Ind. Eng. Chem. Res. 1987, 26, 2337-2344
2337
The Paragon Process: A New Hydrocracking Concept Dennis J. O'Rear Chevron Research Company, Richmond, California 94802
The Paragon Process couples a conventional hydrocracker with a low-pressure reactor which contains a shape-selective zeolite. This combination permits each class of molecules in a vacuum gas oil (VGO) feed to be converted to the most desirable products: S and N heteroatoms H2S and salable NH,, aromatics and naphthenes gasoline-boiling naphthenes, n-paraffins and slightly branched paraffms C3-C8 olefins, and highly branched isoparaffins gasoline-boiling isoparaffins. When compared to conventional hydrocracking, Paragon increases the research octane number of the gasoline; improves the manufacture of jet and diesel fuels from waxy stocks; improves the quality of the reformer feed and reduces the load on this plant; and provides a new source of olefins which are heteroatom-free. rich in isoolefins. and highly concentrated. Experimental results are presented with a process analysis.
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Mobil's shape-selective zeolite ZSM-5 has been used to improve the selectivity of several commercial refining processes and has also led to the development of new ones. (Chen and Garwood, 1986; Heinemann, 1981; Flanigen, 1980). This paper describes the application of shape-selective zeolites to hydrocracking; the modified process developed a t Chevron is called the Paragon Process. The objective of hydrocracking is to convert nonsalable vacuum gas oil (VGO) into transportation fuels (jet fuel, diesel fuel, and naphtha for reforming), while minimizing the formation of light saturate gases and the use of expensive hydrogen. The Paragon Process converts the least valuable hydrocarbon components in a VGO to a new source of olefins, and it provides the opportunity to improve the manufacture of transportation fuels. Conventional hydrocrackers often have two stages operated at high pressure. The first removes almost all the heteroatoms from the VGO and does a little cracking. The second operates with extinction recycle and cracks the heteroatom-free VGO to naphthas and transportation fuels. (An intermediate stripper, operated at near atmospheric pressure, removes the H,S and NH, formed in the first stage.) However, not all the hydrocarbon classes in a VGO react with equal selectivity. Paraffins in the VGO tend to form light saturate gases and high yields of low boiling naphtha (typically C,-160 OF), which has a low octane and cannot be reformed (Figure 1). In addition, paraffins also are converted to the reformable naphtha where they contribute to difficulties in reforming because of their relatively poor selectivity for aromatic formation when compared t o naphthenes. n-Paraffins in the VGO are a particular problem if they are present as overlap (material in the VGO with the same low boiling range as the product). These light n-paraffins pass through the first stage, are distilled, and end up in the product unchanged. For jet- and diesel-fuel-mode hydrocrackers, these nparaffins cause freeze and cloud point problems and force the hydrocracker to recycle and crack the heavy portion of the product. This decreases the yield. Since shape-selective zeolites are used in commercial dewaxing processes to crack paraffins, they should be able to selectively act on paraffins in the hydrocracking process and improve its operation. The marriage of these two processes (dewaxing and hydrocracking),subject to general refinery needs and economics, leads to the Paragon Process. Figure 2 shows a conventional two-stage hydrocracker and how it can be modified to give the Paragon Process (hatched lines). Experimental Section A pilot plant simulating the Paragon Process was operated for many thousands of hours (O'Rear and Mayer,
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Table I. Properties of the Denitrified Feedstocks crude source Arab light Minas Arab light hydrocracker mode gasoline jet fuel diesel gravity, OAPI 41.4 36.9 36.0 aniline pt, O F 173.2 210.8 193.4 10 5.6 s, PPm 0.19 0.42 0.46 N, PPm R.1." at T, "C 1.4555 at 20 1.4420 a t 70 1.4480 at 70 0.8137 a t 20 0.7982 a t 70 0.8077 a t 70 density a t T, "C MW 243 306 293 pour pt, O F 40 95 70 n-oaraffins. wt % 12.75 36.05 6.56 group types by MS, LV % paraffins 47.7 51.8 26.9 naphthenes 38.0 42.5 52.6 aromatics 14.3 5.7 20.5 ndM anal., wt % 78.9 74.2 68.6 CP 13.8 23.7 25.1 Cn Ca 7.3 2.1 6.3 D1160 by LV % St/5 3031361 3991522 2211316 10130 4141555 5621651 3671592 50 653 712 750 70190 720 7591816 8271887 95/EP 8451867 9061972 % recovery a3 98 "R.I. = refractive index.
Table 11. Operating Conditions hvdrocracker mode gasoline jet fuel diesel fuel Hydrocracker 1100 1470 2200 1.0 1.4 2.3 5200 5400 6000 ICR 204 ICR 204 ICR 106
pressure, psig LHSV recycle gas, SCFB catalyst pressure, psig LHSV recycle gas, SCFB
Paragon Reactor 0 0 6.3 6.3 0 0
per-pass conv., LV % recycle cut pt, O F
Overall 80 380
65 470-550
0 12.7 0
65 700
1983). For convenience, the first stage was run separately on three feedstocks, and the denitrogenated products were collected (Table I). This first stage operated at commercial conditions with Chevron Catalysts ICR 106 and ICR 117. The second stage and Paragon reactor were operated at the conditions shown in Table 11. The conditions varied depending on the desired product slate
O888-SS8~/87/2626-233~~O~.SO/O 0 1987 American Chemical Society
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Figure 1. Relationship between feedstock paraffm content and the yields of gas and light nonreformable naphtha in conventional hydrocracking. Paragon process
Saturate Gas
Second-Stage Reactor
Naphthas Jet or Diesel
Figure 2. Flow diagram for a conventional hydrocracker and the Paragon Process.
(diesel-fuel mode versus gasoline- or jet-fuel modes). The pilot plant simulating the Paragon Process followed the flow scheme shown in Figure 2. The denitrogenated feedstocks were continuouslymixed with recycle liquid and then fed to the Paragon reactor. The effluent from the Paragon reactor was sent directly to a low-pressure separator and a series of on-line distillation columns. These separated the product olefins (gases and C,-300 O F liquid) from the unreacted 300+ O F liquid. The 300+ OF liquid was compressed, mixed with recycle hydrogen, and fed to the hydrocracker. The effluent from the hydrocracker was sent to a high-pressure separator where hydrogen was recovered and recycled to this reactor. As hydrogen was consumed in the hydrocracker, make-up hydrogen was added to maintain the pressure. The liquid from the separator was sent to another series of low-pressure separators and on-line distillation coluhns. These separated product gases and liquids from the unreacted recycle liquid. The latter was mixed with fresh denitrogenated feed to complete liquid recycle loop. The temperature of the second-stage reactor was adjusted to maintain a constant conversion of fresh denitrogenated feed per unit of time.
For convenience, the gas and liquid products from the hydrocracker and Paragon reactors were often combined. Occasionally, separate samples were taken from the individual reactors. The Paragon reactor and second-stage hydrocracker were operated downflow. The combined liquid products were batch-distilled to give a light naphtha (C5-180 OF), a heavy naphtha (180-400 OF), a jet fuel (400-550 O F ) , and a diesel fuel (300-700 O F ) . These product boiling ranges are typical values, and the exact boiling ranges are shown in the tables. This flow scheme was varied slightly in some of the preliminary experiments described below. The catalyst used for the Paragon reactor was 65% H-ZSM-5 (67/1 molar ratio Si02/A120,) bound with 35% Catapal alumina. The ZSM-5 was made in our laboratories based on Mobil's patents (Argauer and Landolt, 1972). The steps in this preparation included hydrothermal synthesis, filtration, washing, calcination to remove the template, ion exchange to remove the sodium, binding with alumina, extrusion, and a final calcination. The ZSM-5 crystals were roughly 1-10 pm, mostly free of twining, and coffin-shaped. ZSM-5 was chosen because of its commercial availability, ease of synthesis, and demonstrated use in dewaxing. The hydrocracking catalysts used in the second stage (ICR 204 or ICR 106) were commercial Chevron catalysts. The ZSM-5 extrudate was crushed to 8-16 mesh and diluted with an equal amount of Alundum to minimize bypassing in this high LHSV reactor. The hydrocracking catalyst was tested as whole extrudate. After they were loaded in the reactors, both catalysts were dried in flowing nitrogen at 800 OF. The ZSM-5 catalyst was not treated further. The hydrocracking catalyst was reduced and sulfided by commercial techniques.
Results Preliminary Experiments. The shape-selectivezeolite could conceivably be put in the first or second stages, or it could be used in a separate reactor either before or after the first stage. Also, it could be operated at a variety of conditions. These preliminary experiments defined the best setup and led to the flow scheme shown in Figure 2. Figure 3 shows that ZSM-5 is sensitive to nitrogen in the VGO-the 445 ppm nitrogen in a raw VGO reduces the catalyst's activity by 150 OF when compared to a denitrogenated VGO. When the ZSM-5 was incorporated into the first stage directly, it also fouled rapidly due to nitrogen poisoning. So by placing the shape-selective zeolite after the first stage, it lasts longer and can be operated at a higher LHSV. This figure demonstrates the remarkable resistance of ZSM-5 zeolite to coke fouling. When nitrogen is eliminated as a poison, catalysts in this service foul
Ind. Eng. Chem. Res., Vol. 26, No. 11, 1987 2339 585
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primarily by coke deposition. ZSM-5 zeolite is very resistant to coke fouling and can convert a reasonable fraction of denitrogenated feeds for many thousands of hours with little fouling. In comparison, faujasite and amorphous catalysts would foul completely in a few minutes at these conditions. Walsh and Rollman (1979) found that coke yields from cracking n-hexane over ZSM-5 were 2 orders of magnitude less than over mordenite. They attributed the coke resistance of ZSM-5 to spacial constraints in its pore system. Another benefit of processing a denitrogenated VGO is that its products are free of sulfur and nitrogen impurities. This simplifies their processing. Perhaps the most desirable product which can be derived from cracking a paraffin is an olefin. However, if the olefins produced by cracking are in the presence of active metals and hydrogen, they will be immediately saturated to less desirable paraffins. So to produce olefins, the shape-selective zeolite should not be a component of the catalysts in either the first or second stages. However, hydrogen transfer can also destroy olefins by reacting them with naphthenes to form paraffins and aromatics (Voge, 1958; Venuto and Habib, 1979). Figure 4 shows that the yield of olefins is sensitive to hydrocarbon pressure even when the catalyst does not contain metals and added hydrogen is not present. Under high hydrocarbon pressure and without metals, olefins apparently can be converted to saturates by hydrogen transfer from naphthenes which are abundant in the hydrotreated product from the first stage. These studies showed that the best location for the shapeselective zeolite is in a separate reactor (the Paragon reactor) which is located after the first stage. The best pressure is near atmospheric. At this pressure, ZSM-5's resistance to coking is important. This property, coupled with the use of a low nitrogen feed, enables the process to crack paraffins at near atmospheric pressure to form olefins and to do this for long periods of time without regeneration. A low nitrogen feedstock at low pressure is available in many two-stage hydrocrackers. The product from the first stage is stripped free of H2Sand NHBat low pressure before being hydrocracked in the second stage. (The Paragon reactor could be placed after the secondstage hydrocracker, but this is less desirable because most
Figure 5. Pilot plant run plot for conventional hydrocracking and the Paragon Process.
of the paraffins in the VGO would be cracked to low molecular weight paraffins in the hydrocracker and not cracked to olefins in the Paragon Reactor.) In the work described next, both the hydrocracker and the Paragon reactor were run for many thousands of hours. Neither catalyst lost more than a few degrees of activity, except for slight losses during start-up. Gasoline-Mode Hydrocracking. In this mode, the second stage of the hydrocracker is required to convert denitrogenated VGO into a 160-380 OF naphtha which is reformed into gasoline. Conventional hydrocracking also produces a low octane nonreformable naphtha (C5-160OF) which is not a desirable component in the whole gasoline pool. The objectives of conventional hydrocracking are to maximize the production of reformable naphtha, minimize the production of nonreformable naphtha and light saturate gases, and minimize hydrogen consumption. However, the Paragon Process changes the properties of these streams to such a large extent that these objectives must be reconsidered. When the Paragon reactor is turned on, it influences the hydrocracker because both are in the same recycle loop. The total amount of oil cracked per unit time is kept constant by adjusting the temperature of the hydrocracker (Figure 5). When the Paragon reactor is inactive (300 OF), the yields are those of a typical hydrocracker (Table 111). As the temperature of the Paragon reactor is raised, it cracks the paraffins in the oil, and this is shown by a reduction in the operating temperature of the hydrocracker and the hydrogen consumption. The paraffins cracked from the VGO in the Paragon reactor appear in the products as olefins. Looking at each product separately shows the following. C3Product. A large amount of propylene can be made (Figure 6). An analysis of the separate C3 stream from the Paragon reactor shows that it contains between 60 and 75 LV % propylene depending on the reactor temperature. C4Product. This stream from the Paragon reactor contains about 80% olefins, and the isobutylene content is very high, near 40% (Table IV, Figures 7 and 8). This
2340 Ind. Eng. Chem. Res., Vol. 26, No. 11, 1987
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Table 111. Yields from Gasoline Mode operation conventional hydrocracking cat. temp, O F hydrocracker Paragon reactor yields methane, w t % ethylene, wt % ethane, wt % propylene, LV % propane, LV % isobutane, LV % n-butane, LV % 1-butene, LV % isobutene, LV % trans-2-butene, LV % cis-2-butene, LV % C5-160 OF, LV % 160-380 OF, LV % c5+, LV % chem H2consump., SCFB C5-160 O F properties research octane motor octane Br no.
Paragon 16
581 300 (inactive)
557 700
558 750
557 800
0 0 0.06 0 2.92 11.54 3.16 0 0 0 0 20.78 79.61 100.39 964
0 0.11 0.04 6.80 4.26 8.87 3.22 1.16 5.74 2.52 1.66 23.25 63.25 86.50 665
0 0.27 0.08 9.34 4.42 7.19 3.15 1.25 5.44 2.54 1.70 22.85 62.68 85.53 628
0.07 0.64 0.16 13.53 5.08 7.48 3.06 1.48 5.98 2.67 1.83 19.35 60.82 80.17 600
83.1 79.2 0.1
-91 -81 70
90.6 80.0 63
87.1 80.4 47
Table IV. Comparison of the C4Streams from a Paragon Reactor at 700 OF and a FCC at 950 OF unit Paragon reactor FCC composition, LV % isobutane 7.3 34.8 14.3 n-butane 8.6 8.2 1-butene 13.6 40.7 isobutene 16.6 17.8 trans-2-butene 15.9 11.7 cis-2-butene 10.5 78.4 total olefins 56.6 impurities nil butadienes, ppm 1000 nil 100 s, PPm nil 50 N, PPm
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Figure 6. Yield of propylene (LV % ) versus the Paragon reactor temperature: ( 0 )gasoline mode, Arabian light; (a)jet-fuel mode, Minas; (e)diesel-fuel mode, Arabian light.
stream is free of sulfur and nitrogen impurities and has a very low diolefin content. So it is ideally suited for the manufactur of methyl tert-butyl ether (MTBE) or other chemicals like polybutene.
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Table V. Analysis of the C6-160 OF Product Paragon Reactor group types by MS, LV % paraffins naphthenes aromatics olefins physical properties research octane motor octane Br no. I no. detailed anal. of C i s isopentane n-pentane cyclopentane 3-methyl-1-butene 1-pentene 2-methyl-1-butene trans-2-pentene cis-2-pentene 2-methyl-2-butene cyclopentene total etherable isopentenes detailed anal. of C i s isohexanes n-hexane cyclohexane methylcyclopentane benzene 4-methyl-1-pentene 3-methyl-1-pentene 2,3-dimethyl-l-butene cis-2-methyl-2-pentene trans-4-methyl-2-pentene 2-methyl-1-pentene 1-hexene cis-3-hexene 2-methyl-2-pentene cis-3-methyl-2-pentene trans-2-hexene cis-2-hexene trans-3-methyl-2-pentene 2.3-dimethyl-2-butene total etherable isohexenes
from the
13.9 3.1 4.5 78.5 95.8 79.8 163
