FUEL PROCESSING TECHNOLOGY

ELSEVIER Fuel Processing Technology 50 (1997) 275-289

Simulation of partial oxidation of natural gas to synthesis gas using ASPEN PLUS

M. Khoshnoodi **a, Y.S. Lim ’ ’ Department of Chemical Engineering, Unirersi@ of Sistan and Baluchesian. P.O. Box 98135-161. Zahedan.

Iran

h Facult\ of Chemical and Natural Resources Engineering, Univ~witi Teknnologi Malaysia. 54100 Kualrz

Lumpur, Malaysitr

Received 24 July 1996: accepted 24 July 1996

Abstract

Conversion of natural gas to liquid fuels is a challenging issue. In SMDS process natural gas is first partially oxidized with pure oxygen to synthesis gas (a mixture of H, and CO) which is then converted to high quality liquid transportation fuels by utilizing a modernized version of the Fischer-Tropsch reaction. This paper presents a computer simulation of the first stage of the process, i.e. the synthesis gas production from natural gas. ASPEN PLUS equipped with a combustion databank was used for calculations. Concentrations of over 30 combustion species and radicals expected in the synthesis gas have been calculated at equilibrium and several non-equi- librium conditions. Using a sensitivity analysis tool, the relative feed flow rates and reactor parameters have been varied searching to maximize the CO/O, yield as well as to minimize the undesired nitrogen compounds in the product stream. The optimum reactor temperature for maximizing the CO mole fraction in the synthesis gas was also calculated. 0 1997 Elsevier Science B.V.

Kqwnrdrt Natural gas; Synthesis gas; Combustion simulation: ASPEN PLUS

1. Introduction

The first commercial plant for conversion of natural gas to liquid fuels via synthesis gas was started up by Shell in Bintulu, East Malaysia in 1993. The technology is based on the classical Fischer-Tropsch reaction originally explored and developed as far back

_ Corresponding author.

037%3820/97/$17.00 0 1997 Elsevier Science B.V. All rights reserved. PII s037x-3x20(96)01079-x

as the 1920s. It was used in Germany during the second world war to produce motor fuels from coal [I].

Conversion of natural gas to liquid fuels has been a challenging issue. The proposed processes may be classified into two groups; direct and indirect processes [2]. Direct processes tend to be less selective towards the desired products and are limited by the thermodynamic equilibrium:

nCH, * (CHZ),> + nH, (I)

Below 1000 “C the equilibrium is towards the CH 4 side. To overcome this barrier one should either apply a high temperature or create a sink for Hz. This may be achieved by oxidation or by addition of halogens, preferably chlorine. An example of a direct process is the ARC0 process for oxidative coupling reaction in which methane is converted first to ethane and then either thermally or catalytically dehydrogenated to ethylene [3,4]:

2CH, + l/20, + C,H, + H,O - _ (2)

C,H, + C,H, + H, (3)

Liquid fuels are produced by oligomerizaion and aromatization of ethylene over acidic zeolites. This process is best suited for production of gasoline and, because of a conversion yield of less than 25%, cannot compete with indirect processes.

Indirect processes include conversion of natural gas to synthesis gas which is mainly a mixture of H, and CO. Steam reforming of methane (SMR) uses a nickel catalyst and is operated at about 850°C and 30 bar:

CH, + Hz0 + CO + 3H, (4)

It produces a synthesis gas with a HZ/CO ratio of at least 3. As the H,/CO usage ratio in the Fischer-Tropsch reaction is about 2, it is clear that the steam reforming will always result in surplus hydrogen production. A commercial example of such process is the Mobil plant in New Zealand commissioned in 1986 where natural gas is first converted to synthesis gas and then to methanol. Liquid fuel is produced via the Mobil methanol to gasoline (MTG) process over zeolite catalysts. Another example is the Synthol process at SASOL, South Africa. where natural gas conversion to a CO/H, mixture is carried out via steam reforming. Synthesis gas is then converted to high quality gasoline by the Fischer-Tropsch process. It should be mentioned that synthesis gas with different H,/CO ratios is used in the production of many other chemicals [5].

Synthesis gas with a H,/CO ratio of about 2 to 1.7 can be produced by the non-catalytic autothermal partial oxidation process:

CH, + l/20, + CO + 2H, (5)

This reaction is the basis of the Shell Middle Distillate Synthesis (SMDS) process for conversion of natural gas to synthesis gas and then to transportation fuels [6]. The oxidation process is operated at 1300 to 1500°C and pressures up to 70 bar with carbon efficiency in excess of 95% and a methane slip of about I %. The synthesis gas produced consists mainly of Hz and CO (up to 95% vol.), the remainder being H,O, CO?, NT and traces of hydrocarbons and nitrogen compounds as detailed in Section 2. The non-cata-

lytic partial oxidation of natural gas has also been widely used in the production of hydrogen.

The heart of the SMDS process is a modernized version of the classical Fischer- Tropsch reaction in which carbon monoxide and hydrogen are converted to paraffinic hydrocarbons over a cobalt, ruthenium or iron catalyst:

CO+2H.+-CH2-+H,O (61

In this reaction -CH,- represents a segment of a straight chain paraffin in highly linear products. The reaction conditions are chosen so that formation of long chain liquid paraffinic molecules is favoured, whilst that of gaseous components such as butanes and lighter hydrocarbons is minimized. The Fischer-Tropsch reaction is very exothermic and operates in a relatively narrow temperature range of 200-250°C. Taking into account the requirement for massive removal of the reaction heat several types of reactors have been considered: Fluidized bed, Multitubular fixed bed. Three phase fluidized bed. and Bubble column. The basic limitations of these reactor-catalyst systems have been studied in detail [7,8]. The Fischer-Tropsch process may lead to a raw product of a rather waxy nature which is unsuitable for transportation fuels. Incorporation of a hydroisomerization and mild hydrocracking stage reduces heavy paraffins to middle distillates which are then fractionated to kerosene. gasoil and some naphtha. In the Shell MDS Bintulu plant. 100 million scfd natural gas is partially combusted with pure oxygen. produced in a 2500 ton dd ’ air separation unit, resulting in 1200 bbl d-’ liquid fuels. These fuels are completely free from aromatics and sulphur compounds and upon combustion in vehicles result in much less particulate and sulphur emissions when compared with middle distillates originating from crude oil [9].

This project is a modeling investigation of the first stage of the SMDS process; that is synthesis gas production which involves more than 50% of the total capital cost.

2. Combustion simulation by ASPEN PLUS

ASPEN PLUS (Advanced System for Process Engineering) Release 9.1-3 (1994) was employed to simulate the partial oxidation reactions of natural gas for the production of synthesis gas. This process simulator has been developed at MIT, USA, and is equipped with up to date databanks for thermochemical properties based on the American Institute of Chemical Engineers DIPPR data compilation project as well as a combustion databank based on the JANAF Tables including 59 combustion species and radicals at temperatures up to 6000 K [IO].

ASPEN PLUS incorporates most unit operations and several types of reactors, including the Gibbs reactor [I I]. For a multireaction system such as the partial combustion of natural gas which involves numerous dissociation, recombination and elementary reactions, the Gibbs reactor was preferred because it is based on the minimization of the total Gibbs free energy of the product mixture. In this method. for the calculation of the equilibrium composition in a set of chemical reactions. the reactants with their initial amounts and conditions are specified. It is also required to specify what species are expected in the product stream. After defining reactor tempera-

ture and pressure, the composition of all species present in the product mixture is calculated by solving a set of simultaneous equations [ 12.131. In this simulation work the product synthesis gas was considered to contain over 30 species: Hz. CO. H,O. CO,. N?, CH,, H,N. H. CHN. CH,O. CHO, HO, NO. CHz, CN, 0, HN. C,H,. HNO, 02, N. N,O, H,O. CH, n-C,H,,,. C. I’-C,H,,,. NO,. I’-C,H ,?. n-C5H,, and N,O, respec- tively.

3. Simulation procedure and results

The flow diagram for the simulation of the partial oxidation of natural gas to produce synthesis gas is shown in Fig. 1. Separate streams of natural gas (14946 kg hh’), pure oxygen (17783 kg h- ’ ) and some hydrogen (57 kg hh ’ ), for desulphurization, were directed into a virtual Gibbs reactor and converted to product synthesis gas (32787 kg h- ’ ). Stream flowrates, temperatures, pressures, and other process parameters were obtained from the operational data provided by the manufacturing offices of the Shell MDS Malaysia Sdn. Bhd. in Sarawak. The composition of natural gas from the central Luconia field in South China sea is shown in Table 1. The main objective of this modeling project was to vary the relative feed flowrates and reactor parameters to maximize the CO/O, yield as well as to minimize the undesired trace nitrogen

( Temperature (“C)

a Pressure (BAR )

> Flow Rate ( KGIHR)

Fig. 1. Flow diagram for the simulation of the partial oxidation of natural gas.

M. Khoshnoodi, Y.S. Lirn / Fuel Processing Technolog! 50 C 1997) 275-280 ?I9

Table I

Composition of the natural gas from South China sea, central Luconia field

No. Component Mole %

I CH, 87.975 2 C,H, 3.543 3 co, 3.543 4 C,H, 2.194 5 N2 0.97 I 6 i-C,H,,, 0.574 7 n-C,H,,, 0.536 8 i-C,H,? 0.207 9 n-C,H,, 0.126 10 &HI, + 0.331

TOTAL 100.00

compounds found in the synthesis gas stream [14]. The different cases studied and the corresponding results are as follows:

3. I. General simulation

A general simulation of the partial oxidation of natural gas for the production of synthesis gas was carried out. Perfect mixing of the reactants and ideal gas behavior of the hot gases were the main assumptions. Complete equilibrium was assumed when the reactor temperature was 1390°C and pressure 46 bar. The mole fractions of the outlet synthesis gas containing over 30 species were calculated; H, = 0.5502, CO = 0.3326, H,O = 0.0954, CO, = O-0177,... to as low as C,H ,? = 1.8E - 22. A complete report for the streams specifications was generated, as shown in Table 2.

3.2. DMDS injection

The general simulation was repeated with a small amount of Dimethyl-Disulfide (DMDS, 0.64 kg h-r ) injected into the reactor. Although it was advised by the Shell MDS that the additive may improve the CO concentration in the synthesis gas, the simulation results did not show any significant changes.

3.3. Non-equilibrium simulation

Experimental data provided by the Shell MDS indicated a slight methane slip through the reactor. The synthesis gas contained 0.197 mole % of unreacted CH, which corresponded to a 0.39% departure from the equilibrium conditions of the reacted mixture. Thus, non-equilibrium simulation was carried out and departures from the equilibrium were set at 0.29%, 0.39% and 0.49%. Streams specifications reports were generated for each case. Table 3 shows the results for the 0.39% case. There is 4.944 Kmol hh ’ unreacted methane in 25 11.967 kmol hh ’ synthesis gas, corresponding to 0.1968 mole %. This is very close to the above value, read from the plant operation data.

Pamal oxidation of tialurill gas

Stream ID From

Tll Phaw

Subatream: MIXED

Mole Flow

CO CO2

H20

N2

H2

02

HO

NO

H

0 HO2

NO2 N30

CN

HNO

N

CHO

CH

CHN

HN

Cl

C2

C3

IC1

NC4

ICS

NC5

Ch

cos

H2S

H3N N203

CH30

CH2

C

DMDS

cs2 Total Flow Total Flow

Total Flow Temperature Pressure

Vapor Frx.

hmol h

kmol II

kgh-’ cutnh-’ “C bar

<;a\

Kcactol Vapol

0.0 27.65403

0.0

7.578905

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

6X6.6675

27.65303

17.12373

4.4x02 IX

4.1X361X

I.615688

.Y83462S

7.583530

0.0

0.0

0.0 0.0

0.0 0.0

0.0

0.0

0.0 780.5258

13945.83 879. I222

350.0000

46.00000 I .oOOoOO

ouygen

Vapol-

0.0 X37.272 I

0.0 14SYYY6

0.0 3-w.3262

0.0 7.S187.57

0.0 I3XS.203

5ss.7500 3.OXl5E- IO 0.0 2.41 125E-3

0.0 3. I Y690E - 7

0.0 .028X121

0.0 3.7462E-Y

0.0 5.1284E- I2

0.0 5.3x!- IS 0.0 6.Y848E - I2 0.0 1.6093SE - Y

0.0 6.OXX2E - IO 0.0 5.1259E ~ I I 0.0 3.779054

0.0 2.3llXE- I7 0.0 .0159220 0.0 1.7417x-9 0.0 2.25x51.1

0.0 2.38524E - 5

0.0 6.6XlXE- IO 0.0 6.326YE ~ IS 0.0 I .89XOE - I1 0.0 0.0

0.0 0.0 0.0 0.0

0.0 0.0

0.0 0.0 0.0 IO43728

0.0 0.0

0.0 Y23636E - 3

0.0 2.98888E - X

0.0 I.UIOE- I-1

0.0 0.0

0.0 0.0 555.7500 25 17.339

177x3.33 327X6.50

525.5029 7576.326

250.0000 13 I9.000

46.00000 46.00000 I .ooOoOO I .OOOOOO

Partial oxidation of natural gas

Liquid Frac.

Solid Frdc.

Enthalpy kcal tnol~ ’ Enthalpy kcal kg -I

Enthalpy mmkcal h- ’ Ent1-q’) CAL/MOL-K

Entropy cal/gm-k

Density kmol/cum

Density kg/cum

A\eragc MW

0.0 0.0

0.0 0.0 - I7.08035 2.267678

- XY 1.9982 I 114.007

~ 13.33185 .0634956

--- I Y.72065 - 7.338130

- I .02988 I - I .709606

.8X78567 .8878467

17.00086 I .7x9791

IY.11832 2.01588U

0.0

0.0 I .624103

50.76449

.YO27139

-3.530153

-.I lO3?1-!

I .057558

33.81060

3 I .99X80

0.0 0.0

- 5.33248 I - 409.1261

- 13.12385

I3.55003 I .04036X

.337658Y

4.332640

13.02-1’7

3.4. Muximixkm of CO / 0, yield _

A sensitivity analysis tool is one of the ASPEN PLUS features. Keeping the reactor conditions and inlet parameters constant, as in the general simulation case, the oxygen tlowrate was increased gradually from 12500 to 21000 kg h- ’ with a step change of 200 kg h- ’ This was done to monitor the variation of CO concentration in the product stream. Graphical results showed that CO mole fraction reached a maximum of 0.3454

LIi /, I I1 1 i j / a, I1 / 1. : 3 / ,

13100 13250 13400 13550 13700 13850 14000 14150 14300 14450 14600 14750 02 MASS FLOW RATE (KGIHR)

Fig. 2. Maximization of CO concentration in the synthesis gas as d function of the oxygen mass flowrate. at

1390°C and 46 bar.

Partial oxidation of natural gas

Stream ID From To Phase Substream: MIXED

Gas

Reactor Vapor

Mole Flow co CO2 H20 N2 H2 02 HO NO H 0 HO2 NO2 N20 CN HNO N CHO CH CHN HN Cl C2 C3 lC4 NC4 lC5 NC5 C6 cos H2S H3N N203 CH20 CH2 C Total Flow Total Flow Total Flow Temperature Pressure Vapor Frac. Liquid Frac. Solid Frac.

kmolhK

kmolh-’ kgh-’ cumh-’ “C bar

0.0 0.0 21.65403 0.0

0.0 0.0 7.578905 0.0 0.0 28.439 I9 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0

686.6675 0.0 27.65403 0.0 17.12413 0.0 4.4802 I8 0.0 4.183618 0.0 I .6 I5688 0.0 .9834625 0.0

2.583540 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0

780.5258 28.43919 14945.83 57.33000

879.1222 32.03 165 350.0000 350.0000

46.00000 46.OOOOO I .ooGooo 1.OOOOOO 0.0 0.0 0.0 0.0

Hydrogen Oxygen

Reactor Reactor Vapor Vapor

0.0 834.1025 0.0 45.08443 0.0 242.5269 0.0 7.5 19222 0.0 1377.632

555.7500 3.166OE- IO 0.0 2.437418-4 0.0 3.24055E - 7 0.0 .0287026 0.0 3.7932lE-9 0.0 5.2602E - 12 0.0 5.4772E - 15 0.0 7.0879E - I2 0.0 4.53034E - 9 0.0 6.16lOE- 10 0.0 5.4203E - I I 0.0 3.75846E-4 0.0 2.2659E- I2 0.0 .0156229 0.0 1.73704E-9 0.0 4.944006 0.0 3.28102E-5 0.0 6.2659E - IO 0.0 5.8179E- 15 0.0 I .7429E - 14 0.0 4.1771E- I9 0.0 4.1415E- I9 0.0 8.27218-24 0.0 0.0 0.0 0.0 0.0 1037427 0.0 3.75008 - 28 0.0 9.16073E-3 0.0 2.92462E - 8 0.0 l.4148E- 14

555.7500 25 I I.967 17783.33 32786.50

525.5029 7551.178 25O.ooOO 1390.000

46.ooooO 46.00000 1.OOOOOO I .oOOOOO 0.0 0.0 0.0 0.0

Product Reactor

Vapor

M. Khoshnoodi. Y.S. Lim / Fuel Processing Technology 50 (19971275-289 283

Table 3 (continued)

Partial oxidation of natural gas

Enthalpy kcal mol _ ’ Enthalpy kcal kg- ’ Enthalpy mmkcal h- ’ Entropy cal/mol-k Entropy cal/gm-k Density kmol/cum Density kg/cum Average MW

- 17.08035 2.267678 I .624403 -5.402717 - 891.9982 I 124.907 50.76449 - 4 13.9340

- 13.33185 .06449 IX .9027749 - 13.57164 - 19.72065 - 2.438420 -3.530152 13.54269

- I .029884 - 1.209606 -.I103214 I .037585 .8878467 .8878467 1.057558 .3326589

17.00086 I .789792 33.84060 4.341905 19.14842 2.015880 3 I .99880 13.05212

when 0, flowrate was 13650 kg hP ’ , resulting in a maximum CO/O, yield of 1.758 by mass. Fig. 2 shows this maximization when the oxygen flowrate range was narrowed down to 13000- 15000 kg h- ’ with a step change of 50 kg h- ’ . Streams specifications at this condition of maximum CO production are shown in Table 4. CO concentration goes through a maximum because the oxidation process goes through CO production. An increased supply of oxygen results in further oxidation of CO to CO,.

3.5. Minimization qf trace nitrogen compounds

A sensitivity analysis was carried out to investigate variation in concentration of other species as a function of the oxygen flowrate, especially the undesired nitrogen com-

fi[ ‘%

ot d ,,1,,,,,,,,,,1,,,,,,,,,,,,,,,,,,,,,,, ,,,,,,,,,,,,,,,, LIII,

16650 17200 17550 17900 18250 16600 18950 19300

=~+&

19650 20000 20350 20700 02 MASS FLOW RATE (KGIHR)

Fig. 3. Variation of H,N concentration in the synthesis gas as a function of the oxygen mass flowrate

Stream ID From

To

Phase

S&stream: MIXED

Mole Flow

CO

CO2

H20

N2

H?

02

HO

NO

H

0

HO2

NO2

N20

CN

HNO

N

CHO

CH

CHN

HN

Cl (‘2

C3

1C4

NC4 ICS

NC5 C6

COS

H2S H3N

N203

CH20

CH2 C

DMDS

CS2

Total Flow

Total Flow Total Flow Temperature

Pressure

Vapor Frac.

Gl\

Reactm

VLlp0!-

kmol h

0.0

27.65403

0.0

7.578905

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

6X6.6675

27.65303

17.12473

1.4802 IX

-1. I X36 IX

I .61568X

.9X34625

2.583540

0.0

0.0 0.0

0.0

0.0

0.0 0.0

0.0

0.0 kmolh ’ 780.5258

kph-’ 14945.X3 cumh-’ x79.1222 “C 350.0000 bar 46.00000

I .oooooo

Hydrogen oxygen

0.0

0.0

0.0

0.0

2X.34085

0.0

0.0

0.0

0.0

0.0

0.0 0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0

0.0 28.44085

57.33333

32.0335 I ~50.0000 46.00000

I .oooooo

0.0 XS6.7 I 83 0.0 6.555397

0.0 3X.62486

0.0 7.355856

0.0 I54Y.839

426.5785 h2646E - I? 0.0 3.63662E - S

0.0 -1.S3914E-X

0.0 .0302SO8

0.0 S.03020E - I3

0.0 1.1 I IOE- 13

0.0 I.O86lE- 16

0.0 Y.Y494E ~ I3 0.0 3.79396E ~ X

0.0 9.2117E- I I 0.0 S.3632E - I I 0.0 1. I20hSE - 3 0.0 I .7549E - I I 0.0 .I112573 0.0 I .X3463E - Y

0.0 20.73656

0.0 I .7YX9OE - 3

0.0 4.13734E - 7

0.0 3.2159& 1 I 0.0 YWYOE-II

0.0 I .Y33SE - I4

0.0 I.Y17OE- I3

0.0 3.20hYE -- IX

0.0 0.0

0.0 0 0

0.0 ,I218352

0.0 1 .OS49E ~ 2X

0.0 .o IO7207

0.0 2.4 I771E ~ 7

0.0 I .026SE ~ 13

0.0 0.0

0.0 0.0 426.S785 74x0.129

I3650.00 28653. I I

403.3616 7455.77 I

250.0000 131’).000

46.00000 46.00000 I .nnoooo I .nnoooo

Vapor

Table 4 (continued)

Partial oxidation of natural gas Liquid Frac. Solid Frac. Enthalpy kcal mol- ’ Enthalpy kcalkg-’ Enthalpy mmkcal h- ’ Entropy cal/mol-k Entropy cal/_gm-k Density kmol/cum Density kg/cum Average MW

(0.0 0.0 0.0 0.0

- 17.08035 2.267b78 - X9 L .9982 1124.907 - 13.33185 .0644956 - 19.72065 - 2.438420

- I xl2988 1 - 1 .I!09606 .X878467 .8878467

17.00086 I .x+9792 19.14842 2.015880

0.0 0.0 0.0 0.0 i .624403 - .0491280

50.76449 - 4.252540 .6929539 - .1218505

- 3.530152 13.78313 -.1103214 1.193073

I It57558 .3326589 33.84060 3.843086 31.99880 1 1.55263

pounds, detected in the synthesis gas. Keeping all reactor parameters constant as before, the oxygen flowrate was increased from 16000 to 21000 with a step change of 200 kg h-r. The composition of the product stream was calculated for each step, streams specifications tables were generated and species concentration variation curves were plotted. Figs. 3-7 show that by increasing the OJCH, ratio in the reactor feed, the trace concentrations of H,N, CHN, CN and HN in the synthesis gas decreased proportionally. However, the mole fractions of NO, HNO, N,O and N,O, increased due to increased oxidation. Sensitivity analysis results also indicate that further reduction in

16650 17200 17550 17900 18250 18600 18950 19300 19650 20000 20350 ‘20700 02 MASS FLOW RATE (KGfHR)

Fig. 4. Variations of CHN, H and CHZO concentrations in the synthesis gas as functions of the oxy@~ mas: flowrate.

,, ,,? :’

,ir ,A7 ’

02 MASS FLOW RATE (KG/HIV

Variations of CN, 0 and HN concentrations in the synthesis gas as functions of the oxygen mass flowrate

16660 17200 17550 17900 16250 16600 16950 19900 19650 20000 20350 20700 02 MASS FLOW RATE (KG/HR)

Fig. 6. Variations of C,H,, O2 and HNO concentrations in the synthesis gas as functions of the oxygen ~I~dss flowrate.

M. Khoshnoodi, Y.S. Lim / Fuel Processing Technology 50 (1997) 2 75-289 287

02 MASS FLOW RATE (KG/HR)

Fig. 7. Variations of NO and CH, concentrations in the synthesis gas as functions of the oxygen mass flOWdtf2.

R- /- 2; C-G a 1 6:

,I %I

z dC 9 2% 1

E 0 1 22. /’ 0 “: 20. !’

0 1 j 0% “! 8

0 : gi: / m- d

i ! i

gt / “1 ’ ??i

ad “1

I,, , , , 1 I I I Ia 10 14 1 ” e ’ 1000 1100 1200 1300 1400 1500 1600 1700 1600 1900 2000 REACTOR TEMPERATURE (C)

Fig. 8. Variation of CO concentration in the synthesis gas as a function of the reactor temperature at 46 bar.

oxygen flowrate, would not change the trend of the concentration curves and they do not go through a maximum or a minimum. The range of the oxygen flowrate was chosen so that most nitrogen compounds had comparable concentrations. However, these concen- trations at the condition of maximum CO production are calculated and shown in Table 4.

3.6. Optimization of reactor temperature

A sensitivity analysis was carried out to find the reactor temperature which results in maximum CO concentration in the synthesis gas stream. Keeping all parameters constant as in the general simulation case (46 bar), the reactor temperature was increased gradually from 1000 to 2000°C with a step change of 10 “C. Fig. 8 shows that the CO mole fraction rose steeply to 0.3328 at 14OO”C, and after that it remained almost constant. The experimental value for the synthesis gas reactor temperature was 1390 “C resulting in a CO mole % of 0.3326.

4. Conclusion

The ASPEN PLUS process simulator was found to be a powerful package to simulate the process of synthesis gas production from partial oxidation of natural gas. The combustion databank which contains the thermochemical properties of radicals allowed modeling of complex steady state combustion calculations. Considering the variety of applications for synthesis gas in the production of many chemicals especially liquid fuels, the process simulator proved helpful for operational modifications as well as design considerations. General simulation results were comparable with previous work and other simulation software [ 15,161. Sensitivity analysis modeling allowed the predic- tion of the composition of the synthesis gas when the relative feed flowrates or reactor parameters were varied over a wide span, without real experimentation on the plant which could disturb production operation. Most experimental data from the Shell MDS, Bintulu plant could be repeated successfully in virtual operations and some production guidelines were suggested by utilizing sensitivity analysis and design specification tools incorporated in the simulator 1141.

Acknowledgements

The authors would like to express their gratitude to the Shell MDS Malaysia Sdn. Bhd. for the technical and financial support of this project. Assistance from the staff of the computer center at the Faculty of Chemical and Natural Resources Engineering, Universiti Teknologi Malaysia and in the University of Sistan and Baluchestan is also appreciated.

M. Khoshnoodi, Y.S. Linl/ Fuel Processiq Technology 50 (1997) 275-289 289

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