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Volume 6, Issue 2 2011 Article 3 Chemical Product and Process Modeling CAPE FORUM 2011 Modelling and Optimization of Crude Oil Hydrotreating Process in Trickle Bed Reactor: Energy Consumption and Recovery Issues Aysar T. Jarullah, University of Bradford Iqbal M. Mujtaba, University of Bradford Alastair S. Wood, University of Bradford Recommended Citation: Jarullah, Aysar T.; Mujtaba, Iqbal M.; and Wood, Alastair S. (2011) "Modelling and Optimization of Crude Oil Hydrotreating Process in Trickle Bed Reactor: Energy Consumption and Recovery Issues," Chemical Product and Process Modeling: Vol. 6: Iss. 2, Article 3. DOI: 10.2202/1934-2659.1600 Brought to you by | Hemeroteca de Estudios Económicos - Banco de la República Authenticated Download Date | 4/19/15 6:11 PM
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Volume 6, Issue 2 2011 Article 3

Chemical Product and ProcessModeling

CAPE FORUM 2011

Modelling and Optimization of Crude OilHydrotreating Process in Trickle Bed Reactor:

Energy Consumption and Recovery Issues

Aysar T. Jarullah, University of BradfordIqbal M. Mujtaba, University of BradfordAlastair S. Wood, University of Bradford

Recommended Citation:Jarullah, Aysar T.; Mujtaba, Iqbal M.; and Wood, Alastair S. (2011) "Modelling andOptimization of Crude Oil Hydrotreating Process in Trickle Bed Reactor: Energy Consumptionand Recovery Issues," Chemical Product and Process Modeling: Vol. 6: Iss. 2, Article 3.DOI: 10.2202/1934-2659.1600

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Modelling and Optimization of Crude OilHydrotreating Process in Trickle Bed Reactor:

Energy Consumption and Recovery IssuesAysar T. Jarullah, Iqbal M. Mujtaba, and Alastair S. Wood

Abstract

Energy consumption is a very important consideration for reducing environmental impact andmaximizing the profitability of operations. Since high temperatures are employed in hydrotreating(HDT) processes, hot effluents can be used to heat other cold process streams. The aim of thepresent paper is to describe and analyze the heat integration (during hydrotreating of crude oil intrickle bed reactor) of a hydrotreating plant’s process based upon experimental work.

In this work, crude oil is hydrotreated upon a commercial cobalt-molybdenum on aluminacatalyst presulfided at specified conditions. Detailed pilot plant experiments are conducted in acontinuous flow isothermal trickle bed reactor (TBR) in which the main hydrotreating reactions,are hydrodesulfurization (HDS), hydrodenitrogenation (HDN), hydrodeasphaltenization (HDAs)and hydrodemetallization (HDM). The latter includes hydrodevanadization (HDV) andhydrodenickelation (HDNi). The reaction temperature, the hydrogen pressure, and the liquidhourly space velocity (LHSV) are varied within certain ranges, with constant hydrogen to oil ratio(H2/Oil).

Experimental information obtained from a pilot plant, together with kinetics and reactormodeling tools,and a commercial process data are employed for heat integration process model.The optimization problem to minimize the overall annual cost is formulated as a Non-LinearProgramming (NLP) problem, which is solved using Successive Quadratic Programming (SQP)within gPROMS.

KEYWORDS: hydrotreating, trickle-bed reactor, integrated process, energy recovery

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1. Introduction

One of the tasks with which chemical engineers are continually addressing included is the scale up of laboratory experiments to full-scale production. Due to the high cost of a pilot plant study, this step is starting to be surpassed in several instances via designing the full scale unit based on the process of a bench scale plant called a microplant. In order to make this jump successfully, a thorough understanding of the chemical kinetics and transport limitations is required. However, energy conservation is significant in operation design. The estimation of the minimum cooling and heating requirements reveal important energy savings. For instance, Union Carbide in the USA and Imperial Chemical Industries in the UK have both reported the results of numerous case investigations that reference 30% to 50% energy savings in comparison to conventional practice (Douglas, 1988).

Concentrations of carbon dioxide (CO2) in the atmosphere have increased from 270 ppm before the industrial age to 380 ppm by 2006: a 31% increase since 1870 and a 41% increase over preindustrial values. The primary human source of carbon dioxide in the atmosphere is from the burning of fossil fuels towards energy production and transport. To avoid or reduce global warming, dramatic cuts in all carbon dioxide emissions must be achieved, 25% to 40% below 1990 levels by 2020, and 80% to 95% below 1990 levels by 2050 (www.greenpeace.org). However, more efficient utilization of energy consumption results in the reduction of the negative effects of carbon dioxide emissions. Thus, process integration is an efficient design methodology, which addresses cases related to energy efficiency, waste minimization and an efficient utilization of raw materials.

Operation units, oil refineries, petrochemical complexes and gas units all generate big amounts of low grade heat. This energy is usually rejected to the atmosphere utilizing either air or cooling water systems. However, there are opportunities for recovering some of this energy, and using it either as part of an operation integration scheme or to heat in domestic and industrial specifications by the installation of a hot water system. Energy integration is a very beneficial tool and is a significant phase in estimating the cost of preliminary design, where recovery of waste heat provides both financial and environmental benefits to process unit operators. From an energy saving point view the significant field of energy use improvement are the heat exchange retrofit projects for maximizing the existing heat recovery (Khalfalla, 2009).

Traditional design methods begin by designing the reactor, the separation system, the heat exchanger and finally end by utilizing utilities for supplying residual needs (Douglas, 1988). The utility includes hot and cold utility units. Typically, hot utility units are furnaces, turbines, generators, boilers and motors

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providing the necessary power, hot water and steam. Cold water from external sources is employed as the cold utility, which provides the required cooling in the operations. In the recovery system, the process streams exchange heat so as to reduce the cold and hot utility requirements. The heat exchangers are the only units in a heat recovery system (Khalfalla, 2009).

The main focus of this paper is to maximize heat recovery within the crude oil hydrotreating process in a trickle bed reactor based, upon experimental work and analysis the heat integration process.

2. Experimental Work

Briefly, hydrotreating experiments were conducted in an isothermal pilot-plant continuous flow trickle bed reactor (TBR). The reactor tube was made from stainless steel with an inside diameter of 2 cm and a length of 65 cm. The temperature of the reactor was maintained at the desired value by utilizing independent temperature control of five zone electric furnaces that provide an isothermal temperature along the active reactor section. A commercial cobalt-molybdenum on alumina (Co-Mo/γ-Al2O3) catalyst was used for all experiments. 60.3 g of the fresh catalyst was charged to the HDT reactor and in situ activated by a solution of 0.6 vol% of CS2 in commercial gas oil. Iraqi crude oil was employed as a feed for the pilot plant hydrotreating studies. The main properties of the feedstock and the catalyst used in this work are shown in Tables 1and 2, respectively. The main hydrotreating reactions considered in this study are HDS, HDN, HDAs and HDM, which involves HDV and HDNi. The experimental studies were carried out by varying the temperature (335-400 ºC), the pressure 4-10 MPa and the liquid hourly space velocity (LHSV) 0.5-1.5 hr-1, maintaining constant H2/Oil ratio at 250 L/L. Further details of the experimental pilot plant, reactor, equipment and procedure, experimental runs, catalyst loading and catalyst presulfiding used in this study can be found in Jarullah et al. (2011a,b,c,d).

Table 1: Feedstock specifications Specific Gravity at 15.6 °C 0.8558 Viscosity at 37.8 °C 5.7 cSt Pour point -36 °C Sulfur content 2.0 wt % Nitrogen content 0.1 wt % Asphaltenes content 1.2 wt % Vanadium content 26.5 Nickel content 17 ppm

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Table 2: Catalyst commercial specifications (Co-Mo/-Al2O3 )

Chemical propertiesMoO3 15 wt % NiO 3 wt % SiO2 1.1 wt % Na2O 0.07 wt % Fe 0.04 wt % SO2 2 wt % Al2O3 Balance Physical properties Form Extrude Pore volume 0.5 cm3/g Surface area 180 m2/g Mean particle length 4 mm Bulk density 0.67 g/cm3 Mean particle diameter 1.8 mm

3. Energy Consumption and Recovery Issues

Pilot plant trickle bed reactor experiments show that conversion in hydrotreating reactions is better at high reaction temperatures. Energy consumption for the pilot plant scale was negligible and natural cooling after the reaction was sufficient (no additional utility was required as the amounts reactants and products were small at pilot plant scale), thus heat recovery was not taken into consideration. On the other hand, heat recovery was not an issue in the pilot plant scale process. In industrial processes, energy consumption will be a large issue and heat recovery must be taken into account, especially when the type of reactions are exothermic. Therefore, scaling up a heat integrated hydrotreating process was considered for reducing overall energy consumption (hence reducing environmental effect). However, this leads to the addition of a number of exchangers in the system, requiring capital investment. The objective is to calculate a retrofit design, which can reduce the energy consumption, maximize energy recovery and consequently minimize capital investment.

Generally heat exchangers operate in series with a heater and a cooler. The heater regulates the final temperature of the cold fluid to the required reaction temperature, and the cooler adjusts the final temperature of the hot fluid to requirements of the next step of the process. The exchangers, heaters and coolers used for heat integration and energy consumption in this study are shown in Figure 1.

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Figure 1: Process of heat integrated reaction system

As depicted in Figure 1, the crude oil feedstock (cold stream) is pumped by P1 before preheating from TC0 to TC1 in heat exchanger H.E1. Then, the crude oil is fed into furnace F1 in order to preheat from TC1 to the reaction temperature (TR). The second main feedstock, which is hydrogen (cold stream) is fed into heat exchanger H.E2 to preheat from TH0 to TH1. After this, its temperature rises from TH1 to the reaction temperature (TR) by the furnace F1. The product stream leaving the reactor (hot stream) is cooled from TP1 to TP2 by contacting with the main crude oil feedstock in heat exchanger H.E1. Due to high reaction products

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temperature, the products stream is cooled again from TP2 to TP by in contact with hydrogen feed in heat exchanger H.E2, then its temperature is reduced from TP to the final product temperature (target temperature) TFP in cooler CO1 by using cold water at TW1.

4. Model Equations

The major focus of this paper is to minimize energy consumption and maximize heat recovery during crude oil hydrotreating. The main mass balance equations, energy balance and reaction rate equations for all processes (HDS, HDN, HDAs, HDV and HDN) are presented in Table 3. Other correlations for estimating gas and liquid properties and characteristics of the catalyst bed used at process conditions can be found in Jarullah et al. (2011b,c,d). Also, all the parameters required for the industrial trickle bed reactor are estimated based on the information presented in the literatures (Tarhan, 1983; Wauquier, 1995; Rodriguez and Ancheyta, 2004; Bhaskar et al., 2004; Mederos and Ancheyta, 2007; Alvarez et al., 2009; Chen et al., 2010).

a) Heat Exchanger (H.E1)

The products stream that leaves the reactor is used for preheating the crude oil feedstock from TC0 to TC1 through H.E1 and at the same time is cooled from TP1 to TP2. The heat duties for these streams are described as follows:

)( )( 0111 CCLLpLHE TTcVQ (1)

)( )( 1212 PPG

GpGLLpLHE TTcVcVQ (2)

1112 HEHEQ Q (3)

11

111

lmHE

HEHE TU

QA

(4)

2

1

211

lnT

T

TTTlm (5)

111 CPTT T (6)

022 CPTT T (7)

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Table 3: Mass balances and reaction rate equations (Jarullah et al., 2011b,c,d) Mass balance equations

Gases compounds (H2, H2S) in gas phase

L

ii

GiL

Lig

Gi C

h

Pak

u

RT

dz

dP

Gases compounds (H2, H2S) in liquid phase

S

iLi

SSi

Li

i

GiL

Lil

Li CCakC

h

Pak

udz

dC 1

Liquid compounds (S, N, Asph, V and Ni) in liquid phase S

iLi

SSi

L

Li CCak

udz

dC

1

H2 in solid phase (j =HDS, HDN, HDAs, HDV and HDNi)

jjBSH

LH

SSH rCCak

222

H2S in solid phase HDSHDSBS

SHL

SHSS

SH rCCak 222

S, N, Asph, V and Ni in solid phase (j =HDS, HDN, HDAs, HDV and HDNi)

jjBSi

Li

SSi rCCak

Energy balance equationa l

LpLL

GgpGg

ljBjR cucu

rHdz

dT

Chemical reaction rates

HDS 2

22

2

1 SSHSH

mSH

nSsul

HDSHDSCK

CCKr

HDN, HDAs, HDV and HDNi mjSH

njSijj CCKr

2

HDS , HDN, HDAs, HDV and HDNi

RT

EAAK j

jj exp0

H2S

RTK SH

2761exp841141769

2

a taken from Tarhan (1983)

Q1HE1 and Q2HE1 are heat duties of H.E1, TC0 is the inlet temperature of the cold fluid, TC1 is the outlet temperature of the cold fluid, TP1 is the inlet temperature of the hot products mixture, TP2 is the outlet temperature of the hot products mixture, VL is the volumetric flow rate of crude oil, VG is the volumetric flow rate of hydrogen (includes the main hydrogen feed plus quench feeds), ρL is the liquid density, ρG is the gas density (includes all reacting gases), cp

G is the specific heat capacity of gas (involves all gases (reactants and products)), cp

L is the specific heat capacity of liquid, AHE1 is the heat transfer area of H.E1, UHE1 is

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the overall heat transfer coefficient for H.E1 and ∆Tlm1 is the log mean temperature difference for H.E1.

b) Heat Exchanger (H.E2)

The main hydrogen feedstock is heated from TH0 to TH1 in H.E2 by contact with the products stream that leaves H.E1, which is cooled at the same time from TP2 to TP in H.E2. The model equations for H.E2 are:

)( )( 01212

2 HHGHpHHE TTcVQ (8)

)( )( 222 PPGGpGL

LpLHE TTcVcVQ (9)

2122 HEHEQ Q (10)

22

212

lmHE

HEHE TU

QA

(11)

4

3

432

lnT

T

TTTlm (12)

123 HPTT T (13)

04 HPTT T (14)

Q1HE2 and Q2HE2 are heat duties of H.E2, TH0 is the inlet temperature of the cold fluid, TH1 is the outlet temperature of the cold fluid, TP2 is the inlet temperature of the hot products mixture, TP is the outlet temperature of the hot products mixture, VH2 is the volumetric flow rate of hydrogen (without quench feeds), cp

H2 is the specific heat capacity of hydrogen, AHE2 is the heat transfer area of H.E2, UHE2 is the overall heat transfer coefficient for H.E2 and ∆Tlm2 is the log mean temperature difference for H.E2.

c) Cooler (CO1)

The product stream that leaves H.E2 is cooled from TP to the final product temperature TFP in the cooler by using water at TW1. The model equations for CO1 are shown below:

)( 1211 WWwWCO TTcpmQ (15)

)( )(12 PFPGGpGL

LpLCO TTcVcVQ (16)

1112 COCOQ Q (17)

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lmcCO

COCO TU

QA

1

111 (18)

6

5

65

lnT

T

TTTlmc (19)

25 WPTT T (20)

16 WFPTT T (21)

Q1CO1 and Q2 CO1 are heat duties of CO1, TW1 and TW2 are the inlet and outlet temperature of the cooled water, TFP is the final products temperature, ACO1

is the heat transfer area of CO1, UCO1 is the overall heat transfer coefficient for CO1, ∆Tlmc is the log mean temperature difference for CO1, cpw is the heat capacity of water and mw is the mass flow rate of cooling water.

The total heat transfer area (At) is given in the following equation:

121 COHEHEt AAAA (22)

d) Furnace (F1)

The main feedstocks, crude oil and hydrogen are fed into furnace F1 separately in order to preheat from TC1 to the reaction temperature (TR) for crude oil and from TH1 to TR for hydrogen. The heat equations for F1 can be written as follows:

)( )( 12

22 HRGHpHH TTcVQ (23)

)( )( 1CRLLpLCrude TTcVQ (24)

CrudeHF QQQ 2

(25)

Note, all the physical properties of gases and liquid (such as gas and liquid

density, heat capacity of gas and liquid, gas compressibility factor) are estimated at average temperature (Tav) for each equipment using the following equation:

2outin

av

TTT

(26)

Tin and Tout are the inlet and outlet temperatures for item of each equipment.

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5. Optimization Problem Formulation

The optimization problem can be stated as follows:

Given inlet temperature of crude oil (TC0) and hydrogen (TH0), outlet temperature of products mixture (TP1), reaction temperature (TR), inlet water temperature(TW1), volumetric flow rate of liquid (VL) and gas (VH2 and VG).

Optimize TP2, TH1, TW2

So as to minimize the total annual cost of the process (Ct).

Subject to process constraints and linear bounds on all decision variables.

Mathematically, the optimization problem can be presented as:

Min Ct TP2, TH1, TW2 s.t f(x(z), u(z), v) = 0 (model, equality constraints)

TP2L ≤ TP2 ≤ TP2

U (inequality constraints) TH1

L ≤ TH1 ≤ TH1U (inequality constraints)

TW2 L ≤ TW2 ≤ TW2

U (inequality constraints) ∆TF

L ≤ ∆TF ≤ ∆TFU (inequality constraints)

∆TWL ≤ ∆TW ≤ ∆TW

U (inequality constraints) TR = TR

* (equality constraints) TFP = TFP

* (equality constraints)

∆TW and ∆TF are the temperature differences between the inlet and outlet temperatures of water in the cooler and liquid crude oil in the furnace, respectively. TR

* is the required reaction temperature (which is 400°C) and TFP* is

the target final temperature of product (which is 26°C). Note, in practice, that the best temperature difference between inlet and outlet water in the cooler is within 5-20°C, and between inlet and outlet crude oil in the furnace within 100-130°C, which are quite practical. U and L are the upper and lower bounds.

5.1 Cost Function

The objective function is the overall annual process cost (Ct), which can be calculated using the following expression:

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Ct ($/yr) = Annualized Capital Cost ($/yr) + Operating Cost ($/yr) (27)

To calculate the annualized capital cost (ACC) from capital cost (CC), the following equation is used (Smith, 2005):

1)(1

)(1

n

n

i

iiCCACC (28)

n is number of years and i is the fractional interest per year; n = 10 years, i = 5% (Smith, 2005).

Capital Cost (CC, $) = Reactor Cost (CR) + Compressor Cost (CComp) + Heat Exchanger Cost (CHE) + Pump Cost (CP) + Furnace Cost (CF) (29)

The operating cost is calculated as shown below:

Operating Cost ($/yr) = Heating Cost (CH) + Compression Cost (CCmpr) + Pumping Cost (CPU) + Cooling Cost (CCol) (30)

The capital costs of equipment can be estimated using the following equations (Douglas, 1988; Smith, 2005; Quintero and Villamil, 2009):

a) Reactor Cost (CR)

)18.2(9.101280

&($) 802.0066.1

CRRR FLDSM

C

(31)

pmC FFF (32)

b) Compressor Cost (CComp)

)11.2())(5.517(280

&($) 82.0

dComp FbhpSM

C

(33)

ise

hpbhp

(34)

11003.3 5

in

outinin P

PQPhp (35)

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2

2

2

2

1

H

H

H

H

cv

cp

cv

cp

(36)

Rcpcv GG (37)

c) Heat Exchanger Cost (CHE)

)29.2(78.210280

&($) 65.0

CtHE FASM

C

(38)

mpdC FFFF )( (39)

d) Pump Cost (CP)

55.0

3

41084.9

280

&($)

P

CP

QF

SMC (40)

TpmC FFFF (41)

e) Furnace Cost (CF)

)27.1( 1052.5 280

&($) 85.03

CFF FQSM

C

(42)

pmdC FFFF (43)

M&S is the Marshall and Swift index for cost escalation (M&S = 1468.6 (chemical engineering, 2010)), bhp is the brake horsepower required in the compressor motor, hp is the compressor horsepower, ise is isentropic efficiency, γ

is the specific heat ratio, Qin is the volumetric flow rate at compressor suction, cvG is the specific heat capacity at constant volume and Pin and Pout are the pressure in the compressor inlet and outlet, respectively. ηise ranges between 70-90% (here it is assumed to be 80%) (Douglas, 1988). Qp is the pump power, QF is the heat duty of the furnace, LR is the reactor length, DR is the reactor diameter and FC, Fm, Fp, FT and Fd are the dimensionless factors that are functions of the construction material, operating pressure and temperature, and design type.

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The operating cost is calculated using the following expressions:

a) Heating Cost (CH)

yr

day

day

h

kWhkWQyrC FH 1

340

1

24$062.0)()/($ (44)

b) Compression Cost (CCmpr)

yr

day

day

h

kWhhp

kWhpbhpyrCCmpr 1

340

1

24$062.0

341.1

1

9.0

)()/($ (45)

c) Pumping Cost (CPU)

yr

day

day

h

kWhkWQyrC PPU 1

340

1

24$062.0)()/($ (46)

d) Cooling Cost (CCol)

yr

day

day

h

kg

f

h

kgmyrC w

wCol 1

340

1

24$)/($ (47)

fw is the price of cooling water, which is 0.0305 $ per ton (Khalfallah, 2009). The compression cost is based upon a motor efficiency of 90% (Bouton and Luyben, 2008) and an average power price of 0.062$/kWh (Alvarez et al., 2009).

6. Results and Discussion

The industrial trickle bed reactor (TBR) in this study has a processing capacity of 10000 bbl/day. The reactor is assumed to operate for 340 days/yr. In our previous study (Jarullah et al., 2011a), it has been found that the crude oil hydrotreatment process (from experimental data) achieved maximum conversion of the impurities (sulfur, nitrogen, vanadium, nickel and asphaltene) at reaction temperature 400 ºC, liquid hourly space velocity (LHSV) 0.5 hr-1 and hydrogen pressure 10 MPa. These conditions are utilized as a typical operating condition for the industrial trickle bed reactor (TBR).

The values of constant parameters with factors and coefficients used in this work (Douglas, 1988; Sinnott, 2005; Smith, 2005) are listed in Table 4.

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Table 4: Values of constant parameters, factors and coefficients used in this work Parameter Unit Value TC0 °C 28.00 TH0 °C 70.00 TP1 °C 406.02 TW1 °C 20.00 Cpw kJ/kg K 4.189 UHE1 W/m2 K 250 UHE2 W/m2 K 113 UCO1 W/m2 K 400

Dimensionless Factors Fm Fp Fd FT

Reactor 3.67 3.93 0.0 0.0 Compressor 0.0 1.0 0.0 0.0 Furnace 0.75 0.15 1.0 0.0 Pump 1.0 1.9 0.0 1.0 Heat Exchanger 3.75 0.625 1.0 0.0

6.1 Kinetic Parameters Estimation

The optimal set of kinetic parameters for crude oil hydrotreating reactions (involve HDS, HDN, HDAs, HDV and HDNi) have been evaluated based upon minimizing the sum of squared error between experimental and estimated results. These kinetic parameters were validated against pilot plant experiments reported elsewhere (Jarullah et al., 2011b,c,d). Such kinetic parameters (summarized in Table 5) used for HDS, HDAs, HDN, HDV and HDNi have been developed accurately among all results for all reactions with average absolute error less than 5%, which can be confidently applied for reactor design, operation and control.

Table 5: Kinetic parameters of hydroprocessing reactions (Jarullah et al., 2011b,c,d) Reactions Reaction

order (n) Hydrogen order (m)

Activation energy (EA) (J/mol)

Pre-exponential factor (A0) (mol/cm3)1-n.(cm3/g.sec). (mol/cm3)-m

HDS 1.147 0.4709 50264.10 2026.23 HDN 1.672 0.3555 71775.5 2.85×107 HDAs 1.452 0.3068 104481 2.56134×108 HDV 1.251 0.6337 46181.6 126566 HDNi 1.688 0.5667 37678.3 1.045×108

Note: these parameters are used in the model equations presented in Table 3.

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Table 6: Results of optimization problem for heat integration process Variables Without heat

integration With heat integration

Decision variable type

Optimized value

At (m2) 475.3148 1033.207 TP2 (C) 202.99

Ct ($/yr) 9744870.8 4310909.8 TH1 (C) 200.57 CS (%) ------- 55.76 TW2 (C) 40.00 mw (kg/hr) 711873.9 272716.56 TR (C) 400.00 Qt (kJ) 1386189676.8 496317300.5 ∆TF (C) 110.25 QC (kJ) 1431379037 548357040 ∆TW (C) 20.00 QHE1 (kJ) ------- 784918598.4 TFP (C) 26.00 QHE2 (kJ) ------- 104953777.9 ------- ------- Qr (kJ) ------- 889872376.3 ------- ------- ES (%) 0.0 64.2 ------- -------

CS = Cost saving, ES = Energy saving, Qr = Heat recovery, Qt = Total heating, QC = Heat duty of cooler, QHE = Heat duty of exchanger1,2.

6.2 Optimization Results The optimization results, which are summarized in Table 6, show that the minimum total cost (Ct) and cooling water amounts with heat integration of the hydrotreating process are less than those without the heat integration at specified variables. The cost saving is around 56% in comparison with the cost obtained without heat integration to achieve the reaction temperature (400°C) and to reduce the final product temperature up to 26°C. Also, the amount of cooling water required for reaching the final product temperature is greater than that with heat integration due to the heat recovery, and consequently the cost of the cold utility in addition to capital cost of cooler will decrease.

The results also show that the minimum energy requirement was reduced by 64.2%. Therefore, the CO2 emissions will be reduced by 64.2%, which has the added benefit of significantly reducing environmental impact.

7. Conclusions

Heat integration and energy consumption in a hydrotreating process were investigated. It has been observed that the energy consumption and heat recovery is considered a big issue that should be taken into account in industrial operations particularly when the type of reactions are exothermic, the recovery of which is very significant for maximizing profitability of process. Optimization problem was formulated in order to optimize some of the design and operating parameters of integrated process while minimizing an objective function, which is a coupled function of capital and operating costs including design parameters. Optimal

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Chemical Product and Process Modeling, Vol. 6 [2011], Iss. 2, Art. 3

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minimum energy requirements, heat recovery and cost saving were obtained utilizing optimization process. The results show that the cost saving are 55.76% and the energy saving is 64.2% in comparison to the process without heat integration. This saving in the energy consumption provides better minimum energy requirement, hence reducing environmental effect and maximum heat recovery.

Nomenclature

aL Gas–liquid interfacial area, cm-1

aS Liquid–solid interfacial area, cm-1

ACC Annualized capital cost, $/yr ACO1 Heat transfer area of CO1, m2

AHE1 Heat transfer area of H.E1, m2 AHE2 Heat transfer area of H.E2, m2 At Total heat transfer area, m2

A0j Pre-exponential factor for reaction j, (mol/cm3)1-n(cm3/g.sec)(mol/cm3)-m

HDAs Hydrpdeasphaltenezation HDM Hydrodemetalization HDN Hydrodenitrogenation HDNi Hydrodenickelation HDS Hydrodesulfurization HDT Hydrotreating HDV Hydrodevanadization CF Furnace cost, $

CH Heating cost, $/yr CHE Heat exchanger cost, $

LiC Concentration of i compound in the liquid phase, mol/cm3

SiC Concentration of i compound in the solid phase, mol/cm3

CP Pump cost, $ cp Specific heat capacity, J/g.K CPU Pumping cost, $/yr CR Capital cost of the reactor, $ Ct Overall annual process cost, $/yr cpw Heat capacity of water, J/g.K cv Specific heat capacity at constant volume, J/g.K DR Reactor diameter, cm

jEA Activation energy for j process, J/mol

ih Henry’s coefficient, MPa.cm3/mol

hp Compressor horsepower

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jK Reaction rate constant for j reaction, (mol/cm3)1-n(cm3/g.sec)(mol/cm3)-m

SHK2

Adsorption equilibrium constant of H2S, cm3/mol

Lik Gas–liquid mass transfer coefficient for i compound, cm/sec

Sik liquid–solid mass transfer coefficient for i compound, cm/sec

LR Length of reactor bed, cm mj Order of reaction of hydrogen in reaction j

mw Mass flow rate of cooling water, g/sec nj Order of reaction of i compound in reaction j M&S Marshall and Swift index for cost escalation N Nitrogen Ni Nickel P1 Pumps1

GiP Partial pressure, MPa

Pin Inlet pressure, Ib/ft3 Pout Outlet pressure, Ib/ft3

q Quench mass flow rate, g/sec Qi Heat duties of H.E1,2, CO1, C1 and F1, J/sec Qin Volumetric flow rate at compressor section, ft3/min Qp Power pump, kW rj Chemical reaction rate of j reaction per unit mass of catalyst, mol/g.sec-1

R Universal gas constant, J/mol.K S Sulfur T Temperature TBRs Trickle Bed Reactors TC0 Inlet temperature of the cold fluid to H.E1, K TC1 Outlet temperature of the cold fluid from H.E1, K TFP Outlet product temperature from CO1, K

TH0 Inlet temperature of the cold fluid to H.E2, K TH1 Outlet temperature of the cold fluid from H.E2, K TP Outlet temperature of the hot products mixture from H.E2, K TP1 Inlet temperature of the hot products mixture to H.E1, K TP2 Outlet temperature of the hot products mixture from H.E1, K TR Inlet reactor temperature, K Tw1 Inlet water temperature to CO1, K Tw2 Outlet water temperature from CO1, K UCO1 Overall heat transfer coefficient for CO1, W/m2 K UHE1 Overall heat transfer coefficient for H.E1, W/m2 K

UHE2 Overall heat transfer coefficient for H.E2, W/m2 K

ug Velocity of the gas, cm/sec uL Velocity of the liquid, cm/sec

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Chemical Product and Process Modeling, Vol. 6 [2011], Iss. 2, Art. 3

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V Vanadium VG Volumetric flow rate of gas, cm3/sec VH2 Volumetric flow rate of hydrogen, cm3/sec

VL Volumetric flow rate of crude oil, cm3/sec z Axial position along the catalyst bed, cm

Greek letters

∆HR Overall heat of reaction, J/mol

B Bulk density of the catalyst particles, g/cm3

L Liquid density at process conditions, g/cm3

G Gas density, g/cm3

j Catalyst effectiveness factor for j reaction

ise Isentropic efficiency

g Gas phase fraction

l Liquid phase fraction

Specific heat ratio

∆Tlm log mean temperature difference

Superscripts

G Gas phase H2 Hydrogen L Liquid phase or gas-liquid interface

S Solid phase or liquid-solid interface

Subscripts

B Bulk F Furnace g Gas

G Gases H2 Hydrogen

H2S Hydrogen sulphide

L Liquid

R Reactor sul Sulfur

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Bouton, G.R. and Luyben, W.L. (2008). Optimum Economic Design and Control of a Gas Permeation Membrane Coupled with the Hydrotreating (HAD) Process. Ind. Eng. Chem. Res., 47, pp 1221.

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Chen, H., Zheng, J., Xu, P., Li, L., Liu, Y. and Bie, H. (2010). Study on real-gas equations of high pressure hydrogen. International Journal of Hydrogen Energy, 35, pp 3100.

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Jarullah, A.T., Mujtaba, I.M., Wood A.S. (2011a). Improvement of the Middle Distillate Yields during Crude Oil Hydrotreatment in a Trickle-Bed Reactor. Energy and Fuels, 25, pp 773.

Jarullah, A.T., Mujtaba, I.M., Wood A.S. (2011b). Kinetic parameter estimation and simulation of trickle-bed reactor for hydrodesulfurization of crude oil. Chemical Engineering Science, 66, pp 859.

Jarullah, A.T., Mujtaba, I.M., Wood A.S. (2011c). Kinetic Model Development and Simulation of Simultaneous Hydrodenitrogentaion and Hyhdrodemetallization of Crude Oil in a Trickle-Bed Reactor. Fuel, 90, pp 2165.

Jarullah, A.T., Mujtaba, I.M., Wood A.S. (2011d). Hydrodeasphaltenization of Crude Oil in a Trickle Bed Reactor: Kinetic Model Development and Process Simulation. Submitted to Chemical Engineering Science.

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Quintero, A.R. and Villamil, V.F.D. (2009). On the Multiplicities of a Catalytic Distillation Column for the Deep Hydrodesulfurization of Light Gas Oil. Ind. Eng. Chem. Res., 48, pp 1259.

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Rodriguez, M.A. and Ancheyta, J. (2004). Modeling of Hydrodesulfurization (HDS), Hydrodenitrogenation (HDN), and the Hydrogenation of Aromatics (HDA) in a Vacuum Gas Oil Hydrotreater. Energy Fuels, 18, pp 789.

Sinnott, R.K. (2005). Chemical Engineering, Volume 6: Chemical Engineering Design. 4th ed. UK, Elsevier Butterworth-Heinemann.

Smith, R. (2005). Chemical Process Design and Integration. UK, John Wiley & Sons.

Tarhan, O.M. (1983). Catalytic Reactor Design. New York, McGraw-Hill. Wauquier, J.P. (1995). Crude Oil: Petroleum Products; Process Flowsheets.

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