gas J . 1983 (Feb 21), 116-128. Kalnes, T. N.; Lamb, P. R.; Tajble, D. G.; Pegg, D. R. Mild Hydrocracking: A Low Cost Route for Middle Distillate. Prep.-Natl. Pet. Refin. Assoc. (Annual Meeting, San Antonio, TX) 1984, paper AM-84-36. Plantenga, F. L.; Sonnemans, J. W. M. Hydroconversion of Vacuum Gas Oil and Atmospheric Residua. Prepr.-Am. Chem. Sot., Diu. Pet. Chem. 1983, 28(3), 621-632. Wilson, M. F.; Simmons, R. A,; Notzl, H. Hydrocracking of Gas Oil from Athabasca Syncrude. Prepr.-Am. Chem. Sot., Diu. Pet. Chem. 1987, 32(2), 383-390. Yui, S. M. Hydrotreating of Bitumen-Derived Coker Gas Oil: Kinetics of Hydrodesulfurization, Hydrodenitrogenation, and Mild Hydrocracking, and Correlations to Predict Product Yields and Properties. AOSTRA J . Res. 1989, 5(3), in press. Yui, S. M.; Sanford, E. C. Diesel and Jet Fuel Production from Athabasca Bitumen, and Cetane Number Correlation. Prepr.4th UNITARjUNDP AOSTRA Int. Conf. on Heavy Crude and Tar Sands (Edmonton) 1988, paper 113. Yui, S. M.; Sanford, E. C. Kinetics of Aromatics Hydrogenation of Gas Oils. Submitted for publication in Can. J . Chem. Eng. 1989.
Received f o r review October 28, 1988 Accepted June 9, 1989
