Hydrocarbon Processing Copying and distributing are prohibited without permission of the publisher
Email a friend
  • Please enter a maximum of 5 recipients. Use ; to separate more than one email address.

Debottleneck crude-unit preheat exchanger network inefficiencies

02.01.2012  |  Al-Zahrani, S.,  Saudi Aramco, Saudi ArabiaBright, E. ,  Saudi Aramco, Dhahran, Saudi ArabiaRoy, S. ,  Saudi Aramco, Dhahran, Saudi Arabia

Simulation models can be effectively used to optimize heat transfer and boost operational performance

Keywords: [CDU] [exchanger] [debottleneck] [furnace] [preheat train]

In this case history, a crude distillation unit (CDU) preheat train network in a Saudi Aramco refinery was simulated and analyzed for anticipated modifications to the network. This analysis helped eliminate inefficiencies in the network, and, based on the insights from the analysis, various options were generated and the existing network was reconfigured. The reconfiguration allowed the temperature of the crude preheat network, which processes Arab Light crude oil, to be increased to the maximum of 277°C from a previous temperature of 261°C.

Existing configuration.

Desalted crude from the tank is heated by the crude column top pumparound, light gasoil (LGO) product, heavy gasoil (HGO) product, LGO pumparound (LGO PA), HGO pumparound (HGO PA), heavy vacuum gasoil (HVGO) pumparound and vacuum residue (VR) product, as shown in Fig. 1 in exchangers E1 to E7, respectively. The current crude preheat temperature entering the CDU furnace is around 261°C. This exchanger network is validated using heat exchanger design software and by adjusting the fouling coefficients.


  Fig. 1. Current configuration of CDU preheat train. 

Modifications required.

The base-case network was altered for anticipated modifications in the future. The reasons for the modifications are listed below:

• Vacuum slop circuit. In the current configuration (Fig. 2), the vacuum slop is recycled to the vacuum tower through the vacuum furnace. The purpose of this recycle is to recover the VGO components and send the VGO to the hydrocracker; however, this is not achieved in the current operation due to vacuum furnace limitations and insufficient separation in the wash section. As a result, this vacuum slop stream (which is lower in viscosity) goes with the vacuum tower bottoms. The mingling of streams deteriorates the feed to the asphalt oxidizer and creates operational problems in meeting the penetration property of the asphalt.


  Fig. 2. Current configuration of vacuum slop circuit. 

To address this concern, the vacuum slop stream from the vacuum tower is available at a temperature of 380°C, which is withdrawn as a separate cut and is used to increase the preheat temperature of the crude. This proposed new exchanger is configured to be in parallel with the existing heat exchanger E4 in Fig. 1. Fig. 3 shows the rerouting of the vacuum slop.


  Fig. 3. Modifications in vacuum slop circuit. 

• Future splitter configuration. To meet the clean gasoline specification of 1% benzene in gasoline, the existing naphtha splitter must remove the benzene precursors in the catalytic reformer feed by increasing the initial boiling point of the heavy naphtha. This process requires a higher reboiler duty. In addition, the heavy naphtha from the hydrocracker needs to be processed in the naphtha splitter, as this feed also contains benzene precursors.

Currently, hydrocracker heavy naphtha is not part of the naphtha splitter feed. The hydrocracker heavy naphtha feed volume is 12,500 barrels per day (bpd), and the existing naphtha splitter capacity is 23,000 bpd. Figs. 4 and 5 show the naphtha system’s current and planned configurations, respectively. As the current naphtha splitter cannot handle this higher throughput with higher reboiler requirement, the existing naphtha splitter will be mothballed. The existing reboiler, which uses HGO PA flow and gives a duty of 10.4 million kilocalories per hour (MMkcal/hr), will also be mothballed. High-pressure steam will be used in the reboiler of the new naphtha splitter to meet the higher reboiler requirements. For the column to be in heat balance, this 10.4 MMkcal/hr of heat removal is required. In the proposed exchanger network, this stream (HGO CR) will be used to preheat the crude. 


  Fig. 4. Current configuration of naphtha circuit. 


  Fig. 5. Configuration of naphtha management system after
  clean-fuel implementation. 

Synthesis of crude preheat train.

A new, preliminary heat exchanger network (Fig. 6) was synthesized to accommodate the above modifications. While modifying the crude preheat train network, the following impact on the equipment was kept in mind:
• Prevention of vaporizations in the furnace pass-control valves, as it is difficult to control two-phase flows across pass-control valves. Inadequate flow in the furnace pass flows will also lead to coking.
• Column heat balance.
• Column hydraulics.
• Impact of hot streams going directly to the other unit.


  Fig. 6. Base-case network after modifications. 

The changes made in the base-case network are listed below:

• Exchanger N1 was added parallel to E4 (see Fig. 6) using vacuum slop (vacslop) and vacuum residue ex-E7 as the hot fluid. This modification is required to improve the viscosity of the vacuum residue to the asphalt oxidizer. The current viscosity of the feed to the asphalt oxidizer is 1,500 centistokes (cst), and the required viscosity is 2,000 cst.

• Another exchanger N2 (E5-2, similar to E2) was added parallel to E2 using HGO PA fluid ex-E5 (hereafter referred to as E5-1) as the hot fluid. This modification is performed to accommodate the 10.4-MMkcal/hr duty in the HGO PA circuit.

• Increased area in E4 from the 2-parallel-1-series arrangement to a 2-parallel-2-series design and added cooler N3 downstream of E4.

Due to the first two modifications, the inlet temperature to E4 has increased, which decreases the logarithmic mean temperature difference (LMTD) available across the unit. Since E4 is the LGO PA exchanger, the column will not be in heat balance if the required heat removal is not performed. The required duty was 18.8 MMkcal/hr, and the available duty was 12.7 MMkcal/hr (see Table 1). Therefore, additional area and a cooler were added in the LGO PA circuit to meet the duty requirement of the column.


The required HGO PA duty is 26.8 MMkcal/hr, and the available duty is 29.8 MMkcal/hr. As the heat removed in HGO PA is higher by 3 MMkcal/hr, the requirement of LGO PA duty will come down by 3 MMkcal/hr. As both LGO and HGO are mixed outside of the column and go to the diesel hydrotreater (DHT), the splitting of the duty between LGO and HGO pumparound is not a concern from a separation point of view. However, it does impact the column draw temperature, which will slightly reduce the LMTD across E3 (HGO product/crude exchanger) and E5 (HGO PA/crude exchanger).

Results of network modification.

In the modified network, the obtained preheat temperature was 266°C. The duty, LMTD and area of each exchanger in the network are presented in Table 1. From Table 1, it can be observed that:

• Exchanger E6, which has a higher area, is experiencing the lowest LMTD; therefore, any modification that increases the LMTD will significantly increase the heat recovered from E6.

• The exchanger preceding exchanger E6 is heated by HGO circulating reflux (CR), which is at 337°C; this is higher than the hot stream (HVGO CR) temperature of E6, which has decreased the LMTD in E6.

This preliminary network was analyzed for possible improvement in the preheat temperature. The analysis indicated that heat recovery can be increased by 45% by boosting the area by 56% (see Table 2).


The analysis also indicated that the driving force across exchanger E7 further limited the heat recovery. Fig. 7 displays the driving-force plot. The figure indicates that the driving force in E7 can be increased by decreasing the inlet temperature in E7. This temperature adjustment can be achieved by operating E5 in parallel with E7.


  Fig. 7. Driving-force plot for base-case network. 

Case 1. Based on the insights derived from Table 1 and Fig. 7, to improve the heat recovery, the crude stream in E7 and E5 was split by operating E5 in parallel with E7. The objective of this modification is to increase the LMTD across E7 and E6. However, it also decreases the LMTD across E5-1. The net effect is shown in Table 3, and the modified network is shown in Fig. 8. With this arrangement, the preheat temperature has increased from 266°C to 269°C.


  Fig. 8. Modified network based on E5 operating in parallel with E7. 

Case 2. From LMTD and approach data in Table 3, it can be inferred that heat recovery in E5-1 can still be improved by increasing the area. Hence, another case study was performed by adding two similar exchangers in a series in E5-1. The results are tabulated in Table 4. The preheat was found to be increased to 277°C.



The HGO PA is now providing an extra 4.2 MMkcal/hr more than required, which will reduce the LGO PA duty requirement by the same amount for the column to be in heat balance. Then, the required LGO PA cooler duty comes down to 2.6 MMkcal/hr. HP

The authors 

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 AC Tech, Anna University, Chennai. He also earned a master’s degree in management from the Asian Institute of Management in Manila. 

Samit Roy is an engineering consultant at Saudi Aramco’s downstream process engineering division under the process control and systems department. A graduate in chemical engineering, he has more than 33 years of broad experience in the process engineering and technical services areas of oil refining and gas processing plants. His experience includes 21 years in Saudi Aramco refining and engineering services and 12 years at Indian refineries. He has worked at most of the refinery process units associated with distillation, hydroprocessing and gas treating plants. 

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. 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. 

Have your say
  • All comments are subject to editorial review.
    All fields are compulsory.

Related articles


Sign-up for the Free Daily HP Enewsletter!

Boxscore Database

A searchable database of project activity in the global hydrocarbon processing industry


Is 2016 the peak for US gasoline demand?




View previous results

Popular Searches

Please read our Term and Conditions and Privacy Policy before using the site. All material subject to strictly enforced copyright laws.
© 2016 Hydrocarbon Processing. © 2016 Gulf Publishing Company.