Debottleneck Crude-unit Preheat Exchanger Network Inefficiencies _ Hydrocarbon Processing _ February 2012

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Debot Debottle tlene neck ck crude-u crude-unit nit prehe preheat at exchan changer ger netw network ork ineff inefficien iciencies cies | Hydro Hydroca carbo rbon n Proce Processin ssing g | Februa February ry 2012 2012

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Debottleneck crude-unit preheat exchanger network inefficiencies inefficiencies | Al-Zah A l-Zah ran i, S., S., Saudi Aramc o, Saudi Arabia; Bright , E. , Saudi Ar amc o, Dhahran, Dhahran, Saudi Saudi Ara A rabia bia;; Roy , S. S. , Saudi Ar amco , Dhahran, Dhahran, Saudi Saudi Arabia 02.01.2012

Simulation Simulation models can be effectiv effectiv ely used to o ptimize ptimize heat transfer transfer and bo ost oper ational performance performance Keywords: In this case history, history, a c rude distillation distillation unit unit (CD (CDU) preheat train net work in a Saudi Saudi Ar amco refinery refinery was was simulated and analyzed for anticipated modifications modifications to the network. This analysis helped eliminate ineffi inefficiencies ciencies in the ne twork, and, based on the insights insights from the analysis, v arious options were generated and the ex isting isting network was reconfigured. The The rec onfiguration onfiguration allowed the temperature o f the crude pre heat network, which which proc esses Arab Light crude oil, to be increased increased to the max imum of 27 7 °C from a previous temperature o f 261°C. 261°C. Existing c onfiguration. onfiguration. Desalted crude from the tank is heated by the cr ude c olumn top pumparound, light gasoil (LG (LGO) produc t, heav y gasoil (HGO (HGO) prod uct , LGO LGO pumpar ound (LGO (LGO PA), HGO HGO pumpar ound (HGO (HGO PA), heav y v acu um gaso il (HVGO (HVGO)) pumparound and v acuum residue (V R) product, as shown in Fig. Fig. 1 in ex changers E1 to E7 , respec tively . The current c rude preheat te mperature entering the CDU CDU furnace furnace is around 2 61°C. This This ex changer network is v alidated using using heat ex changer design software and by adjusting adjusting the fouling fouling coefficients.

Fig. 1. Current 1. Current co nfiguration of CDU CDU prehea t train.

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Debottleneck crude-unit preheat exchanger network inefficiencies | Hydrocarbon Processing | February 2012

Modifications required. The base-case netwo rk was altered for anticipated mo difications in the future. The reasons for the modifications are listed below: • Vac uum slo p circ uit. In the current c onfiguration (Fig. 2), the vacuum slop is recy cled to the v acuum tower through the vac uum furnace. The purpose of this recy cle is to recov er the VGO compo nents and send the VGO to the hydro crac ker; howev er, this is not achieve d in the current ope ration due to vac uum furnace limitations and insufficient separation in the wash section. A s a re sult, this vac uum slop stream (which is lower in visco sity) goes with the vac uum tower botto ms. 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 vac uum slop circuit.

To address this conc ern, the vac uum slop stream from the v acuum tower is available at a temperature of 380°C, which is withdrawn as a separate c ut and is used to increase the preheat temperature of the crude . This propo sed new ex changer is co nfigured to be in parallel with the existing heat exc hanger E4 in Fig. 1 . Fig. 3 shows the rerouting of the vac uum slop.

Fig. 3. Modifications in vacuu m slop circ uit.

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• Future splitter configuration. To meet the clean gasoline spec ification of 1 % benzene in gasoline, the ex isting naphtha splitter must remove the benzene prec ursors in the catalytic re former feed by incre asing the initial boiling point of the heav y naphtha. This proce ss requires a higher re bo iler duty. In addition, the heavy naphtha from the hy droc racker needs to be pro cessed in the naphtha splitter, as this feed also contains be nze ne p re curso rs. Currently, hyd rocr acker heav y naphtha is not part of the naphtha splitter feed. The hydro crac ker heavy naphtha feed v olume is 12,50 0 barrels per day (bpd), and the existing naphtha splitter capacity is 23,0 00 bp d. Figs. 4 and 5 sho w the naphtha sy ste m’s c urre nt and planned c onfigur ations, respe ctiv ely . A s the curre nt naphtha splitter c annot handle this higher throughput with higher re bo iler requirement, the ex isting naphtha splitter will be mothballed. The ex isting reboiler, which uses HGO PA flow and gives a duty of 10 .4 million kilocalories pe r hour (MMkcal/hr), will also be mothballed. High-pressure steam will be used in the rebo iler of the new naphtha splitter to meet the higher re bo iler requirements. For the co lumn to be in heat balance, this 1 0.4 MMkc al/ hr o f heat re mo v al is r eq uired . In the prop osed ex changer ne two rk, this stream (HGO CR) will be used t o pre heat the cr ude.

Fig. 4. Current co nfiguration of naphtha circuit.


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

Sy nthesis of crude preheat train. A new, pre liminar y heat e x changer ne two rk (Fig. 6) was sy nthesized to ac commo date the abov e modifications. While modifying the c rude pr eheat train network, the following impact o n the e quipment was kept in mind: • Preve ntion of vaporizations in the furnace pass-co ntrol v alves, as it is difficult to c ontrol two-phase flows acro ss pass-control v alves. I nadequate flow in the furnace pass flows will also lead to c oking. • Column heat balance. • Column hy draulics. • Impact o f hot streams going directly to the other unit.

Fig. 6. Base-case netwo rk after modifications. www.hydrocarbonprocessing.com/Article/2972114/Debottleneck-crude-unit-preheat-exchanger-network-inefficiencies.html?Print=true

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The changes made in the base-case network are listed below: • Ex changer N1 was added parallel to E4 (see Fig. 6) using vac uum slop (v acslop) and v acuum residue e x-E7 as the hot fluid. This modification is required to improv e the v iscosity o f the v acuum residue to the asphalt ox idizer. The current visco sity of the feed to the asphalt oxidizer is 1,5 00 c entistokes (cst), and the required v isc os ity is 2,0 00 c st. • Ano the r ex changer N2 (E5-2 , similar to E2) was added pa rallel to E2 using HGO PA fluid ex -E5 (he re afte r referred to as E5-1 ) as the hot fluid. This modification is performed to ac co mmodate the 1 0.4-MMkcal/hr duty in the HGO PA c ircuit. • Incr eased area in E4 from the 2-parallel-1-series arrangement to a 2-parallel-2-series design and added co oler N3 do wnstream of E4. Due to the first two modifications, the inlet temperature to E4 has increased, which de creases the logarithmic mean temperature difference (LMTD) av ailable across the unit. Since E4 is the LGO PA e xc hanger, the c olumn will not be in he at b alanc e if the req uired heat remov al is not performed. The re quire d duty was 1 8.8 MMkcal/hr, and the available duty was 1 2.7 MMkcal/hr (see Table 1). Therefore, additional area and a coo ler were added in the LGO PA circ uit t o mee t the duty re quire me nt o f the c olumn.

The required HGO PA duty is 26.8 MMkcal/hr, and the av ailable duty is 29.8 MMkcal/hr. A s the heat remov ed in HGO PA is higher by 3 MMkcal/hr, the requirement o f LGO PA duty will come do wn by 3 MMkcal/hr. As bo th LGO and HGO are mix ed ou tsid e o f the c olumn and go to the die sel hy drot re ate r (DHT), the splitting o f the duty between LGO and HGO pumparound is not a co ncern from a separation point of v iew. Howev er, it does impact the column draw temperature, which will slightly reduce the LMTD acro ss E3 (HGO produc t/cr ude ex changer) and E5 (HGO PA/c rude ex changer). Results of network modification. In the modified network, the o btained preheat temper ature was 266°C. The duty , LMTD and area of eac h ex changer in the network are presented in Table 1. From Table 1, it can be observ ed that: • Ex changer E6, which has a higher are a, is ex periencing the lo west LMTD; therefore, any modification that increases the LMTD will significantly increase the heat rec ov ered from E6. www.hydrocarbonprocessing.com/Article/2972114/Debottleneck-crude-unit-preheat-exchanger-network-inefficiencies.html?Print=true

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• The exc hanger prece ding exc hanger E6 is heated by HGO circulating reflux (CR), which is at 3 37 °C; this is higher than the hot str eam (HVGO CR) temper ature o f E6, which has d ec rease d the LMTD in E6. This preliminary network was analyzed for possible improv ement in the preheat temperature. The analy sis indicated that heat recov ery can be increased by 45% by boo sting the area by 5 6% (see Table 2).

The analy sis also indicated that the driv ing forc e acro ss exc hanger E7 further limited the heat recov ery . Fig. 7 displays the driving-forc e plot. The figure indicates that the driving force in E7 can be increased b y decr easing the inlet temperature in E7 . This temperature adjustment can be achiev ed by operating E5 in parallel with E7 .

Fig. 7 . Driving-forc e plot for b ase-case network.

Case 1. Based on the insights derive d from Table 1 and Fig. 7 , to improve the heat rec ov ery , the crude stream in E7 and E5 was split by o perating E5 in parallel with E7 . The objec tive of this modification is to increase the LMTD across E7 and E6. Howev er, it also decr eases the LMTD across E5-1 . The net e ffect is shown in Table 3, and the mo dified network is shown in Fig. 8. With this arrangement, the pre heat temperature has increased from 26 6°C to 2 69°C.


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

Case 2. From LMTD and approach data in Table 3, it can be inferred that heat reco v ery in E5-1 can still be improv ed by increasing the area. Hence, another c ase study was performed by adding two similar exc hangers in a series in E5-1 . The results are tabulated in Table 4. The preheat was found to be increased to 27 7 °C.


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The HGO PA is now pro v iding an ex tra 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 co oler duty co mes down to 2.6 MMkcal/hr. HP

T he authors Edwin Bright h as over 17 y ears of experience in the petroleum refining industry . Before joining Saudi A ramc o, he worked for Reliance I ndu stries Ltd., Indian Oil Corp., AT V Petrochemicals and Foster Wheeler India Ltd. He ho lds a bach elor’s degree in ch em ical engineerin g and m aster’s degrees in petroleum refining and petrochemicals from AC T ech, Ann a University, Chenn ai. He also earned a master’s degree in m anagement from the A sian I nstitute of Management in Manila. Samit Roy is an engineering c onsultant at Saudi Aramc o’s downstream process engineering division under the process control and system s department. A graduate in ch em ical engineering, he h as more th an 33 y ears of broad experience in the process engineering and tech nic al serv ices areas of oil refining and gas processin g plants. His experience in clu des 21 y ears in Saudi A ram co refining and engineering serv ices and 12 years at Indian refineries. He has wo rked at m ost of th e refinery process un its associated with distillation, hy droprocessing and gas treating plants. Said A. Al-Zahrani is the general supervisor in the process and control sy stems department at Saudi Aramco. He is the ch airm an of the m ulti-disciplinary product specifications com m ittee, tasked with m anaging various issues related to Saudi Aramc o products and fuel specifications. Al-Zah ran i holds a degre e in chem ic al en gin eerin g from King Fah d Univ ersity of Pet rol eu m and Minerals, and began h is career at Saudi Aram co as a process engine er in the Ras T anura refinery . He is a mem ber of several local and internation al societies and an officer of the Am erican I nstitute o f Ch em ic al En gin eers, Saudi Arabian chapt er.


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Debottleneck Crude-unit Preheat Exchanger Network Inefficiencies _ Hydrocarbon Processing _ February 2012

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