Thermal Process Integration in the AVDU A12/2 Crude

ISSN 0040-5795, Theoretical Foundations of Chemical Engineering, 2009, Vol. 43, No. 6, pp. 906–917. © Pleiades Publishing, Ltd., 2009.Original Russian Text © L.L. Tovazhnyanskii, P.A. Kapustenko, L.M. Ul’ev, S.A. Boldyrev, O.P. Arsen’eva, M.V. Tarnovskii, 2009, published in Teoreticheskie OsnovyKhimicheskoi Tekhnologii, 2009, Vol. 43, No. 6, pp. 665–676.

906

Reduction in the consumption of fuel and powerresources in industry can be achieved by the wide imple-mentation of modern energy-saving technologies and thecreation of highly efficient power engineering complexes.Although this challenge is far from being new for the pet-rochemical industry, two factors appeared in the last twodecades which attached special significance to energy sav-ing in CIS countries and prompted the extension of therange of methods and devices used for this purpose.

First, the rise in energy prices has encouraged more sav-ing of energy to reduce total costs. Moreover, the operationof all of the plants designed and built in the time of lowenergy prices is currently far from optimal in terms ofenergy consumption.

Second, the rate of economic growth has significantlydecreased, and, therefore, so has the potential of puttingnew plants into operation and implementing new pro-cesses; therefore, more and more attention is shiftedtowards increasing the efficiency of using existing equip-ment.

Oil refining and petrochemical production are energyintensive, and energy consumption considerably affects theproduction cost. Depending on the degree of oil conver-sion, oil composition, the range and quality of products,equipment performance, and other factors, auxiliaryenergy consumption by refineries is equivalent to 6–10%of refined oil. Of the total energy consumption, 55–65% isfor process fuel, 30–35% is for heat, and 8–12% is for elec-tric power [1].

The most energy-intensive processes are hydrocrack-ing, catalytic cracking and reforming, coking, and the pro-duction of oil. Crude distillation processes are less energyintensive, but all of the crude oil entering a refinery is sub-jected to atmospheric/vacuum distillation, which takes up

about 50% of the total energy consumption. Consequently,a reduction of energy consumption is equally important forall oil-refining processes. The most energy-intensive petro-chemical production processes are the syntheses of lowerolefins, methanol, hydrogen, and aromatic hydrocarbons.

The energy consumption by modern non-CIS oil refin-eries is 3000–3500 MJ, or 100–114 kg of standard fuel perton of refined oil. For CIS oil refineries, the energy con-sumption in terms of standard fuel is approximately thesame, but the degree of oil conversion is much lower. Withan increase in the degree of oil conversion and an expan-sion of petrochemical production, the energy consumptionincreases, and so does the importance of energy saving.The main ways of improving the competitiveness of CISoil refining and petrochemical production are to increasethe degree of oil conversion and significantly reduce theenergy consumption by all of the processes [2].

In CIS oil refineries, most of the processing units wereconstructed in the 1960s–1970s, when the cost of energyresources was very low and energy saving was of little con-sequence, and almost no energy-saving arrangements wereundertaken. Therefore, the energy consumption by thebasic oil refining and petrochemical production processesin these processing units is 30–60% higher than that inmodern units with significant thermal processes integra-tion.

The specific energy consumption can be reduced byupgrading individual production systems, units, and plantsand by rationalizing and improving production processes.

Previously [3], we designed a project of partial processintegration in a crude distillation unit of the KremenchugOil Refinery in Ukraine. It was shown that the implemen-tation of the process retrofit project designed by pinch anal-ysis methods can reduce the consumption of hot and cold

Thermal Process Integration in the AVDU A12/2 Crude Distillation Unit during Winter Operation

L. L. Tovazhnyanskii

a

, P. A. Kapustenko

a

, L. M. Ul’ev

a

, S. A. Boldyrev

a

, O. P. Arsen’eva

a

, and M. V. Tarnovskii

b

a

Kharkov Polytechnic Institute, National Technical University, Kharkov, Ukraine

b

OptimEnergo Energy Service Firm, Kharkov, Ukraineemail: [email protected]

Received June 15, 2009

Abstract

—A pinch analysis of the AVDU A12/2 crude distillation unit processing 2 million tons of crude oilper year is made. A pinch retrofit of the recuperative heat-exchange system of the unit is performed. It is shownthat the implementation of the proposed project will reduce the arbitrary energy consumption by 60%, whichcorresponds to a reduction in specific fuel consumption from 36 to 13 kg per ton of refined oil.

DOI:

10.1134/S0040579509060086

THEORETICAL FOUNDATIONS OF CHEMICAL ENGINEERING

Vol. 43

No. 6

2009

THERMAL PROCESS INTEGRATION IN THE AVDU A12/2 CRUDE 907

utilities by 9 and 12%, respectively. However, in that work,the decrease in the number of external utilities fed to thedistillation process was only determined without analyzingtheir production, i.e., without considering the heating ofcrude oil and its products in a pipe still; moreover, the ther-mal process integration covered only some of the processflows in the unit.

It was also shown [4] that energy saving due to mea-sures designed by pinch analysis in three European oilrefineries of a small scale of ~20 000–50 000 t of crude oilper year is $13.7 million per year. Using pinch analysismethods in the Guangdong Oil Refinery in China, $16 mil-lion was saved per year owing to an increase in the energyefficiency of the plant, making this oil refinery one of theworld’s best oil refineries (3%) [5].

The retrofitting of an oil refinery in Indonesia by pinchanalysis led to a reduction in energy consumption by 39%with a 6-month return on investment [6].

It was demonstrated [7, 8] that using pinch analysismethods in the design of recuperative heat-exchange sys-tems for crude distillation units can reduce the consump-tion of hot and cold utilities by 29 and 34%, respectively.

In this work, we perform a pinch analysis of the systemof process flows in one of the early crude distillation units,the A12/2 atmospheric/vacuum distillation unit (AVDU),designed in 1956 and upgraded for refining 2 million tonsof oil per year in 1967. Here, we make a pinch analysis ofonly one of the operating modes of the unit, namely, themode of winter operation of A12/2 with a vacuum section.

PINCH ANALYSIS AND INTEGRATIONThis plant currently uses a countercurrent pattern of

heat recuperation, which is typical of the 1970s and is gen-erally not optimal for large systems of heat exchangers [3].

We analyzed four operating modes of the unit: opera-tion with and without a vacuum section in winter and sum-mer. Figure 1 presents a flow sheet of the A12/2 unit with avacuum section.

A flow of crude oil at ambient temperature is fed to theunit where it is split into two flows, each being heated in aseries of heat exchangers and delivered through a mixer toa first-stage electric dehydrator (Fig. 1). Then, partiallydesalted oil is conveyed to a second-stage electric dehydra-tor. After the dehydrators, desalted dehydrated oil enterscolumns C1 and C1a in two flows, in each of which a partof the oil flow is heated in a pipe still and the other part isheated by the separation products in shell-and-tube heatexchangers.

Stripped oil from the bottom of columns C1 and C1aat a temperature of 200–250

°

C is fed to the coils of pipestills PS1 and PS2. A tail fraction of gasoline and dieselfuel in the vapor phase from the top of columns C2 andC2a is transferred through vapor lines to column C3.

A fraction of distillate oil and diesel fuel enters thetop section of stripping column C5/1, from which theflow is delivered through circulation reflux heatexchangers to columns C2 and C2a.

Vapors from the stripping section of column C5/1 entertray 16 of column C3. From trays 11 and 13 of column C2and trays 9, 11, and 13 of column C2a, atmospheric gasoilcan be removed; it is fed to the stripping section of strip-ping column C5/3, from which it is taken by a vapor pumpand pumped into diesel fuel or fuel oil. Vapors from thestripping section of stripping column C5/3 enter columnsC2 and C2a.

The tail gasoline fraction is distilled from the top of col-umn C3. Gasoline vapors enter through the vapor lines tocondensers and gas separators, and then gasoline flows bygravity to a buffer vessel.

Diesel fuel from the bottom of column C3 is pumpedthrough heat exchangers to a pool. If there is a vacuum sec-tion, part of the fuel oil from the bottom of columns C2 andC2a is pumped through the vacuum coil of pipe still PS1,where it is heated to a temperature of ~380–420

°

C, to vac-uum column C4. Fuel oil from the bottom of column C2can be directed to flow by gravity through a bypass line tovacuum column C4.

In vacuum column C4, a wide fraction—vacuum gas-oil—is distilled from fuel oil. Vacuum gasoil is thenpumped through heat exchangers to a pool.

Vacuum gasoil can be delivered to fuel oil. Water vaporand a part of the oil gases from the top of the vacuumreceiver enter a barometric condenser and are condensedusing vacuum gasoil.

The temperatures and flow rates of all the process flowsin the unit and the flows of the utility system were mea-sured with both stationary and mobile instruments. Theproperties of the flue gases of the utility system were mea-sured with gas analyzers and pyrometers. Data on the ther-mophysical properties of the products of the process flowsat the moments of time of the measurements were obtainedfrom the Central Plant Laboratory.

After classifying the material and heat balances of theunit on the basis of the collected data, the flow data neces-sary for integration were summarized in a flow table. Inmost of the process flows, there are either phase transitionsor the specific heats significantly change within the rangesof temperature variations; therefore, for further processing,the data were segmented according to their temperatureranges. In each of such intervals, the properties of a flowcan be considered as constant [9].

A comparison of the energy consumption by the unit inwinter and summer at the same production capacity withconsideration for the difference between the temperaturesof crude oil supplied from the storage tanks revealed signif-icant heat losses to the environment. Therefore, the temper-atures of all of the open surfaces of the transport system andthe heat-exchange equipment were measured. The mea-surement results allowed us to calculate the heat losses tothe atmosphere by convection and radiation. These lossesamounted to 8 MW, from which only the losses from theheat-exchange system constituted about 6 MW.

The power released from the air and water coolersand the heat loss power withdrawn with the flue gaseswere measured, as were the temperatures of the heat-

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TOVAZHNYANSKII et al.

23

°

C

CW

8

8

88

To

the

Gas

oil

¸

SC-1

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pool

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to p

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7a,7

b,7

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°

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339

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C-1


Vol. 43

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2009


transfer media in each working heat exchanger. Thisallowed us to calculate the recuperation power in theheat-exchange system in terms of both hot and coldflows. The two values differ by the heat losses from theheat-exchange system of the unit.

The parameters of the flows in the pipe stills, whichwere necessary for calculating the payload in the utilitysystem, were measured. The calculation showed thatthe payload in the pipe still was ~50 MW.

The formulation of the material and heat balances ofthe unit and the extraction of the data on the processflows allowed us to construct composite curves [9, 10]of the process flows of the unit (Fig. 2). The relativepositions of the composite curves for the hot and coldutilities were determined using both the utility mea-surement data (the power consumed by the processfrom the hot utilities and the power withdrawn from theprocess by the cold utilities), and the results of measur-ing the temperatures and flow rates of the heat-transfermedia in the heat-transfer equipment.

The composite curves contain much information onthe system of process flows, the utility system, and theefficiency of using heat in the process.

The projection of the composite curve for the hotutilities to the enthalpy (abscissa) axis presents the heatpower that can be withdrawn from the system of hotflows and can be used to heat the cold process flows.This value is 59.09 MW. Similarly, the projection of thecomposite curve for the cold utilities to the enthalpyaxis shows the heat power that should be supplied to thecold flows to perform crude distillation in the AVDUA12/2 unit.

Performing the process requires a power of~79.8 MW, but this means that all of this power must beobtained from the utility system of the unit, in our case,from the system of pipe stills. Part of the energy may berecuperated using the system of heat exchange betweenthe cold and hot flows. The power is shown by the spanbetween the curves (Fig. 2). The recuperation power inthe heat-exchange system is

Q

REC

≈

30 MW.The heat power consumed by the process from the

utilities is

Q

H

min

~ 49

.8 MW, and the load on the coldutilities is

Q

C

min

~ 28

.1 MW.The pinch is located at temperatures of hot and cold

flows of

t

Phot

= 210°

C and

t

Pcold

= 50°

C

,

respectively,and

Δ

T

min

= 160°

C.

Δ

t

min

in this case is not the minimal dif-

2

100

4 6 8

200

300

400

500

HE

1

HE

2

HE

4

HE

13T18T

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33T32

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28

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34

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14

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20

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6

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21

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24

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24

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1

PS

1

PS

1

PS1PS1

PS2PS2

PS2

1

2

Pinch

QCmin

QREC

H × 10–4, kW0

t, °C

QHmin

St

Fig. 2. Composite curves for crude distillation in the AVDU A12/2 unit with the current heat recuperation system (QHmin =49.8 MW, QCmin = 28.1 MW, QREC ≈ 30 MW): (1) composite curve for hot flows and (2) composite curve for cold flows (St, steamheater; PS1, PS2, pipe stills; HE1–HE34, heat exchangers).

Fig. 1. Flow sheet of the AVDU A12/2 crude distillation unit with the current heat recuperation system. The numbers within thecolumns are the numbers of trays. The numbers of flows are taken according to the table. The temperatures in the columns are thetemperatures of the bottom and top trays, respectively. A, BU, B, E, BE, vessels; ACC, air-cooled condenser; AVDU, atmo-spheric/vacuum distillation unit; BC, barometric condenser; C1, C1a, atmospheric distillation columns; C2, C2a, columns for thedistillation of stripped oil; C3, distillation column for the production of diesel oil; C4, vacuum column; C5, stripping column; CC,cooler–condenser; CO, cooler; CW, cooling water; DF, diesel fuel; ED, electric dehydrator; EDU, electric desalting unit; GS, gasseparator; HE1–HE34, heat exchangers; P, pump; PS, pipe still; PW, process water; SC, submerged cooler; VES, vapor ejectionsection; VP, vapor pump.

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THEORETICAL FOUNDATIONS OF CHEMICAL ENGINEERING Vol. 43 No. 6 2009


Properties of process flows involved in process integration for the winter operation of AVDU A12/2 with a vacuum section

No. Flow Type tS, °C

tT, °C

G, t/h

C, kJ/(kg °C)

r, kJ/kg

CPS, kW/°C

ΔH, kW

α, kW/(m2 °C)

1 Crossflow from C1, C1a to C3 A* 173 54 8.20 2.11 – 4.81 571.93 0.402.1 Gasoline vapor cooling in C1, C1a A 145 50 29.69 2.52 – 20.78 1974.36 0.102.2 Cooling of gas from C1, C1a A 145 40 2.78 2.52 – 1.95 204.33 0.102.3 Gasoline vapor condensation in C1, C1a A 50 50 29.69 – 75 – 618.53 12.4 Cooling of gas from C1, C1a A 50 40 29.69 2.11 – 17.40 174.01 0.403.1 Gasoline vapor cooling in C3 A 157 46 17.40 2.52 – 12.18 1351.98 0.103.2 Gasoline vapor condensation in C3 A 46 46 17.40 – 75 – 362.50 13.3 Water vapor cooling in C3 A 157 100 1.05 2.00 – 0.58 32.96 0.123.4 Water vapor condensation in C3 A 100 100 1.05 – 2256 – 658.00 23.5 Water cooling in C3 A 100 46 1.05 4.19 – 1.22 65.99 0.804 Diesel fuel A 229 200 71.11 2.70 – 68.29 1980.27 0.20

A 200 100 71.11 2.42 – 58.81 5880.67 0.20A 100 58 71.11 2.10 – 51.76 2174.05 0.20

5 Central reflux in C3 A 165 74 75.79 2.11 – 55.87 5083.79 0.156 Central reflux in C2, C2a A 295 144 30.73 2.11 – 29.85 4507.90 0.157 Fuel oil A 360 300 79.79 2.90 – 89.94 5396.67 0.15

A 300 250 79.79 2.71 – 81.41 4070.55 0.15A 250 200 79.79 2.54 – 73.74 3687.10 0.15A 200 150 79.79 2.36 – 65.87 3293.68 0.15A 150 90 79.79 2.17 – 57.37 3442.40 0.15

8.1 Vacuum gasoil A 330 300 43.96 2.52 – 30.77 923.16 0.408.2 Vacuum gasoil vapor condensation A 300 300 43.96 – 75 – 915.83 18.3 Vacuum gasoil cooling A 300 200 43.96 2.54 – 42.74 4273.89 0.15

A 200 150 43.96 2.36 – 36.29 1814.57 0.15A 150 68 43.96 2.17 – 30.66 2514.12 0.15

9 Tar A 336 250 9.33 2.70 – 9.63 828.14 0.1010 Salt solution from EDU A 115 40 15.05 4.21 – 17.14 1285.61 0.8011 Crude oil B** 10 50 250.85 1.95 – 142.55 5701.96 0.10

B 50 118 250.85 2.10 – 170.22 11574.78 0.1012 Desalted oil B 112 150 248.34 2.28 – 192.17 7302.34 0.15

B 150 200 248.34 2.47 – 214.40 10720.07 0.15B 200 214 248.34 2.65 – 231.36 3239.06 0.15

13 Stripped oil to right coil of PS1 B 214 250 56.00 2.75 – 54.87 1975.49 0.15B 250 300 56.00 2.84 – 59.75 2987.44 0.15B 300 371 56.00 3.18 – 66.73 4737.86 0.15

Stripped oil to left coil of PS1 B 214 250 26.00 2.75 – 25.48 917.19 0.15B 250 300 26.00 2.84 – 27.74 1387.03 0.15B 300 310 26.00 2.97 – 29.42 294.16 0.15

14 Stripped oil to left coil of PS1 B 214 250 64.00 2.75 – 62.71 2257.71 0.15B 250 300 64.00 2.84 – 68.28 3414.22 0.15B 300 370 64.00 3.18 – 76.17 5331.73 0.15

Stripped oil to right coil of PS2 B 214 250 64.00 2.75 – 62.71 2257.71 0.15B 250 300 64.00 2.84 – 68.28 3414.22 0.15B 300 370 64.00 3.18 – 76.17 5331.73 0.15

15 Fuel oil heating in PS1 B 360 407 53.29 2.90 – 65.68 3086.84 0.1516 Vacuum gasoil to HE24 B 15 87 43.96 2.10 – 25.64 1838.99 0.1517 Water to EDU B 10 80 12.54 4.20 – 14.60 1021.87 0.8018 Vapor overheating in PS1 B 135 450 1.05 2.69 – 2.60 819.88 0.6019 Gas to pipe stills (fuel) Å 31 80 3.00 2.52 – 2.10 102.90 0.1020 Fuel oil to pipe still (fuel) Å 49 80 3.79 2.05 – 2.32 71.80 0.10Notes: * A – hot flow.

** B – cold flow.



ference in temperature between the heat-transfer media inthe heat exchangers, because there is a crisscross heatexchange in the heat-exchange system (Fig. 2). Conversely,Δtmin represents the minimal temperature difference for thefully vertical heat exchange [9], for which the loads QHmin,QCmin, and QREC would reach the existing values. Theintense crisscross heat exchange results not only in a con-siderable increase in the heat-exchange surface area, butalso in significant heat transfer through the pinch, which iswell seen in the grid diagram of the existing network ofheat exchangers (Fig. 3).

In the existing system, 15 heat exchangers participate inthis transfer, and flow 15 (below the pinch) uses a hot utility,which is inadmissible in energy-efficient chemical engi-neering systems.

Even the first introduction with the processing systemrevealed a significant potential of energy saving.

The calculated load on the utility system is ~50 MW,and if we take into account the power withdrawn by fluegases, it is ~20.4 MW; then, we can estimate the powerreleased in fuel combustion in the pipe stills at ~70.4 MW,or ~36 kg of standard fuel per ton of refined oil.

Composite curves allow one to estimate the necessaryheat-exchange surface area for a designed or retrofitted pro-cess with sufficient accuracy. For this purpose, using themeasurement results obtained, we estimated the heat-trans-fer coefficients for each process flow (see table).

Further, using heat-exchange equipment prices reportedby manufacturers, we can estimate the necessary capitalinvestment and the period of the return on this investment.The cost of shell-and-tube heat exchangers is [8]

Capital cost = ÄHE + ÇHE(S)c,

where ÄHE = $5000; ÇHE = $120; and S is the heat-exchange surface area, m2. For shell-and-tube heatexchangers, as a rule, Ò = 0.87 [11].

The energy cost can be estimated from the market priceof fuel oil of $112 per ton at $87 per kilowatt per year.

Using these values and composite curves, diagrams canbe drawn for the determination of the dependence betweenΔtmin, the investment in an additional area to the existingheat-exchange surface area, and the period of the return oninvestment of the proposed project (Fig. 4), from which anadmissible Δtmin value can be found.

It can be seen that the problem of an energy-saving ret-rofit of the AVDU A12/2 unit is of a threshold nature, i.e.,once a certain investment value is exceeded, the annualsavings cease to increase (Fig. 4b). This threshold is at aninvestment of $2.23 million, and the period of the return oninvestment is 0.82 year. The minimal difference in temper-ature between the heat-transfer media in the heat exchang-ers near the threshold value of energy consumption is vir-tually equal to Δtmin = 20°ë.

The composite curves drawn for the considered operat-ing mode of the unit at Δtmin ≈ 20°ë (Fig. 5) show that coldutilities are virtually unused. The heat recuperation power

in the heat-exchange system is doubled, and the power ofthe hot utilities is lower by a factor of 2.3.

At the found optimal Δtmin value, using pinch analysismethods, we synthesized a grid diagram for the project ofretrofitting the heat-exchange system of the unit that lacksheat transfer through the pinch (Fig. 6). Constructedaccording to the grid diagram, the energy flow sheet of theretrofit project for the AVDU A12/2 unit (Fig. 7) has alower heat power consumed by the process (21.7 MW) anduses almost no cold utilities.

The heat-exchange surface area in the new heat-exchange system should be increased by 19.5 × 103 m2,with its current value being 12.6 × 103 m2. It is this thatrequires capital investment for retrofit.

For the optimal arrangement of utilities, the pinch anal-ysis uses the grand composite curve [9, 11], which indicatesthe energy requirements of the flow at any temperature fromthe temperature range characterizing the process flows.

Figure 8 presents the grand composite curve and thetemperature profile of flue gases of the pipe stills of theunit. The point of intersection of the temperature profileand the ordinate axis indicates the temperature at whichflue gases enter the chimney intake.

The power of the hot utilities shows the payload on thepipe stills, and the power of the cold utilities demonstratesthe heat power withdrawn from the process (Fig. 8). Theprojection of the temperature profile of flue gases on theabscissa axis presents the power released during fuel com-bustion in the pipe stills. In the case of the existing heat-exchange system of the process, the pipe stills release79.8 MW, which is equivalent to the combustion of 9.3 t ofstandard fuel per ton of refined fuel. After retrofitting theheat-exchange network of the unit according to the pro-posed project, e.g., for operation with a vacuum section inwinter, no cold utilities are used and the power of the hotutilities becomes 2.71 MW. Because of this, the powerreleased during fuel combustion in a pipe still is also lower,QHmin~ 32 MW, i.e., lower by a factor of 2.5, which corre-sponds to the combustion of 3.72 t of standard fuel per hour,or 15.5 kg of standard fuel per ton of refined oil (Fig. 8).Such a reduction in the fuel consumption by the pipe stillsallows one to abandon the use of fuel oil as a fuel for thepipe stills and to heat the products in a single pipe stillinstead of two.

The slope of the temperature profile of flue gases isdetermined by the fuel consumption of the pipe stills [9].The lower this consumption, the steeper the slope. Fortechnological reasons, the temperature of the flue gases inthe chimney intake upstream from the smoke stacks shouldnot be below 300°ë, which determines the minimum pos-sible fuel consumption by the pipe stills (Fig. 8). In thiscase, the power released in the pipe stills is ~27 MW, whichcorresponds to the consumption of 3.13 t of standard fuelper hour or 13 kg of standard fuel per ton of refined oil.

As the result of the retrofit, the specific fuel con-sumption by the process in the unit becomes ~15 kg ofstandard fuel per ton of refined oil, which is lower than

912



572

4.81

25.8

4

22.2

6

58.6

8

55.8

7

29.8

573

.67

39.8

69.

6317

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159.

97

208.

44

89.1

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07

65.6

825

.64

14.6

02.

60

72

2.10

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2971

2471

1003

5

5084

4508

1989

0

1044

282

812

86

1727

7

2126

1

1229

9

2200

7

3087

1839

1022

819

103

54°C

31°C

31°C

46°C

58°C

60°C

47°C

74°C

66°C

40°C

68°C

88°C

90°C

59°C

47°C

46°C

47°C

47°C62

°C

33°C

83°C

92°C

62°C

62°C

33°C

92°C

C C

C

C

CCCC

C

C

C

C

C

C C

AC

C3,

4C

O4

CC

3

ï8ï5

CC

1

CC

2SC5

SC1

SC7

AC

C5,

6

AC

C11

,12 AC

C9,

10 AC

C7,

8

AC

C1,

2

SC19SC

16A

CC

1311

6°C

SC18

SC18

‡

150í

125°

C14

5°C

130°

C

158°

C

155°

C

115°

C

122°

C

1

2 3

10

5

157°

C

145°

C

173°

C

138°

C

165°

CT

1T

2

T3

T4

T5

T6

T7

T19 T33

T13

T34

T3217

5°C

210°

C

200°

C

200°

C

200°

C13

5°C

135°

C

27°C

45°C 26

°C

10°C

120°

C17

7°C

T25

T28

11

144°

C

T14 T9

T20

T3

T30

PS

12229°

C

250°

C

225°

C

220°

C

112°

C

295°

C

243°

C

240°

C27

5°C

300°

C

T16

T17

T26

T21

T1

T22

T27

13

6

4

56°C

80°C 55

°C

50°C

40°C

49°C

31°C

23°C

14°C

16

17

19

20

18

14

ç

214°

C

135°

C

120 °

C

214°

C

è1

è1 è2

è2

è2è1

è1

T11

T12

T24

360°

C7

300°

C 360°

C1537

0°C

370°

C

310°

C

371°

C20

7°C

173°

C

214°

C21

0°C

205°

C

214°

C

132°

C15

0°C

166°

C18

6°C

450°

C

166°

C

407°

C

336°

C

153°

C14

6°C

203°

C

133°

C

330°

C

è1

8 9

130°

C12

8°C 1

24°C

120°

C11

4°C

53°C

123°

C11

8°C

93°C

72°C

86°C

Cps

, kW

/°ë

DH

,kW

T1

ç

C C

C

C

C

81°C

100°

CHE

2

HE

4

HE

1

PS1

HE

1H

E19

HE

32H

E33

HE

28

HE

34

PS2

HE

5H

E6

HE

3

HE

3

HE

20

HE

14

HE

17H

E16

HE

21

HE

26H

E1

HE

7

PS1

PS1

PS2

PS1

HE

22

HE

27

HE

12

HE

24

HE

11

HE

9



the current value by a factor of 2.4, and if the reductionin emissions with flue gases and the decrease in theirtemperature downstream from the chimney intake are

taken into account, then the specific fuel consumptionfor refining 1 ton of crude oil in the unit is ~13 kg ofstandard fuel.

1

0.5

2

1,0

0 I × 10–6,

1

1

2

0 I × 10–6, USD

2

1

0 40 80 Δtmin, °C

E × 10–6, USDE × 10–6, USD

P, year (‡)

(b) (c)

USD

Fig. 4. Determination of the optimum Δtmin value according to the chosen criterion of retrofitting the heat-exchange system:(a) retrofit project return on investment period versus retrofit investment, (b) economic efficiency, and (c) annual profit versus Δtmin.The dashed line shows the relationship between the minimal temperature difference in the heat-exchange system, annual profit,project investment, and return on investment of the project.

100

1 20 3 4 5 6 7

200

300

400

QREC

QHmin

H × 10–4, ÍW

t, °C

1

2

Δtmin = 20°C

Fig. 5. Composite curves for crude distillation in the AVDU A12/2 unit for Δtmin = 20°C (QHmin = 21.7 MW, QREC = 58.1 MW):(1) composite curve for hot flows and (2) composite curve for cold flows.

Fig. 3. Grid diagram for the existing heat-exchange system in the AVDU A12/2 unit, shown with the separation of process flows atthe pinch. The dashed line represents the location of the pinch; H, vapor heater; ΔH, flow heat change, W; C, cooling; ACC, air-cooled condenser; PS1, PS2, existing pipe stills; SC, submerged cooler; HE1–HE34, heat exchangers; CO, cooler; CC, cooler–con-denser.

914



Fig

. 6. G

rid

diag

ram

for

the

proj

ect o

f re

trofi

tting

the

heat

-exc

hang

e sy

stem

in th

e A

VD

U A

12/2

uni

t. T

he n

umbe

rs a

t the

poi

nts

of th

e sp

littin

g of

the

flow

s ar

e th

eir

perc

enta

gera

tios.

The

das

hed

line

repr

esen

ts th

e lo

catio

n of

the

pinc

h. T

he n

umbe

rs u

nder

the

heat

exc

hang

ers

are

thei

r he

at lo

ads

in k

W. H

E1–

HE

34 a

re h

eat e

xcha

nger

s.

572

4.81

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4

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6

58.6

8

55.8

7

29.8

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69.

63

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4

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97

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44

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2

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1.52

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1989

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8

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60

Cps

, kW

/°ë

DH

,kW

T2

T9

T1

15°C

46°C

88°C 40

°C

10°C

100°

C

74°C

58°C

90°C

144°

C

46°C

10°C

31°C

70°C

112°

C

87°C

80°C

135°

C10

22

115°

C

118°

C

112°

C

112°

C

124°

C

124°

C

124°

C

60

1267

572

1457

4421

2971

4056

1962

2410

112°

C

132°

C

153°

C

157°

C

68°C

150°

C

132°

C

132°

C

3745

2350 22

35

776

663

1278

1940

5689

2585

4671

9441

1791

229°

C

250°

C235°

C

124°

C

214°

C

235°

C

250°

C

214°

C

215°

C

217°

C

215°

C

209°

C

828

T3

T4

T6

T5

T7

T8

360°

C

278°

C35

2°C

360°

C

370°

C

407°

C

450°

C

330°

C

336°

C

295°

C

275°

C

292°

C

292°

C

292°

C

5891

1173

8

3087 81

9

3

4

5

10 12

18

11

16

17

19

31.1

%

68.9

%

32.1

8%

30.6

1%

19.7

1%

17.5

0%

14.8

%

12%

24.9

%

16.3

%

24.2

% 8%

29.1

%

13.7

%

39%

18.2

%

12.8

%

6

25.7

%

56.4

%

5.1%

CP

= 6

3.47

292°

C

CP

= 2

9.85

CP

= 8

5.02

CP

= 3

9.77

13.1

4

7

8

9

15

H H

H H

T16

T13

HE

17

T15

T10

165°

C

T12

T11

T14

T21

T20

T23

HE

22H

E22H

E21

HE

20

HE

15

HE

14

HE

11

HE

10

HE

12 HE

13

HE

23

HE

17

HE

1

HE

2

HE

4HE

3H

E9

HE

16

HE

5

HE

6

HE

7

HE

8



Fig

. 7. E

nerg

y flo

w s

heet

for

the

proj

ect o

f re

trofi

tting

the

heat

-exc

hang

e sy

stem

in th

e A

VD

U A

12/2

uni

t. T

he n

otat

ion

is a

s in

Fig

. 1.

9

3

8

15P-

5,6,

7

16

336°

C

C-4 P-

12,1

4

331°

C

P-10

,11

VP-

6,6‡

VP-

7a,7

·,7‚

E-8

BC

Gas

oil

to p

ool

3 st

2

st

1 st

To

atm

osph

ere

To

afte

rbur

ning

VE

SV

apor

181°

C

Bar

omet

ric

trap

2 G

asol

ine

C-1

from

uni

t

E-2

1Wat

erSa

ltso

luti

on

5

T-2

2

HE

-9H

E-1

0H

E-1

1H

E-1

2H

E-1

3

HE

-6H

E-7

P-1,

2,8

4

3

To

afte

r-bu

rnin

g

P-26

,27

Gas

olin

efr

om u

nit

BE

-2

E-1

0

VP-

25

VP-

0B

E-1G

S-1

55°C

GS-

246

°C

156°

C

Gas

to T

-23

A-5

5

5

217°

CC-39111516192123 4

HE

-19

Vap

or

1 R

eflu

x b

y s

trip

ped

oil

362°

C

C-2

‡9112123 1213

47 1V

apor

6

7

15

VP-

8

ç-3

ç-4

C-5

/1

C-5

/2

C-5

/3

C-2

9112123 1213

47 1

358°

C

Ref

lux

by

oil

str

ippe

d

Vap

or

VP-

5,5a

,13

P-16

,17,

17a,

17b

287°

C

P-24

P-23

11

15

P-20

P-18

P-19

P-20

a

287°

C

2

2

23 47

111315

C-1

‡C

-1

215°

C

23 47

111315

212°

C

148°

C14

2°C

12H

E-8

HE

-5

HE

-1H

E-2

HE

-3H

E-4

HE

-23

PS-1

PS-2

B-2

B-3

13

2.2

11

1210

118°

CE

D-1

ED

-211

2°C

80°C11

8°C

14

17

HE

-21

E-2

1

HE

-20

HE

-18

HE

-17

HE

-16

HE

-15

HE

-14

Tar

from

uni

t

Fue

l oil

from

uni

t

DF

from

uni

t Gas

from

E10

Cru

deoi

l

Wat

er

10°C

11

916



CONCLUSIONS

Thus, after retrofitting the heat-exchange system forimproving the thermal process integration in the systemof process flows in the crude distillation unit, no morethan 13 kg of standard fuel will be consumed per ton ofrefined oil, which is 3 times lower than that consumedcurrently. The period of the return on investment of theretrofit project will not exceed 8 months.

NOTATION

Aí—cost of the installation of a heat exchanger,USD;

Bí—coefficient that is equivalent to the cost of 1 m2

of heat-exchange surface area, USD;C—specific heat, J/(kg°ë);Ò—coefficient that characterizes the nonlinear

dependence of the cost of a heat exchanger on its heat-exchange surface area;

CPS—flow heat capacity (water equivalent), W/°ë;E—annual savings, USD;G—mass flow rate, kg/s;H—flow heat, W;ΔH—flow heat change, W;I—investment, USD;P—payback period, year;QCmin—power of cold utilities, W;QHmin—power of hot utilities, W;QREC—power of heat energy recuperation, W;Qstill—power released in a pipe still, W;r—latent heat of phase transition, J/kg;

S—heat-exchange surface area, m2;St—steam heater;t—temperature, °ë;tcold—temperature of pinch of cold flows, °C;tPhot—temperature of pinch of hot flows, °ë;tS—initial flow temperature (supply temperature),

°ë;tT—final flow temperature (target temperature), °ë;Δtmin—minimal difference in temperature between

heat-transfer media in recuperative heat exchangers,°ë;

α—characteristic heat-transfer coefficient,W/(m2°ë);

A, BU, B, E, BE—vessels;ACC—air-cooled condenser;AVDU—atmospheric/vacuum distillation unit;BC—barometric condenser;C1, C1a—atmospheric distillation columns;C2, C2a—columns for the distillation of stripped

oil;C3—distillation column for the production of diesel

oil;C4—vacuum column;C5—stripping column;CC—cooler–condenser;CO—cooler;CW—cooling water;DF—diesel fuel;ED—electric dehydrator;EDU—electric desalting unit;

0

1000

800

600

400

200

0

1 2 3123

1

2

Qfurnace

QHmin

Qëmin

H × 10–4, kW

t, °C1000

800

600

400

200

1200

t, °C

1 10 QHmin

Qfurnace

Q'furnace

H × 10–4,kW

1

23

(‡) (b)

Fig. 8. Determination of the fuel consumption of the pipe stills by the pinch analysis: (a) for the existing heat-exchange system inthe AVDU A12/2 unit and (b) for the proposed retrofit project; (1) grand composite curve for crude distillation, (2) temperature pro-file for flue gases from the pipe stills of the unit, and (3) temperature profile for flue gases from the pipe stills of the unit after reduc-ing the temperature in the chimney intake.



GS—gas separator;H—heating of flows in Figs. 3 and 6;HE1–HE34—heat exchangers;P—pump;PS—pipe still;PW—process water;SC—submerged cooler;VES—vapor ejection section;VP—vapor pump.

SUBSCRIPTS AND SUPERSCRIPTS

Cmin—power of cold utilities that is required for theprocess;

Hmin—power of hot utilities that is required for theprocess;

HE—heat exchanger;Pcold—temperature of pinch of cold flows;Phot—temperature of pinch of hot flows;PS–—flow;REC–—recuperation;S—initial flow temperature;Still—power released during fuel combustion in a

pipe still;T—target flow temperature.

REFERENCES

1. Stepanov, A.V., Sul’zhik, N.I., and Goryunov, V.S., Rat-sional’noe ispol’zovanie syr’evykh i energeticheskikhresursov pri pererabotke uglevodorodov (Rational Use

of Raw and Energy Resources in Hydrocarbons Process-ing), Kiev: Tekhnika, 1989.

2. Stepanov, A.V. and Goryunov, V.S., Resursosberegay-ushchaya tekhnologiya pererabotki nefti (Resource-Sav-ing Technology of Oil Refining), Kiev: Naukova Dumka,1993.

3. Klemes, J., Kostenko, Yu.T., Tovazhnyanskii, L.L.,Kapustenko, P.A., Ul’ev, L.M., Perevertailenko, A.Yu.,and Zulin, B.D., The Pinch Design Method for Energy-Saving Oil-Refining Plants, Teor. Osn. Khim. Tekhnol.,1999, vol. 33, no. 4, p. 420 [Theor. Found. Chem. Eng.(Engl. Transl.), vol. 33, no. 4, p. 379].

4. Ricci, G. and Bealing, C., Using an Integrated ApproachConserves Energy and Hydrogen, Neftegaz. Tekhnol.,2004, no. 3, p. 83.

5. Mallik, S. and Dhoul, V., Integrated Development andPlant Design, Neftegaz. Tekhnol., 2008, no. 4, p. 95.

6. Azimi, A.S., Klemes, J., and Ngatirin, S., Optimizationof Heat–Exchange System for Crude Oil Rectification,Integr. Tekhnol. Energosberezh., 2003, no. 3, p. 97.

7. Promvitak, P., Siemanond, K., Bunluesriruang, S.,Ragharentai, V., Retrofit Design of Heat Exchanger Net-works of Crude Distillation Unit, Chem. Eng. Trans,2009, Part 1, vol. 18, p. 99.

8. Promvitak, P., Siemanond, K., Bunluesriruang, S., andRagharentai, V., Grassroots Design of Heat ExchangerNetworks of Crude Distillation Unit, Chem. Eng. Trans.,2009, Part 1, vol. 18, p. 219.

9. Smith, R., Klemes, J., Tovazhnyanskii, L.L., Kapuste-nko, P.A., and Ul’ev, L.M., Osnovy integratsii teplovykhprotsessov (Foundations of Heat Processes Integration),Kharkov: Izdatel’skii Tsentr NTU “KhPI”, 2000.

10. Smith, R., Chemical Process Design and Integration,Chichester: Willey, 2005.

11. Linnhoff, B., Townsend, D.W., Boland, D., et al., UserGuide on Process Integration for the Efficient Use ofEnergy, Rugby: IChemE, 1991.

Documents

Thermal Process Integration in the AVDU A12/2 Crude