2752
Ind. Eng. Chem. Res. 2009, 48, 2752–2769
REVIEWS Lowered Emissions Schemes for Upgrading Ultra Heavy Petroleum Feeds Edward Furimsky IMAF Group, 184 Marlborough AVenue, Ottawa, Ontario, Canada, K1N 8G4
Several commercial processes have been successfully used for upgrading heavy petroleum feeds containing less than 300 ppm of V + Ni. For problematic petroleum feeds, metal content exceeds this level and may approach 1000 ppm of V + Ni. Frequently, the upgrading battery is located in remote locations in the proximity of the heavy crude producing wells. In this case, the objective of upgrading is a syncrude suitable for pipelining. The complexity of upgrading increases with increasing content of metals, resins, and asphaltenes. The database of processing parameters has to be established for selecting optimal upgrading schemes, e.g., catalytic versus noncatalytic routes. From a processing point of view, the information on the yield of syncrude, steam, and electricity requirements, as well as hydrogen and catalyst consumption, are of primary interest. The parameters determining the environmental impact on the upgrading schemes have to be considered as well. The processes suitable for upgrading problematic feeds under evaluation include catalytic hydroprocessing, slurry bed hydrocracking, and coking. Various combinations of these processes have been used as well. Deasphalting combined with catalytic processes may be an attractive route providing that the rejected asphalt can be efficiently utilized on site. Lowered emissions can be achieved by utilizing asphalt and other residues from upgrading on gasification island for production of hydrogen, carbon dioxide, steam, and electricity. Compared with combustion, an integrated gasification-combined cycle operates at higher overall thermal efficiency with all emissions being significantly lower. Introduction Several types of processes are available for upgrading heavy petroleum feeds. They involve either hydrogen addition to the feed or carbon rejection from the feed. The former requires the presence of an active catalyst. For catalytic processes, it is more difficult to upgrade vacuum residues (VR) than atmospheric residues (AR), whereas few problems have been experienced with catalytic upgrading of vacuum gas oil (VGO), heavy gas oil (HGO), and deasphalted oil (DAO). The difficulty of upgrading increases with increasing content of metals and asphaltenes. The metals of primary importance include the porphyrin forms of vanadium (V) and nickel (Ni). There is little information suggesting that petroleum feeds containing more than 300 ppm of metals were upgraded via a catalytic route on a commercial scale, although some processes were designed to handle heavy feeds containing as much as 700 ppm of the metals.1-3 Biasca et al.4 published an extensive database on properties of petroleum feeds. This database was used by Rana et al.5 to establish ranges of properties of feeds suitable for residue fluid catalytic cracking (RFCC), hydroprocessing, and thermal processes. These correlations are shown in Figure 1. These results indicate that feeds with metal contents between 300 and 1000 ppm may still be upgraded via hydroprocessing although with moderate conversions. However, it is unlikely that even moderate conversions can be achieved without significant catalyst inventory and hydrogen consumption.6,7 The simplified flow-sheet of an upgrading complex is shown in Figure 2. The complex may be situated near the site of heavy oil production wells. Prior to upgrading heavy crude is subjected to dewatering/desalting. Subsequently, it is either topped to remove distillates or directly enters upgrading plant. The primary products from upgrading are hydroprocessed to attain specifications of the synthetic crude (syncrude) ready for pipelining. With primary upgrading being performed on site, the cost of diluent,
which otherwise would have to be used for preparation of the mixture with heavy crude suitable for transportation by pipeline, can be eliminated.8 For the purpose of this study, the residue from primary upgrading enters the integrated gasification combined cycle (IGCC) plant employing a Texaco gasifier which operates in a partial heat recovery mode.9 If necessary, the residue can be subjected to visbreaking to decrease its viscosity. This results in additional liquid products. The flexibility of the IGCC plant operation provides an opportunity for simultaneous production of H2 required for hydroprocessing, as well as CO2 and steam which may be utilized during the enhanced oil recovery (EOR), as well as the electricity necessary for operation of the complex.
Figure 1. Effect of properties of residue (343 °C+) on selection of upgrading technology. Reprinted with permission from ref 5. Copyright 2007 Elsevier.
10.1021/ie800865k CCC: $40.75 2009 American Chemical Society Published on Web 01/29/2009
Ind. Eng. Chem. Res., Vol. 48, No. 6, 2009 2753
2. Carbon Rejecting Processes
Figure 2. Simplified flowsheet of complex for production and upgrading problematic petroleum feeds. Table 1. Typical Emissions (lb/MW h) from Power Generating Technologies (petroleum coke; 100% capacity) emissions
natural gas comb cycle
gasification comb cycle
circulating fluid bed
PC boiler FGD and SCR
SO2 NOX CO VOC particulates CO2 solid waste
0.0 0.3 0.2 0.02 0.05 820 0
0.6 0.4 0.4 0.07 0.07 1930 9.1
3.7 0.9 1.5 0.08 0.2 2170 350
3.6 1.5 N/A N/A 0.2 2120 190
In some cases, the heavy oil producing wells are in very remote locations. Therefore, electricity, H2, steam, and/or CO2 must be generated on site. The environmental parameters in Table 1 confirm the advantages of gasification over combustion. Moreover, the slags produced during gasification are virtually nonleachable contrary to those from combustion. A high concentration CO2, a byproduct of H2 production from synthesis gas, could be utilized on site for EOR rather than disposed of to the atmosphere. In some situations, the synthesis gas produced by gasification may be used as the feed for the Fischer-Tropsch (FT) synthesis to produce blending components having high octane and cetane numbers. At the same time, the waxy byproduct of the FT synthesis may be hydroisomerized to a high quality lube-based oil and middle distillates. It should be noted, that these liquid products from FT synthesis contain neither sulfur nor nitrogen. This suggests that coke may displace natural gas as the source of H2, chemicals, and high quality blending components for transportation fuels. Then, the additional natural gas can become available for more noble applications. Ikbal et al.10 compared five different configurations of petroleum refinery to obtain the most efficient utilization of the residues from carbon rejecting processes employing the gasification route. The primary objective of this study is the comparison of processing schemes used for the upgrading of heavy feeds to syncrude. The focus is on the feeds containing more than 300 ppm of metals (V + Ni). For the purpose of this review, these feeds are referred to as problematic and/or ultra heavy feeds compared with extra heavy feeds the metals content of which may range from 200 to 300 ppm. The primary focus is on the operating parameters and on the environmental performance. It has been noted that a volume of relevant information can be found in the proceedings from international conferences organized under the auspices of the United Nation Organization. One of the objectives of this review is to make such information more readily available to other researchers in the field.
The carbon rejection from problematic feeds can be achieved either thermally or by separation of metals and heaviest components of the feed via deasphalting (DAS). In the former case, a sufficient viscosity ensuring a steady feeding to the reactor is an essential requirement. Such a level of the viscosity of the heavy feed can be attained either by preheating or by blending with lighter fractions. Apparently, carbon rejecting processes can handle any heavy petroleum feed regardless the content of metals and asphaltenes. However, other factors, e.g., yield of distillates, utilization of rejected carbon (residue), emissions, etc., have to be taken into consideration as well. It should be noted that information on carbon rejecting processes is rather extensive. This information can be readily accessed in various textbooks. Therefore, for the purpose of this review, only essential facts of the individual processes, with the aim to assist reader, are given. 2.1. Thermal Processes. During thermal treatment, portions of the resins and asphaltenes in the feed are converted to distillates and gas, whereas the other portion is rejected as the residue together with most of the metals. The primary products require a hydroprocessing step to attain specifications of the syncrude for pipelining. Thermal processes operate at temperatures approaching 800 K either in the absence of H2 (thermal cracking) or in the presence of H2 (hydrocracking), whereas the temperatures employed during hydroprocessing rarely exceed 700 K. For most carbon rejecting processes, the operation is conducted at a near atmospheric pressure. The heat required for thermal cracking is supplied by combustion of the gaseous byproduct and/or of the solid residue such as pitch and coke. Table 211 compares the yields of gases, distillates, and residues from conventional thermal upgrading processes which have reached commercial stage. For comparison, the last column in Table 2 shows the yields obtained in an ebullated bed reactor. The VR derived from an Arabian crude was used for this comparison. The extension of visbreaking through hydrovisbreaking to thermal hydrocracking (HCR) resulted in the significant increase in the yield of distillates. In this case, the thermal HCR was operated under elevated pressure of H2 compared with a near atmospheric pressure employed during hydrovisbreaking. The yield was further increased by DAS of the solid pitch from HCR as it is shown in the column HCR/ DAS in Table 2. With respect to the yield of distillates, a flexi coking (FC) process is more efficient than the delayed coking (DC) process. Moreover, the former has a provision for conversion of most of the coke to fuel gas which is used on the site for generation of electricity and steam. Only traces of asphaltenes and metals were present in the primary products obtained from thermal processes except for the HCR/DAS case shown in Table 2.11 For the latter, primary liquids also included DAO suggesting that their hydroprocessing requires more severe conditions and a greater amount of catalyst because asphaltenes and metals (∼50 ppm) may be still present. It should be noted that the yields of primary liquids can be increased by modifications of the conventional thermal processes. For example, the liquid yields almost doubled using a modified visbreaker, i.e., the high conversion soaker (HCS) cracking process.12,13 This resulted from the modified reactor internals and from using steam as the fluidizing medium at a near-atmospheric pressure. The decades of practical experience show that large scale plants upgrading extra heavy petroleum feeds usually employ coking technology. The technology has been dominated by DC and FFC processes. Gaseous byproduct formed during coking represents a valuable feedstock for the production of hydrogen,
2754 Ind. Eng. Chem. Res., Vol. 48, No. 6, 2009 Table 2. Yield of Products (wt %) from Thermal Processes1 (Feed VR from Heavy Arabian) coking yield
visbr
Hvisbr
HCR
DC
FFC
HCR/DAS
ebbul
gas naphtha gas oil total dist residue form of res
1.7 6.1 10.4 16.5 81.8 tar
2.5 3.0 34.5 37.5 60.0 liquid pitch
3.0 3.7 53.3 57.0 40.0 solid pitch
10.7 13.9 42.6 56.5 32.8 coke lumps
10.0 13.7 54.4 68.1 22.0 coke powd
3.0 3.7 67.3 71.0 26.0 solid pitch
4 13 66 79 17 vac res
synthesis gas, electricity, etc. Moreover, petroleum coke may no longer be a refinery waste. In fact, as it was indicated earlier, its value can be significantly enhanced when gasification is chosen as the utilization route. Besides DC and FFC processes, which have been used commercially, other configurations of coking processes are in various stages of development. The choice between the DC and FFC mode of the coking process depends on the scale of operation. Generally, the latter process is chosen for large scale upgrading plants compared with small and/or medium scale upgrading for which the DC process is preferred. Moreover, because of a longer residence time, the DC process is more sensitive to the properties of heavy feeds. For example, the increase in the yield of coke relative to liquids with increasing Conradson carbon residue (CCR) of the feed is more pronounced in the DC process than that in the FFC process.14 For ultra heavy feeds, this may be an important factor to be considered while deciding between the coking processes. 2.1.1. Delayed Coking Process. This process comprises two reactors/drums: one operating in the coking mode while the other is simultaneously decoked.15 This ensures the semicontinuous operation. For the operation, the coking drum is filled with VR exiting from the bottom of the vacuum distillation column.16 The heat required for coking is supplied by the internal heat source. The gaseous byproduct and primary liquids exit at the top and enter fractionator. The final boiling point of the primary liquids is determined by the temperature of coking (∼800 K). The heavy gas oil (HGO) fraction of the primary liquids, boiling above 650 K accounts for more than 50% of the product. The coke removed from the drum at the end of coking cycle is in the form of lumps. Therefore, it has to be crushed before its utilization via either combustion or gasification. The DC coke derived from conventional crude may be suitable for preparation of various carbon products. However, similar utilization options may be limited for DC coke obtained from problematic feeds because of the high content of metals and sulfur. 2.1.2. Fluid/Flexi Coking Process. The FFC process operates in a continuous mode.15,17,18 In the flexi-coking mode, the process consists of the coker where the conversion is achieved by injecting heavy feed in a liquid form into fluidized bed of the hot coke particles.19 This ensures a short contact time at coking temperature. Consequently, unwanted cracking of primary products is minimized. This is one of the advantages of the FFC process compared with the DC process. The coke particles with a thin layer (∼5) of the fresh deposit on the external surface are withdrawn from the coker and transferred to the heater where their temperature is increased by contacting with hot gasification products. Part of the coke from heater enters the gasifier where it is converted to fuel gas by reacting with steam and air. In the fluid coking mode, the heater is replaced by burner. Also, the gasifier is no longer part of the scheme as it is in the flexi coking process. During fluid coking, the coke particles withdrawn from the coker are combusted with air under the oxygen starving conditions. The aim of the partial combustion is to increase temperature of the coke particles (to about
Table 3. Comparison of the EUREKA Process with the DC Process (feed: spec gr of 1.030; CCR of 22.4 wt %; 538 + 89.0 wt %) products, wt %
EUREKA
DC process
yields C4-gas C5-370 °C 370 °C+
5.3 33.6 28.4
10.4 39.3 16.3
residue, wt % VR pitch coke
32.7 34.0
900 K). The hot coke particles are then returned to the coker. With this arrangement, all heat required for coking reactions is generated in the process. Most of the produced coke has particle size less than 1 mm.20 Then, little preparation is necessary in the case that either combustion or gasification options are chosen for utilization of coke. During coking, most if not all metals end up in coke. Consequently, the primary liquids contain little metal and asphaltene. The coking temperature ensures that most of the primary products consist of the naphtha and HGO fractions. Similarly as in the DC process, more than 50% of HGO from the FFC process boil above 623 K (350 °C). Always, the primary products from coking have to be stabilized via hydroprocessing to attain specifications of syncrude and/or those of commercial fuels. 2.2. EUREKA Process. An alternative to conventional coking processes for heavy feed upgrading is the commercially proven EUREKA process consisting of the semibatch reactor system.21 In the EUREKA process, a heavy feed (usually VR) is preheated and mixed with the recycle oil before being fed into another preheater. Subsequently, the mixture is injected into the reactor together with the superheated steam. The latter provides additional heat necessary for thermal cracking. It is unlikely, that at temperatures employed (∼850 K) the reaction of steam with the feed yielding H2 can occur to any great extent. Table 321 shows that the yield of distillates from the EUREKA process is higher than that from delayed coking but lower than that from the FFC process when compared with the results in Table 2.21 However, the residue from the former, i.e., EUREKA pitch, has much better combustion and/or gasification properties than delayed coke. It can be also used as a binder for the preparation of metallurgical coke from a lower quality coals. In spite of some positive parameters, the EUREKA process has not yet made an impact on the upgrading of problematic feeds on a commercial scale. In this regard, more demonstration time may be required. 2.3. Deasphalting (DAS). The physical separation of asphaltenes and metals from heavy petroleum feeds termed DAS is a well established commercial process. A brief account of this process in relation to the separation using distillation methods is given only. The former, conducted under rather mild conditions, is based on the difference in solubility of petroleum constituents (e.g., oil, resins, and asphaltenes) in various solvents, while distillation methods are based on the difference
Ind. Eng. Chem. Res., Vol. 48, No. 6, 2009 2755 Table 4. Deasphalting of Boscan and Zuata Crude butane AR DAO
asph
Table 5. Yields and Properties of AR and VR from Boscan and Zuata Crude
pentane VR
DAO
asph
AR DAO
asph
DAO
57 59 0.962 4.8 14 140
33 29 1.137 39 140 2300
31 32 0.986 6.3 13 140
35 31 1.113 38 240 3200
70 70 0.991 8.3 40 440
19 17 1.135 47 300 2700
40 40 1.004 9.7 35 360
26 23 1.128 44 290 4100
Zuata crude yielda wt % vol % spec gr CCR, wt % Ni, ppm V, ppm a
62 63 0.981 4.7 13 45
22 20 1.122 43 270 1260
21 21 0.993 7.6 14 40
30 27 1.109 38 265 1180
68 69 0.994 8.3 30 130
16 14 1.122 49 400 1950
34 34 1.016 13.9 45 175
16 14 1.163 50 370 1730
yield on crude, wt % vol % spec gravity CCR, wt % Ni, ppm V, ppm
100 100 0.999 14.7 107 1122
VR
89 87 1.019 16.5 120 1260
66 63 1.049 22.4 135 1700
84 83 1.016 15.4 93 448
51 48 1.059 28.5 162 706
Zuata crude yield on crude, wt % vol % spec gravity CCR, wt % Ni, ppm V, ppm
100 100 1.006 13.4 81 384
Table 6. Upgrading Orinoco VR (68% on Crude) VR
Yields on crude.
in boiling range of hydrocarbon fractions. Deasphalting involves the formation of two phases as the result of mixing a heavy feed with a solvent. One phase comprises a dissolved portion of the heavy feed, and the other consists of the solid or semisolid residue, i.e., predominantly asphaltenes and metals. To allow countercurrent flow of the two phases in the industrial reactors, the specific gravity of the solvent must be different than that of the feed. This would also ensure an easy separation of the solvent from the dissolved and precipitated portions of heavy feed. In any case, the solvent recovery may account for a relatively large part of the energy requirements of the DAS process. It was reported that from the energy demand point of view, the solvent recovery under supercritical conditions reduced utilities compared with the conventional method.22 In refinery practice, normal alkanes such as propane, butane, pentane, hexane, and heptane have been the solvents of choice. However, the use of a light paraffinic naphtha fraction has also been noted. The early DAS units used propane as the solvent and VGO and AR as the feeds to produce the base oil for lubricant preparation. Gradually, VR and heavy crudes have been included as the feeds and higher alkanes as the solvents. The ability of the alkane solvents to dissolve asphaltenes increases with the increase in their molecular weight.23,24 At the same time, the quality of DAO is decreasing. However, this is offset by higher yields of DAO. This is evident from results in Table 4, which compare pentane and butane as solvents used for DAS of Boscan and Zuata crudes. The properties of these crudes are shown in Table 5.25 The decreased quality (because of increased yield) requires more extensive hydroprocessing of DAO. Therefore, the optimization of the DAS operation involves trade-offs between the higher yield of the more contaminated DAO and the additional cost of hydroprocessing associated with it. Tables 4 and 5 show that, for Boscan crude, vacuum distillation yielded about 34 wt % distillates and 66 wt % VR. Further processing of the VR via DAS using pentane resulted in the additional 40 wt % of liquid products (DAO). However, as it is shown in Table 4,24 significant contamination of DAO with metals and CCR could not be avoided. In this regard, the quality of DAO after butane DAS approaches that of a medium heavy crude and as such may be suitable for pipelining, whereas the DAO from pentane DAS may require additional upgrading before pipelining. The overall yield of distillates may be further increased by integration of DAS with coking of asphalt. As Table 626 shows, in this case, not only the yield but also the
AR
Boscan crude
asph
Boscan crude yielda wt % vol % spec gr CCR, wt % Ni, ppm V, ppm
crude
VR
DC
DAS + DC
74 68 700 15-30 693-753 0.2-1 80-95 COa ∼0.002 160-230
a
CO-continuous operation. b RCC-relative catalyst consumption for the same feed for 1 year.
functionalities which are suitable for the secondary upgrading of DAO differ from those for the upgrading of primary liquids from slurry bed HCR. For the former feed, the catalyst has to exhibit adequate HDM activity. On the other hand, the selection of catalysts fro upgrading the primary liquids from slurry bed HCR has to take into consideration the presence of refractory S- and N-containing heterorings, as well as aromatics because of severe conditions employed. Low reactivity components can be present in liquids from coking processes as well. Therefore, to a certain extent, the requirements for secondary upgrading are determined during the primary upgrading step. These requirements are important for the overall comparison of the processes. 4. Catalytic Processes The metal content in heavy feeds has an over-riding influence on their suitability as the feeds for catalytic upgrading processes.30 While comparing several catalytic reactors, Kressmann et al.2 indicated that ∼700 ppm of metals in heavy feeds may be an upper limit which can be still acceptable for catalytic upgrading routes. However, this can only be achieved a significant cost of catalyst inventory. The operating parameters of these reactors are shown in Table 11.2 The approach used in the present study assumed ∼300 ppm of metals as the lowest limit. This exceeds the amount of metals in most of the feeds which have been included in the studies on catalytic upgrading of extra heavy petroleum feeds found in scientific literature. Therefore, the assessment of catalytic routes for upgrading such feeds could not be made without several assumptions. In this regard, it should be noted that the metal storage capacity (MSC), defined as the ratio of the weight metals which deposited from the feed per unit weight of fresh catalyst, was assumed to be 1.0, although this appears to be rather idealized case. Because of metal deposition, the activity of catalyst is gradually declining to the point that catalyst replacement is necessary.38 Therefore, for heavy feeds, MSC of catalyst is the parameter of primary importance. Figure 439 shows that MSC has a significant impact on the feasibility of catalytic upgrading. It is evident that at least five times more catalyst with MSC of 0.2 would be required for upgrading heavy feeds containing about 700 ppm of Ni + V compared with the catalyst possessing MSC of 1.0. This suggests that the success of catalytic upgrading of problematic feeds depends on the development of catalysts with MSC in the range of 1.0 or higher. For such feeds, the catalyst utilization may also be improved by a proper match of the feed and catalyst with the type of catalytic reactor.6 Three types of catalytic reactors have been used commercially, i.e., the reactors employing fixed bed (FB), moving bed (MB), and expanded and/or ebullated (EBU) bed.2,5,6
4.1. Fixed Bed (FB) Reactors. The best known multistage upgrading processes employing FB reactors include the Hyvahl40-45 and ARDS46,47 processes. The flowsheet of the multistage process shown in Figure 5 approaches configurations of the latter process used for upgrading AR containing less than 100 ppm of metals. The advanced version of the Hyvahl process employs two HDM reactors and two guard reactors upstream of two HDS reactors. The guard reactors operate in a perturbating mode, i.e., one in operation, while the other is in catalyst reloading mode. It was anticipated that with this arrangement upgrading of extra heavy feeds (e.g., ∼300 ppm of metals) could be achieved. The feeds which may be successfully upgraded using the Hyhvahl process include the AR derived from the Lloydminster and Rospomare crudes shown in Table 12.40 The yields of products from these runs are shown in Table 13. It is evident that the efficiency of the Hyvahl process for upgrading these feeds is rather high. Using the amount of the unconverted AR and its metal content, the estimate of catalyst consumption was made for a plant processing 40 000 b/d of the feeds. For this estimate, it was assumed that the MSC of the HDM catalyst was 1.0. This would translate into about 1.32 t/d of the catalyst consumed. It is evident from Table 12, that a significant reduction in sulfur and asphaltenes contents occurred besides that of metals. To accomplish this, the same amount of HDS catalyst was assumed in the HDS reactors of the Hyvahl process, giving total catalyst consumption of about 2.6 t/d. Overall, the amount of processed feed (barrel per kilogram of catalyst) approached 15 for Lloydminster crude and 10 for Rospomare crude. These results are included in Table 12 as well. Peries et al.43 pointed out that for a feed containing 100-120 ppm of metals, the Hyvahl process could achieve continuous operation lasting more than 1 year. For the MSC of HDM catalyst of 1.0 and 40 000 b/d operation, the amount of the catalyst consumed would approach 0.66 t/d. Assuming the same amount of HDS catalyst would result in about 30 b/kg of the upgraded feed. The HDM catalyst employed in the Hyvahl process is of the NiMo formulation supported on the macroporous alumina. The MSC of this catalyst may approach 1.0. It should be, however, noted that such a high MSC requires an even radial deposition of metals across the catalyst particle during the operation. This cannot be achieved without fine-tuning the operation. In this case, the attention has to be paid to the asphaltenes conversion with the aim to minimize coke lay-down on catalyst surface. Otherwise, such a high catalyst utilization could not be attained. Therefore, a high level of HDM can only be achieved by maintaining relatively low rate of HDAs. The HDM activity of this catalyst was tested using the most problematic feeds such as those shown in Table 14.43 The steady performance until an MSC of 1.0 was attained was demonstrated for a heavy feed derived from Boscan crude. The feeds in Table 1443 were used to make the estimate of catalyst consumption and the amount of feed processed per unit weight of catalyst. Assuming 80% metal removal for the plant processing 40 000 b/d of heavy feeds, the annual HDM catalyst consumption would approach 2500, 1100, and 800 t and the corresponding feed/catalyst ratio of 5.8, 13.3, and 18.2 b/d, for Boscan, Cerro Negro, and Laguna, respectively. These results were estimated assuming the MSC of the HDM catalyst to be 1.0. However, at 80% metals removal rate, the products would still contain about 270, 120, and 90 ppm of metals after processing of Boscan, Cerro Negro, and Laguna crude, respectively. With respect to further processing, the presence of the CCR precursors and asphaltenes in such products deserve
Ind. Eng. Chem. Res., Vol. 48, No. 6, 2009 2759
Figure 4. Effect of metals in feed and metal storage capacity (MSC) on cost of catalyst inventory. Table 12. Characteristics of Atmospheric Residues and Products from the Hyvahl Process (40 000 b/d) Lloydminster crude
Rospomare
Aa
B
A
B
72.2
42 59 0.951 0.70 7.7 2.7b 16 0.10
80.4
41 51 0.977 0.56 11.9 3.4b 8 0.003
b
yield , wt % on crude on feed spec gravity sulfur, wt % CCR, wt % asphaltenes C7, wt.% V + Ni, ppm total metals, t/d metals removed, t/d HDM cat consump, t/d total cat consump, t/d barrel/kg a
Figure 5. Simplified flowsheet of fixed bed multistage system for upgrading distillation residues.
attention as well. Therefore, the next stage processing of these products would require a catalyst possessing high MSC and high HCR activity.7 Further decrease in the content of metals in the Boscan, Cerro Negro, and Laguna products from 270, 120, and 90 ppm, respectively, to about 10 ppm would require additional 1.6, 0.6, and 0.5 t/d of the HDM catalyst, respectively. Assuming similar amount of catalyst in the HDS section of the Hyvahl process, would result in a total catalyst consumption of 10.1, 4.2, and 3.2 t/d giving processing ratios of 4.0, 9.5, and 12.5 b/kg for the overall upgrading of Boscan, Cerro Negro, and
1.002 4.30 15.0 11.2 223 1.42 1.32 1.32 2.6 15
1.063 7.75 26.4 23.9 328 2.09 2.09 2.09 4.2 10
A virgin AR; B AR after upgrading. b In parts per million.
Table 13. Yields of Products (wt %) feed AR
Lloydminster
Rospomare
H2S + NH3 C1-C4 gasoline gas oil VGO VR
4.3 3.3 5.4 30.0 43.8 15.0
8.2 4.4 4.8 34.1 36.5 14.5
Laguna crudes, respectively. It should be noted that these results represent a rather idealized case, i.e., the MSC of HDM catalyst 1.0. Otherwise, the HDM catalyst consumption would increase
2760 Ind. Eng. Chem. Res., Vol. 48, No. 6, 2009 Table 14. Properties of Heavy Feeds TBP cut point, °C spec gravity sulfur, wt % CCR, wt % asphaltenes, wt % C5 C7 nickel, ppm vanadium, ppm HDM cat cons t/d Aa Bb Total cat cons, t/d feed/HDM cat, b/kg A B feed total cat, b/kg b
Boscan
Cerro Negro
Laguna Once
200 1.003 5.42
250 1.014 3.95 16.7
200 0.995 2.81 17.6
19.9 13.3 109 1240
17.6 11.7 110 480
13.1 8.0 70 373
6.9 8.5 10.1
3.0 3.6 4.2
2.2 2.7 3.2
5.8 4.7 4.0
13.3 11.1 9.5
17.7 14.8 12.5
a A assumes an HDM rate of 80% for a plant processing 40 000 b/d. B assumes metal removal down to 10 ppm for the same plant as in A.
Table 15. Upgrading Venezuelan Crude Using Hyvahl and DAS Processes (50 000 b/d of feed)
synthetic fuel, wt % crude specific gravity sulfur, wt % nitrogen, ppm CCR, wt % C7 insol, wt % V + Ni, ppm metals in, t/d out, t/d
crude
HDM
HDM + HDS
HDM + DAS
HDM + HDS + DAS
100
96.4
96.3
86.0
90.2
0.995 2.84 4250 14.0 7.9 494 3.9
0.953 1.6 3950 9.8 3.0 110
0.921 0.5 3000 6.7 1.7 55
0.925 1.4 2700 3.1
