Methodology. The methodology used in the study included:
• The primary units (CSF, CDU, VDU, stabilizers and naphtha splitters)
were simulated, and the model was vali-dated with the current crude assay
• The crude assay was replaced with the new crude assay; see the true boiling point (TBP) distillation curve in FIG. 2 for the existing and new crude blends
• For diesel HT units, light gasoil (LGO) and heavy gasoil (HGO) streams were considered
• For HC fresh feed, light vacuum gas-oil (LVGO) and heavy VGO (HVGO) were considered
• For the naphtha hydrotreater (NHT), 85°C to 171°C ASTM D86 cutpoints were considered
• The crude charge rate for the pri-mary units was kept the same as in the base case.
Base case validation. The model is considered to be validated, as there was a close match between the plant data and the simulated data. See TABLES 1–7 for the validation results for the primary units.
The parameters that did not match with the plant data are shown in red. The deviations have been carried forward while predicting the flowrates for the new crude blend case. Deviations found in the crude column top section are not expected to affect the final conclusion, as the new crude blend loads the bottom section rather than the top section.
Target case capacities. The vali-dated models were used to predict the capacities of the downstream units for clean-fuel specifications. TABLE 8 lists the expected incremental unit capacities, as percentages, for the new crude blend.
Results and observations. For the new crude assay, it was found that the product yields from the CDU had de-creased, except for kerosine and RCO.
The reduction in diesel was expected, since Euro 5 diesel specifications require a reduction in the heavy end of the die-sel fraction to meet the T95 distillation specification. Therefore, the reduced crude circuit, including pumps, piping, the vacuum charge heater, the vacuum tower and associated equipment, are prone to bottlenecks due to increased flow, especially when meeting clean-die-sel specifications.
A suggested modification was to avoid revamping the vacuum heater by diverting the CDU overflash material directly to the vacuum tower, bypassing
Current crude New crude oil blend
00
650 700 750 800 850 900 950 1,000 1,050 1,100 1,150 Yields comparison
FIG. 2. Yield comparison between the present crude blend and the new crude blend.
TABLE 1. Simulation operating parameters vs. actual operating parameters for the CSF Simulation
number CSF parameter Unit Plant data
Simulation data
2 Flash drum temperature °C 142.88 142.88
3 Flash drum pressure kg/cm2g .8 .7
10 Debutanizer bottom temperature °C 174.6 174
11 Debutanizer top temperature °C 75.3 75
12 Debutanizer top pressure kg/cm2g 10.33 10.33
13 Debutanizer bottom pressure kg/cm2g 11.27 11.27
14 Debutanizer refl ux rate m3/hr 61 80
15 Tray 6 temperature °C 85.5 91
16 Naphtha splitter (NS) top temperature °C 78.7 64
17 NS top pressure kg/cm2g .84 .84
18 NS bottom temperature °C 227 239
19 NS refl ux rate1 m3/hr 1.7 45
1 A 25-m3/hr quench fl ow at 225°C is taken.
TABLE 2. Simulation lab data vs. actual lab data for the CSF Simulation
number CSF parameter Unit Plant data
Simulation data CSF LN
1 5% distillation temperature °C 44 31
2 Final boiling point °C 91 91
CSF HN
3 5% distillation temperature °C 85 60
4 Final boiling point °C 171 172
5 Specifi c gravity .736 .72
6 Sulfur ppm 178 364
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Refining Developments
the heater. This option will eliminate the need for increased vacuum charge heater capacity.
The vacuum bottoms flow with the new crude blend is expected to be in-creased by about 8%. Therefore, FO sales from the refinery must be raised to maintain the management of black oil (i.e., reduced crude oil, VGO and vacu-um bottoms) in the refinery. Alternative-ly, the capacity of the asphalt plant must be increased to consume extra vacuum bottoms as asphalt product, to maintain sustainable refinery charge rates with the new crude oil blend. TABLES 2, 4, 6 and 7 show the properties of the streams from the crude stabilization facility (CSF), the CDU and the VDU.
In conclusion, black oil management will be impacted significantly at present refinery capacity with 100% new crude oil blend processing, especially when meeting clean-diesel specifications.
Sustaining the present refinery charge rate is not an option for meeting clean-fuel specifications with 100% of the new crude blend, until black oil management is improved.
Expected throughput reduction without investment. The expected throughput reduction for the new crude oil blend, without any capital invest-ment, was determined by keeping the VDU heater capacity the same as the original installed capacity. The CDU throughput will be reduced to 94% if the current diesel specification is to be main-tained, and to 91% to meet the Euro 5 diesel specification.
Expected HC capacity. The VGO production with the new crude oil blend will be higher than with the exist-ing crude oil blend. Therefore, the HC capacity with the new crude oil blend will be constrained during catalyst chan-geout, plant shutdowns and other oc-casions to reprocess the extra VGO in storage. The required HC capacity for processing VGO generated during cata-lyst changeout is 20% higher than the present HC capacity.
Semi-regenerative reformer and HT capacities. The semi-regenerative catalytic reformer will not be operating at full capacity at the present CSF and CDU throughputs, as the heart-cut naphtha yield is lower in the new crude blend.
Moreover, the reformer feed pre-cursor N + 2A content of the heart-cut
naphtha was also found to be lower for the new crude blend. This will severely
affect the octane number of fuel from the reformer; consequently, total
gaso-TABLE 3. Simulation operating parameters vs. actual operating parameters for the CDU Simulation
number CSF parameter Unit Plant data
Simulation data
5 Top tower temperature °C 131 131
6 Top tower pressure kg/cm2g 1.02 1.02
7 Refl ux drum temperature °C 42 42
8 Refl ux drum pressure kg/cm2g .4 .4
9 Refl ux fl owrate m3/hr 8.97 2
10 Top pumparound (PA) fl owrate m3/hr 400 400
11 Top PA drawoff temperature °C 153 158
12 Top PA return temperature °C 63 63
13 Top PA ⌬T °C 90 95
15 Kerosine drawoff temperature °C 195 202
17 LGO drawoff temperature °C 250 255
18 LGO PA fl owrate m3/hr 379 379
19 LGO PA return temperature °C 174 174
20 LGO PA ⌬T °C 76 81
22 HGO drawoff temperature °C 337 329
23 HGO PA return temperature °C 215 215
24 HGO PA ⌬T °C 122 114
25 Flash zone temperature °C 371 371
26 Flash zone pressure kg/cm2g 1.28 1.28
27 Tower bottom temperature °C 363 365.8
28 Tower bottom pressure °C 1.4 1.4
30 Stripping steam at 3.5 kg/cm2g kg/hr 5,511 5,511
TABLE 4. Simulation lab data vs. actual lab data for CDU Simulation
number CDU parameter Unit Plant data
Simulation data
1 Overhead liquid fl owrate m3/hr 138 136
2 Final boiling point °C 172 176.5
3 Specifi c gravity .7 .7265
4 Sulfur wt% .05 .05
5 Kerosine fl owrate m3/hr 88 88
6 Final boiling point °C 267 258
7 Specifi c gravity .789 .8053
8 Sulfur wt% .1871 .382
9 LGO fl owrate °C 176 176
10 Final boiling point °C 363 357
11 Specifi c gravity .8415 .8463
12 Sulfur wt% 1.1 1.14
13 HGO fl owrate m3/hr 97 97
14 100% distillation temperature °C 407 417
15 Specifi c gravity .88 .8957
16 Sulfur wt% 1.9 1.8
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Refining Developments
line production will be reduced. To fill the capacity of the semi-regenerative reformer, a deep stabilization operation must be carried out at the CSF, which will recover more heavy naphtha (HN) from the crude before it is sold to power
plants. This will make crude oil to the CDU relatively heavier; consequently, more black oil production is expected.
Another option is to increase the cut-point of HN from the CDU. Since the gasoline endpoint is 210°C and,
typical-ly, the final boiling point of the reformate increases by around 7°C in the semi-re-generative reformer, the endpoint for the NHT feed can be raised further. How-ever, other contaminants and reformer catalyst performance at such high NHT feed endpoints should be investigated before making such a decision. There-fore, a thorough investigation with the reformer licensor is recommended to evaluate the potential consequences of increasing the HN endpoint.
As the diesel specification has been changed from a T85 of 350°C maximum to a T95 of 360°C maximum for clean-fuels production, the diesel output from the CDU will be reduced by 83%. The tail-end of the diesel fraction goes to the HC, and it is eventually converted to naphtha and diesel. Therefore, neither the semi-regenerative reformer nor the DHT is expected to have capacity con-straints. The capacity available in the distillate hydrotreater (DHT) can be supplemented with kerosine feed and normal sour diesel feed during low jet fuel lifting from the refinery.
Feed sulfur to the kerosine HT unit is around 0.2 wt%. Sour kerosine can be processed in the DHT, along with nor-mal sour diesel feed, during the catalyst changeover in the kerosine unit, through proper tank management. The time re-quired to consume this extra inventory in the DHT is less than one month. Al-though there is no capacity limit in the DHT for processing this kerosine, meet-ing the diesel flashpoint will be a chal-lenge, and the DHT stripper should have adequate capacity to produce an on-specification diesel flashpoint when the DHT feed is blended with atmospheric gasoil and kerosine.
With only LGO and HGO, the flash-point of diesel is 60°C, and the required diesel flashpoint specification is 55°C. A small quantity of diesel is cracked in the DHT, and the DHT stripper removes the light ends to meet the diesel flash-point specification.
The present kerosine flashpoint is 46°C. Blending this kerosine in a die-sel feed pool will decrease the diedie-sel flashpoint; therefore, the DHT strip-per requires an additional reboiling duty to meet the diesel flashpoint. A minor modification is required in the DHT stripper to meet the diesel flash-point. During this operation, the DHT
TABLE 5. Simulation operating parameters vs. actual operating parameters for the VDU Simulation
number VDU parameter Unit Plant data
Simulation data
1 Top tower temperature °C 48 52
2 Top tower pressure mmhg 18 18
3 LVGO PA fl owrate m3/hr 137 158
4 LVGO PA drawoff temperature °C 101 101
5 LVGO PA return temperature °C 49 49
7 HVGO PA fl owrate m3/hr 212 220
8 HVGO PA drawoff temperature °C 281 287
9 HVGO PA return temperature °C 101 101
11 Indirect recycle fl owrate1 m3/hr 50 50
12 Indirect recycle temperature °C 350 350
13 Indirect recycle pressure kg/cm2 3.5 3.5
14 Slop wax fl owrate m3/hr 13 19.7
15 Slop wax drawoff temperature °C 386 381
16 Flash zone temperature °C 392 391.4
17 Flash zone pressure mmhg 35.65 35.65
18 Tower bottom temperature2 °C 366 365.3
19 Vacuum tower bottoms fl owrate (vacuum slop included)
m3/hr 165 161
20 Stripping steam at 3.5 kg/cm2g kg/hr 2,298 2,298
21 RCO fl owrate m3/hr 338 339
1 Indirect recycle is the HC main fractionator bottoms used as a wash liquid.
2 A 25-m3/hr quench fl ow at 225°C is taken.
TABLE 6. Simulation lab data vs. actual lab data for the VDU Simulation
number VDU parameter Unit Plant data
Simulation data LVGO
1 5% distillation temperature °C 287 272
2 90% distillation temperature °C 463 441
3 Specifi c gravity 0.8984 0.8728
4 Sulfur wt% 1.98 1.93
HVGO
5 5% distillation temperature °C 387 386
6 90% distillation temperature °C 539 524
7 95% distillation temperature °C 562 551
8 Specifi c gravity 0.9196 0.9254
9 Sulfur wt% 2.3 2.6
HC feed
10 Specifi c gravity 0.919
11 Sulfur wt% 2.53
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Refining Developments
stripper overhead produces more light naphtha (LN), which must be accom-modated in the overall naphtha balance for the refinery to fill the NHT and the catalytic reformer.
Takeaway. In summary, changing the crude oil slate from semi-light crude to a mixture of light and medium crude will result in several changes. The changes have been identified using process simu-lation software.
Distillate recovery from the CDU will be impacted at the expense of extra VGO and vacuum bottoms production.
The extra VGO production requires an increase in HC capacity and, therefore, an HC revamp is required. The extra vacuum bottoms production increases FO production.
Therefore, with the new crude blend, no spare capacity is available at the HC to reprocess VGO recovered from the HC feed filter backwash, or to maintain the present refinery charge rate during occasions such as catalyst changeout.
The feed quality for catalytic process units, such as the CR, the DHT, the HC and others, will be inferior and will
im-pact yields, catalyst performance and cycle lengths. A revamp is required to process this new crude oil blend and to meet clean-fuel specifications.
SAID A. AL-ZAHRANI is the general supervisor in the process and control systems department at Saudi Aramco. He is the chairman of the multi-disciplinary product specifications committee, tasked with managing various issues related to Saudi Aramco products and fuel specifications. Mr. Al-Zahrani holds a degree in chemical engineering from King Fahd University of Petroleum and Minerals, and began his career at Saudi Aramco as a process engineer in the Ras Tanura refinery. He is a member of several local and international societies and an officer of the American Institute of Chemical Engineers, Saudi Arabian chapter.
SAMIT ROY is an engineering consultant at Saudi Aramco’s downstream process engineering division.
A chemical engineering graduate, he has more than 33 years of experience in process engineering and technical services. His experience includes 21 years in Saudi Aramco refining and engineering services and 12 years at Indian refineries. He has worked at most refinery units associated with distillation, hydroprocessing and gas treating.
EDWIN BRIGHT has over 17 years of experience in the petroleum refining industry. Before joining Saudi Aramco, he worked for Reliance Industries Ltd., Indian Oil Corp., ATV Petrochemicals and Foster Wheeler India Ltd. He holds a bachelor’s degree in chemical engineering and master’s degrees in petroleum refining and petrochemicals from Anna University’s Alagappa College of Technology in Chennai, India.
He also earned a master’s degree in management from the Asian Institute of Management in Manila.
TABLE 7. Key properties of the feed to secondary processing units
Unit parameter Unit Plant data
Simulation:
30% new crude
Simulation:
100% new crude CDU overhead liquid
Final boiling point °C 172 176.5 171
Specifi c gravity 0.7 0.7265 0.7307
Sulfur wt% 0.05 0.05 0.075
Kerosine
Final boiling point °C 267 258 254
Specifi c gravity 0.789 0.8053 0.8055
Sulfur wt% 0.1871 0.382 0.3525
LGO
Final boiling point °C 363 357 360
Specifi c gravity 0.8415 0.8463 0.8492
Sulfur wt% 1.1 1.14 1.22
HGO
100% °C 407 417 410
Specifi c gravity 0.88 0.8957 0.9
Sulfur wt% 1.9 1.8 1.94
LVGO
90% °C 463 441 470
Specifi c gravity 0.8984 0.8728 0.878
Sulfur wt% 1.98 1.93 2.1
HVGO
90% °C 539 524 524
95% °C 562 551 551
Specifi c gravity 0.9196 0.9254 0.923
Sulfur wt% 2.3 2.6 2.6
DHT feed
Specifi c gravity 0.85 0.8638 0.8604
Sulfur wt% 1.4 1.4 1.4
HC feed
Specifi c gravity 0.92 0.919 0.9168
Sulfur wt% 2.3 2.5 2.5
TABLE 8. Required capacities of diff erent units for the new crude blend
Unit parameter
With 100% new crude, specifi cation of T95 at 360°C for diesel,
% of current operation
CSF NS fl owrate 82
NHT feed (excluding
HC naphtha) 86
Kerosine 105
LGO plus HGO 83
RCO 115
HC combined feed1 120
HC fresh feed 124
Vacuum bottoms 108
Indirect recycle2 100
1 A 25-m3/hr quench fl ow at 225°C is taken.
2 Indirect recycle is the HC main fractionator bottoms used as a wash liquid.