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.
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.
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withdrawn as a separate cut and is used to increase the preheat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
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
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 systems 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
Fig. 4. Current
configuration of naphtha circuit.
Fig. 5. Configuration
of naphtha management system after
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.
Impact of hot streams going directly to the other
Fig. 6. Base-case
network after modifications.
The changes made in the base-case network are listed
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
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
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
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
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.
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
bachelors degree in chemical engineering and
masters degrees in petroleum refining and petrochemicals from AC Tech,
Anna University, Chennai. He also earned a
masters degree in management from the Asian
Institute of Management in Manila.
Samit Roy is an engineering consultant
at Saudi Aramcos 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.