1-s2.0-s0255270104001461-main

Chemical Engineering and Processing 44 (2005) 421428

Energetic and economic evaluationne oM. Billstatt

ptembert 2004

Abstract

Acetic ac feed cparameter in compofeed was var ect oxiby methanol ed. 2004 Else

Keywords: A

1. Introduction

The oxidative conversion of light hydrocarbons to morevaluable products is of high industrial interest. Acetic acidis an impotion over eto acetic action procesthe cheap fprocess stucess econom

carbonylatiThe me

1913 and rBASF in 19and high-teand in 197in Texas Ccatalyst). Tmuch mildare two lead

CorrespoE-mail a

BP. The BP process uses novel irridium or rhodium catalyst[2,3]. The Celanese process is based on the rhodium cata-lyst [4]. The world leading acetic acid producers are listed inTable 1. In present study, the set-up of BP was simulated and

0255-2701/$doi:10.1016/jrtant industrial product with world-wide produc-ight million tons. The direct oxidation of ethaneid can be an alternative to methanol carbonyla-s because of its high selectivity and because ofeedstock. The main target of the novel acetic aciddies is to find process conditions, where the pro-

y can compete with the state-of-the-art methanolon process at the same acetic acid quality.thanol carbonylation process was developed inealised (with pioneer cobalt iodide catalyst) by60 in Ludwigshafen, Germany. The high-pressuremperature process was improved by Monsanto

0 commercialised for a production of 135 kt/yearity, Texas (with a new iodide-promoted rhodiumhe operating conditions (30 bar and 453 K) wereer then in the BASF set-up [1]. Nowadays, thereers in acetic acid production, that is Celanese and

nding author. Tel.: +49 30 6392 4456; fax: +49 30 6392 4004.ddress: [email protected] (Q. Smejkal).

compared to the ethane oxidation process due to the moredetailed references and literature published. The BP processcould be taken as a successor of the original Monsanto pro-cess. The ethane direct oxidation (EDO) process is not yetcommercialised, although SABIC announced in July 2002 a30 kt/year acetic acid plant based on the EDO process, whichhas been developed by their Research and Development Cen-ter in Riyadh, Saudi Arabia [5,6].

In the present study, acetic acid production by ethane di-rect oxidation with respect to feed composition and reactorperformance was studied. The evaluation of the process flow-sheet was based on a simulation by ASPEN PLUSTM. Thedata from previous kinetic modelling [7] were used as a basisfor the simulation of a non-isothermal fixed-bed reactor forthe catalytic ethane conversion. The prediction of reactor per-formance as well as cooling down of the product stream andits separation in a series of heat exchangers and rectificationcolumns were implemented into the process simulations.

The aim of this paper is to (i) analyse the influence of feedcomposition, dilution mode and reactor performance and the

see front matter 2004 Elsevier B.V. All rights reserved..cep.2004.06.004of acetic acid via ethaQ. Smejkal, D. Linke,

Institut fur Angewandte Chemie Berlin-Adlershof e.V., Richard-WReceived 10 April 2003; received in revised form 15 Se

Available online 19 Augus

id production via the selective oxidation of ethane was studied. Thereactor and process simulation. The concentration of water (as a

ied. Heat and mass balances were predicted. Finally, the ethane dircarbonylation and the investment and production costs are discuss

vier B.V. All rights reserved.

cetic acid; Ethane oxidation; Reactor simulation; Process flowsheetof the productionxidationaernser-Str. 12, D-12489 Berlin, Germany

2003; accepted 4 June 2004

omposition and mode of dilution was taken as a majornent improving acetic acid selectivity) in the reactiondation process was compared to acetic acid production

422 Q. Smejkal et al. / Chemical Engineering and Processing 44 (2005) 421428

Table 1The world leading acetic acid producers (in thousands of m.t./year, [6])Company Global capacity

Celanese 2065BP Chemicals 1175Millennium Chemicals 450Acetex 400

separation effort based on EDO; (ii) demonstrate the featuresand benefits of direct ethane oxidation to acetic acid com-pared to the industrial methanol carbonylation process; (iii)predict the investment and production costs.

2. Process description

2.1. Ethane oxidation

2.1.1. Reactor designA multi-channel fixed-bed reactor of 30000 tubes (length

3 m, i.d. 25 mm) was described by a pseudo-homogeneousone-dimensional model for acetic acid production of50 kt/year. The mass and energy balances for the steady-stateare given elsewhere [7,8]. The heat transfer coefficient be-tween the catalyst bed and the wall was derived by the correla-tion of Hennecke and Schlunder [9] and Tsotas and Schlunder[10]. Depending on the reaction conditions (gas velocity,composition of gas feed, pellet size), the heat conductivity ofthe catalysThe respec

2

S-

Fig. 1. The scooler, 5: flassteps, p = 16 bflash.

tween 27 and 48 W m2 K1 for a stainless steel tube of i.d.= 25 mm. The differential equations of mass and energy bal-ances were solved applying a numerical integration by theGEAR-algorithm implemented in FORTRAN [11].

2.1.2. Process designThe scheme of the reaction and separation set-up is illus-

trated in Fig. 1 and was based on known technologies forthe production of acetic acid [12,13]. It includes cooling ofthe reaction mixture, flash of non-condensed components andrectification of wateracetic acid mixture with a limitation ofwater content in acetic acid

Q. Smejkal et al. / Chemical Engineering and Processing 44 (2005) 421428 423

Table 2Constants of NRTL equation of state used in simulation of acetic acidwaterseparation

Parameter Value

ai, j 1.9763aj , i 3.3293bi, j 609.8886bj , i 723.8881ci, j 0.3

The phase equilibrium acetic acidwater was describedusing NRTL equation of state. The parameters of the NRTLequation are listed in Table 2 and were taken from ASPENPlus databank (databank VLE-HOC).

2.1.3. ModThe calc

productionther below.selectivityof a polytroPEN PLUSture rangewas derivedof the coolitemperaturrecycled etcompositiotion to the rblack-boxoxidation ato a water-tration waswith coxygeoxygen conThe compaof water onon the sepato the feedfor the ethaence of airthe feed iscomparison

Table 3Description oTfeed = 515 KVariant

ABCDEaF

a When us

CO2). In case of variant F, only a portion of produced CO2is separated from the recycle stream.

2.2. Metha

For mehomogeneo(Rh(CO)x)CO) and pr30 bar, twotank reactoin a flash (5lyst is sepafeed. In a sseparated a

tilledeticalafter

led bae pr

anol ainto t

ompote, 5.6c acid,550 pin the

catalility suthe sito there (8sed in

or, coolloy a

esults

Proce

e resuvarianmma

ced ir the hperatid by the of process simulation for ethane oxidationulations were performed for a constant plant scale50 kt/year for each process variant discussed fur-The feed composition, conversion of ethane and

of acetic acid were obtained from the simulationpic fixed-bed reactor and were used for the AS-TM simulation. Therefore, the optimal tempera-

with respect to thermally stable reactor operation, that is feed temperature 515 K and temperature

ng medium 515 K. The maximal tolerable reactore in the hot spot was set to 550 K. The amount ofhane was obtained by fitting the predicted productn obtained by polytropic fixed-bed reactor simula-esults of process simulation for which a so-calledreactor unit was used. The variants of the ethanere summarised in Table 3. Variants A and B referfree feed. Here, the influence of oxygen concen-studied. Since all fixed-bed reactor simulations

n > 14 mole % led to runaway of the reactor thecentration was limited to 13 mole % in the feed.rison of variants A, C and D shows the influencethe energy balance of the process and, especially,ration effort. A positive effect of water additionon the selectivity to acetic acid is well knownne oxidation reaction [7]. In variant E, the influ-as an oxidising agent in the presence of water inconsidered. Variants A, B, E and F deal with theof different dilution agents (ethane, nitrogen and

f the EDO to acetic acid process variants, GHSV = 1200 h1,, p = 16 bar, calculated for 50 kt/year production of acetic acidMolar fraction

is distheorCO2;recyc

Thmethouslyage cacetaacetisalt),didelationvolat

Inbackmixtuture ureactHaste

3. R

3.1.

Thcess

are su

produhigheare o

duceEthane Oxygen Water Nitrogena CO20.87 0.13 0 0 00.95 0.05 0 0 00.8 0.08 0.12 0 00.85 0.1 0.05 0 00.85 0.02 0.05 0.08 00.73 0.12 0 0 0.15ing nitrogen as dilution agent.

can be reusof the feed

3.1.1. Inubalance an

For analergy balancditions seenol carbonylation

thanol carbonylation, the liquid methanol andus catalyst in form of methyl iodide complexis mixed (Fig. 2) with a gaseous feed (compressede-heated to the reaction conditions [3,14] 462 K,

phase system. The carbonylation in the stirredr is followed by separation of gas and liquid phases). From the liquid phase, the homogeneous cata-

rated by rectification (7) and recycled back to themall rectification column (8), a portion of CO isnd consequently the water and acetic acid mixturein the rectification column (6) with efficiency 47stages. Gas phase from flash consists of CO andpurification (9: CO2 selective absorption) CO isck to the feedgas.ocess specification is as follows: 33.4 wt.%nd 31 wt.% carbon monoxide were fed continu-he reactor held at temperature of 462 K. The aver-sition of the reactor contents was: 2.6 wt.% methylwt.% water, 14.0 wt.% methyl iodide, 61.9 wt.%0.55 wt.% lithium (present at least in part as iodidepm rhodium and 11.6 wt.% iodide. The lithium io-reactor composition functioned both as a carbony-yst stabiliser in the reactor and a water relativeppressant in the flash tank.

mulation, the homogeneous catalyst was recycledreaction feed after separation from the reaction

) [14]. Due to the highly corrosive catalyst mix-the process parts of the unit (at least pre-heater,

ler and catalyst separator) have to be made fromlloy [15,16].

and discussion

ss simulation results

lts of the process simulation for the different pro-ts (see Table 2) for the ethane oxidation processrised in Tables 3 and 4. The energy consumed andn the process is shown in Table 2. Generally, theeat consumption of the process variant, the higher

ng and utility costs of the process. The heat pro-e reaction or the HD steam formed in cooler H-4ed in other process operations, i.e. for pre-heatingor for heating the rectification column.

ence of the mode of dilution on the energyd ethane consumption per mole acetic acidysis of the influence of dilution mode on the en-e, variants A, B, E and F were considered (con-Table 3). The advantage of the process variant A


H-2

CST-3 S-5

P-1

S-9

C-6

C-7

Acetic acid

Water

CO2

Catalyst recycle

Feed

Recycle

H-4

Fig. 2. The sc sor, 2:wateracetic pre-sepreaction temp etters ucompressor, C

and B (whevariant E wfrom Tableto the extenreaction mipression anvariant A want B correoxygen incomparedof energy c

The varreaction mvariant amo

Table 4Summary resmethanol carb

Variant

ABCDEF

BP

ion anis lowen inheme of acetic acid processes: methanol carbonylation by BP; 1: compresacid (purification of acetic acid), 7: separation of homogeneous catalyst, 8:erature to 303 K in one step, p = 30 bar, flash: 303 K, p = 30 bar. Code lST: reactor, S: absorber, separator, flash.

re only ethane and oxygen are fed) compared toith air used as an oxidising agent can be deduced3. Air as an oxidising agent is not favourable due

pressCO2oxygded effort for compression and pre-heating of thexture (variant E). For variant A and B, lower com-d separation energy consumption is predicted forith higher oxygen concentration in the feed. Vari-sponds to the variant with 95% ethane and 5%the feed; i.e. the oxygen content was decreasedto variant A. Obviously, this leads to an increaseonsumption.iant F of EDO process, where the dilution of theixture is realised by CO2 is the most promisingng all process variants discussed above. The com-

ults of the computer simulation of variants AF by EDO process, energy balanceonylation (see Chapter 3.1.3), 50 kt/year production of acetic acid (HD: high dens

Heat consumption (kW)Compressor Heater Cooler Rector Tota

6120 1400 7340 7400 2214000 3630 14050 7200 3810000 7800 16600 21000 55

8500 2300 12100 11250 3442000 16000 45000 27000 130

5100 600 6600 7400 19

670 3600 2900 13000 20

dioxide. Mgen in thethan ethane

3.1.2. Inuthe energyacetic acid

Variantin the feedon the acecomparisonC-8

heating, 3: CST reactor, 4: cooler, 5: flash, 6: rectification ofaration of reaction mixture, 9: CO2 separation, cooling: fromsed in Fig. 1: C: rectification column, H: heat exchanger, P:

d heating/cooling effort for EDO variant F wither compared to variant A with only ethane and

the feed due to the lower heat capacity of carbonof ethane oxidation to acetic acid and BP variant based onity steam)

Heat production (kW)l Reactor HD steam Total

300 29000 2300 31300900 25200 5600 30800400 26500 4100 30600150 29300 3500 32800000 22400 17000 39400700 26600 1800 28400

500 7600 0 7600

oreover, for nearly the same concentration of oxy-feed, the CO2 seems to be better dilution agent(compare variants A and F).

ence of the water concentration in the feed onbalance and ethane consumption per mole

C represents the highest concentration of water(13 mole %). The positive influence of water

tic acid selectivity is obvious from Table 5. Theof variant C with variant A, D and F shows, that


Table 5Selectivity and ethane consumption of variants AF, EDO process, TFeed =515 K, p = 16 bar, 50 kt/year production of acetic acid, 100% conversion ofoxygen

Variant

ABCDEF

the increasconsumptioTable 5). Owith increaantagonistithe selectivconcentratihand variCO2) seem

3.1.3. ComThe com

ethane withcesses is doplant of 50the set-up wby variant12 mole %ture and feefor compreof methanopre-heatingin case of Ereaction iscarbonylatiduced only

4. Cost an

The objeconomicsrelevant mtotal investof:

(1) fixed (buildin

(2) producmateriaand COcatalys

The basgiven in Ta

Table 6Design assumptions for acetic acid production calculated for year 2002

Parameter Value/type

ityure

s locatis contrf proce

onditionre vess

procesng toe pronylatiesultsO, wthe fl

predic

Invest

e bloanol cn in Ftal equion coratuseumma

e pricrial cosive css, themadee specatalyste equge ofp by Breduc

on anthe respecinylati

for catSelectivity ofacetic acid (%)

Consumption of ethane(mol)/mol acetic acid

76.3 1.3177.7 1.2981.0 1.2379.6 1.2580.7 1.2476.1 1.31

e of the selectivity to acetic acid results to lowern of ethane per formed unit of ethylene (seen the contrary, the separation effort is increasingsing amount of water. Taking into account thec requirements high water content to improveity to acetic acid on one hand, and low wateron to minimise the separation effort on the otherant F (73 mole % ethane, 12 mole % oxygen, 15%s to be a best variant of EDO (see Table 4).

parison of the EDO and BP processparison of variant F of selective oxidation ofthe commercial BP methanol carbonylation pro-

cumented in Table 4 for an acetic acid productionkt/year. Among the ethane oxidation processes,ith the highest economic potential is representedF with a feed composition 73 mole % ethane,oxygen and 15 mole % CO2 a cooling tempera-d temperature of 515 K. The energy consumptionssion and product cooling is the lowest in the casel carbonylation. On the contrary, the effort for theand product purification (rectification) is lowerDO variant F; the heat produced by the chemical

higher in the EDO process compared to methanolon. Moreover, the high-density steam can be pro-in case of the EDO process (see Table 4).

alysis

ective of this part of the study is to evaluate theof ethane direct oxidation compared to industry

ethanol carbonylation process by BP [3,14]. Thement and production costs of acetic acid consist

CapacProced

ProcesProcesType oSoil cPressu

acidcordi

Thcarbounit rof EDfromcost

4.1.

Thmethshowon totimatappaare s

Thmatecorro

proceto beThesthe c

Thvantaset-umainficatiAlsoto thecarboumninvestment) costs, including apparatuses, project,gs, labor, analysis and control [17,18]tion costs, i.e. variable costs, including the rawl ethane and oxygen (EDO process) and methanol(BP process) price and the utilities steam, gases,

t, cooling water, fuel, energy, etc.

ic assumptions used in process cost estimation areble 6. A low to high-scale 10200 kt/year acetic

all other coused non-liof compresproductionmajor contcess (almos

The inveent methodpend on a nprocess and10200 kt/year acetic acid(a) EDO(b) Methanol carbonylation

on Europeol Digitalss Grass roots/clear field

Soft clayel design code DIN

s was used located in Europe and designed ac-DIN.cess flowsheet for ethane oxidation and methanolon was adopted for the cost analysis. The plantdiscussed above show the advantage of variant Fhere CO2 is used as a dilution agent. The resultsowsheet simulation were used as an input file fortion routine ICARUS by ASPEN PLUSTM.

ment cost

ck flowseet for ethane direct oxidation andarbonylation is identical with process schemesigs. 1 and 2. The total investment costs are basedipment costs. Therefore, the first step of cost es-mprises the calculation of the purchase costs of

s included in the process simulation. The resultsrised in Table 7.e of the reactors was roughly estimated. Here, thests could play an important role. Due to the highlyatalyst mixture used in methanol carbonylationreactor and also pre-heater, cooler and flash have

from expensive Hastelloy alloy (see Chapter 2.2).ific aspects have been included in the calculation,reactivation has not.

ipment cost summarised in Table 7 shows an ad-EDO process compare to methanol carbonylationP for all production capacities of acetic acid. Thetion on the apparatus cost can be seen on recti-

d pre-heating, the EDO process is more simple.actor price is lower for EDO process mainly dueal and expensive Hastelloy alloy used in methanolon set-up. The costs for a special rectification col-alyst separation (see also Chapter 2.2.) overcomessts and shows a benefit of simple and by material

mited EDO process. On the other hand, the pricesor for EDO process increases dramatically withamount and for 200 kt/year is the compressor a

ributor to the total equipment costs in EDO pro-t 70% of equipment costs).stment costs have been calculated by two differ-

s. In ICARUS calculation, the investment costs de-umber of factors like location, type of controlling,

coil type (the so-called description method was


Table 7Investment costs of major equipment and total investment costs (103 US$) for acetic acid production equipment by EDO and BP methanol carbonylation for10, 50 and 200 kt/year processEquipment Description 10 kt/year 50 kt/year 200 kt/year

EDO BP EDO BP EDO BP

P-6 Compressor 1150 1250 2170 1410 5540 2330H-2 Preheater 45 70 60 110 80 120R-1 Reactor 114 250 130 383 140 450H-4 Cooler 70 130 320 S-5 Flash 50 290 60 560 85 1400C-6 Rectification 540 470 1150 1050 1900 1250C-8 Catalyst separator 1620 4760 8930S-9 Absorber 46 50 46 60 50 60

Equipment cost 2020 4000 3740 8340 8100 14500Investment cost [19] 7020 14000 13020 29030 28200 50630Investment cost [20] 6000 9112 9170 17300 16400 24500

applied). In the module method by Guthrie [19] the invest-ment costs have been obtained multiplying apparatuses pricesby factor 3482. The difference of both calculation methodscomprises the specification of non-direct costs, mainly theprice of special materials used in methanol carbonylation pro-cess. The investment costs calculated by ICARUS [20] arefrom 12 tothe modulecomprehenthe subsequ

4.2. Produ

The inpTable 8.

Five macalculation

(i) Projec(ii) Opera

Table 8Input data for

Parameter ription

Number of ho investmDesired rate o on invesNumber of peTax rate (%)Project capitaProduct escalRaw materialOperating andUtilities escalWorking capi

Operating cha

Plant overhea

G and A expe

(iii) Raw material cost(iv) Utility cost

where the project cost is a fixed cost and all other costs(operating, utility, . . .) are annually based costs, i.e. are cal-culated per year and specified production of acetic acid. The

ct cost (i) means the investment to the plant unit, priceound,ts therial (ii(ethans theuel.e assuent a

umma

nder alationare li

ted fro50% lower compared to the costs resulted frommethod. The ICARUS method processes more

sive information and will be, therefore, used inent calculation of production costs.

ction costs

ut file for ICARUS simulation is illustrated in

in parameter were compared by production cost:

t costting cost

cost estimation

Data Parameter desc

urs 8000 Term used forf return (%) 20 Rate of return

projeof grresen

mateusedmean

and fTh

vestmare s

Ucalcuplantadopriods for analysis 20 Number of periods inc40 Percentage of earnings

l escalation (%) 5 Rate at which projectation (%) 5 Rate at which the saleescalation (%) 3.5 Rate at which the ethamaintenance lab escalation (%) 3 Rate at which the oper

ation (%) 3 Rates at which the utiltal percentage (%) 5 Amount required to op

cient to cover costs exrges (%) 25 Supplies and laborato

per yeard (%) 50 Charges during produc

a percentage of operatnses (%) 8 General and administr

salaries/expenses, proda percentage of operatent analysistment in percent per year

project, buildings, etc. Operating cost (ii) rep-lab, staff, analysis and control spendings. Raw

i) is here defined as a price of a feed componentse and oxygen, methanol and CO). Utility cost (iv)

investment to the steam, cooling water, electricity

mptions used in calculation are related to the in-nalysis and description of the parameters [21] andrised in Table 9.ssumptions from Table 8 the production cost

s were performed. The results for 10200 kt/yearsted in Table 9. The prices of raw material werem literature [2224] and added into the calcula-luded in the cashflow and other project totals and calculationsbefore taxes per year that must be paid to the government

capital expenses may increase expressed in percent per years revenue escalates per yearne price escalates per yearating and maintenance costs escalate per yearities costs escalate per yearerate the facility until the revenue from product sales is suffi-

pressed as a percentage of total capital expense per yearry charges expressed as a percentage of operating labor costs

tion for services, facilities, payroll overhead etc. expressed asing labor and maintenance costs per yearative costs incurred during production such as administrativeuct distribution, research and development etc. expressed as

ing costs per year


Table 9The cost estimation (103 US$) of EDO set-up (F) and commercial BP pro-cess for capacity of the unit 10200 kt/year acetic acid, the conditions seeTables 69

F BP

10 kt/yearInvestment cost 6000 8960Operating cost 3500 3530Raw materials cost 1080 1200Utilities cost 470 210Acetic acid price ($/kg) 0.51 0.49



tion: ethaneprice of C50 US$/m.t

The resument, utilitof 10200 krect oxidatiand separacarbonylatiover, the etsteel and tsame unitfrom expenthe cheape

the same of methanol carbonylation set-up (all comparedvariants). The variant F has a disadvantage, the non-reactedethane and CO2 must be recycled and reused as a feedgas.Thus, the utility cost is for the variant F in all productioncapacity higher than for methanol carbonylation. On con-trary, the utility cost plays not the most important role in totalcosts. Depending on the unit capacity, the utility cost repre-sents from 10% rel. for capacity 10 kt/year up to 15% rel. forcapacity 200 kt acetic acid per year of total annual costs.

The investment and production cost calculation resultsin an estimation of the acetic acid price in all variants andproductions discussed above. The price of acetic acid con-sists of utility, raw material and operating costs and de-creases with the capacity of the unit. In case of both EDOand methanol carbonylation process (50 kt/year) representsnearly 0.36 $/kg. This price is two times lower then the ac-tual market price 0.7 $/kg [6]. For the highest acetic acidproduction 200 kt/year mentioned in this study, the price ofacetic acid is decreasing to 0.33 $/kg, the price for acetic acidcalculated for methanol carbonylation is even lower, 0.3 $/kg.The price of acetic acid must be additionally related to theinvestment costs, which are in all cases higher for traditional

c acid technology.e impto theof pro

ase ofts to ths even

the acquarte2000in Figt threeer 200rise cprice: 72 US$/m.t., methanol: 166 US$/m.t. TheO and oxygen is almost similar, approximately.lts in Table 9 show the relation between invest-

ies, raw material and operating costs. In the caset/year unit, the advantage of process set-up F (di-on of ethane to acetic acid) is shown. The reactiontion unit is more simple compare to the methanolon and therefore the operating cost is lower. More-hane oxidation unit is made from normal stainlesshe investment cost is thus lower compare to thecapacity based on methanol carbonylation madesive Hastelloy (for capacity 10 kt/year). Due to

r feedstock, the raw material cost is lower than

acetiTh

(duetagesincreresulcome

over,

nextfromlistedin lasquartprice0

0,1

0,2

0,3

0,4

0,5

0,6

0,7

0,8

0,9

I/200

0

II/200

0

III/20

00

IV/20

00I/2

001

II/200

1

III/20

01

IV/20

01I/2

002

II/200

2

III/20

02

IVQuarter/year

AA

Pric

e ($/

kg)

Fig. 3. Purchasing prices of acetic acid [/2002

I/200

3

II/200

3

III/20

03

IV/20

03

outlook

6,23].

roved selectivity of acetic acid in EDO processcatalyst optimisation) can bring further advan-cess economy to EDO process; i.e. the selectivity90% calculated for EDO production of 50 kt/yeare acetic acid price of 0.3 $/kg and EDO process be-more interesting for industrial application. More-

etic acid prices seems to stay at about 0.70 $/kg inrs. The acetic acid purchasing prices for the periodto the end 2002 with the forecast for year 2003 are. 3. The purchasing price of acetic acid decreasesquarters from its maximum of 0.79 $/kg in third

1 to actual 0.680.70 $/kg. For year 2003, a slightould be expected to 0.720.73 $/kg of acetic acid.


5. Conclusions

The most attractive variant of the EDO reaction and sep-aration set-up from energy point of view seems to be avariant F, where the water-free reaction mixture consists ofethane, oxygen and CO2. Considering the ethane consump-tion per mole acetic acid as the major objective, the variantC with water in the feed is probably the best one. In gen-eral, the addition of high amount of the water to the feedis not recommended, since the separation effort increasessignificantly.

The comparison of variant F of ethane oxidation to aceticacid with the commercial BP carbonylation of methanolshows that the new technology based on selective oxidationof ethane can compete with the existing processes for aceticacid production.

From the cost analysis the following aspects can beconcluded: acetic acid price is lower in methanol carbony-lation compared to EDO in all cases, but the investmentcosts in state-of-the-art methanol carbonylation process atthe same acetic acid quality are much higher due to theusage of a special resistant materials for construction ofthe plant. If the selectivity to acetic acid can be increasedin EDO from 80 to 90%, the ethane direct oxidationbecomes to be even more attractive for industrial appli-cation.

Acknowled

The studfor EducatiDepartmen

References

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hemicaww.pugment

y was supported by the German Federal Ministryon and Research (Project FKZ03C3013) and thet of Science of the State of Berlin.

[19] GE

[20] A[21] P

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Energetic and economic evaluation of the production of acetic acid via ethane oxidationIntroductionProcess descriptionEthane oxidationReactor designProcess designMode of process simulation for ethane oxidation

Methanol carbonylation

Results and discussionProcess simulation resultsInfluence of the mode of dilution on the energy balance and ethane consumption per mole acetic acidInfluence of the water concentration in the feed on the energy balance and ethane consumption per mole acetic acidComparison of the EDO and BP process

Cost analysisInvestment costProduction costs

ConclusionsAcknowledgmentReferences

1-s2.0-s0255270104001461-main

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