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Energetic and economic evaluation of the production of acetic acid via ethane oxidation
Chemical Engineering and Processing 44 (2005) 421–428
Energetic and economic evaluation of the production of acetic acid via ethane oxidation Q. Smejkal∗ , D. Linke, M. Baerns Institut f¨ur Angewandte Chemie Berlin-Adlershof e.V., Richard-Willst¨atter-Str. 12, D-12489 Berlin, Germany Received 10 April 2003; received in revised form 15 September 2003; accepted 4 June 2004 Available online 19 August 2004
Abstract Acetic acid production via the selective oxidation of ethane was studied. The feed composition and mode of dilution was taken as a major parameter in reactor and process simulation. The concentration of water (as a component improving acetic acid selectivity) in the reaction feed was varied. Heat and mass balances were predicted. Finally, the ethane direct oxidation process was compared to acetic acid production by methanol carbonylation and the investment and production costs are discussed. © 2004 Elsevier B.V. All rights reserved. Keywords: Acetic acid; Ethane oxidation; Reactor simulation; Process flowsheet
1. Introduction The oxidative conversion of light hydrocarbons to more valuable products is of high industrial interest. Acetic acid is an important industrial product with world-wide production over eight million tons. The direct oxidation of ethane to acetic acid can be an alternative to methanol carbonylation process because of its high selectivity and because of the cheap feedstock. The main target of the novel acetic acid process studies is to find process conditions, where the process economy can compete with the state-of-the-art methanol carbonylation process at the same acetic acid quality. The methanol carbonylation process was developed in 1913 and realised (with pioneer cobalt iodide catalyst) by BASF in 1960 in Ludwigshafen, Germany. The high-pressure and high-temperature process was improved by Monsanto and in 1970 commercialised for a production of 135 kt/year in Texas City, Texas (with a new iodide-promoted rhodium catalyst). The operating conditions (30 bar and 453 K) were much milder then in the BASF set-up [1]. Nowadays, there are two leaders in acetic acid production, that is Celanese and ∗
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BP. The BP process uses novel irridium or rhodium catalyst [2,3]. The Celanese process is based on the rhodium catalyst [4]. The world leading acetic acid producers are listed in Table 1. In present study, the set-up of BP was simulated and compared to the ethane oxidation process due to the more detailed references and literature published. The BP process could be taken as a successor of the original Monsanto process. The ethane direct oxidation (EDO) process is not yet commercialised, although SABIC announced in July 2002 a 30 kt/year acetic acid plant based on the EDO process, which has been developed by their Research and Development Center in Riyadh, Saudi Arabia [5,6]. In the present study, acetic acid production by ethane direct oxidation with respect to feed composition and reactor performance was studied. The evaluation of the process flowsheet was based on a simulation by ASPEN PLUSTM . The data from previous kinetic modelling [7] were used as a basis for the simulation of a non-isothermal fixed-bed reactor for the catalytic ethane conversion. The prediction of reactor performance as well as cooling down of the product stream and its separation in a series of heat exchangers and rectification columns were implemented into the process simulations. The aim of this paper is to (i) analyse the influence of feed composition, dilution mode and reactor performance and the
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tween 27 and 48 W m−2 K−1 for a stainless steel tube of i.d. = 25 mm. The differential equations of mass and energy balances were solved applying a numerical integration by the GEAR-algorithm implemented in FORTRAN [11].
Table 1 The world leading acetic acid producers (in thousands of m.t./year, [6]) Company
Global capacity
Celanese BP Chemicals Millennium Chemicals Acetex
2065 1175 450 400
2.1.2. Process design The scheme of the reaction and separation set-up is illustrated in Fig. 1 and was based on known technologies for the production of acetic acid [12,13]. It includes cooling of the reaction mixture, flash of non-condensed components and rectification of water–acetic acid mixture with a limitation of water content in acetic acid <0.1 mass%. As shown in Fig. 1, the ethane and oxygen stream is mixed with water steam and compressed and pre-heated up to the reaction conditions (515 K, 16 bar). The chemical reaction occurs in a multi-tubular reactor. Using thicker tubes than i.d. = 25 mm leads to the formation of severe hot spots and runaway of the reactor because of the non-sufficient heat removal from the fixed-bed reactor to the cooling media. The reaction temperature in the reactor is maintained by an additional heat-exchanger for cooling down the molten salt. The product from the reactor is cooled to 303 K in two steps, highdensity (HD) steam is produced in the first step. Formed gas and liquid mixture is divided in a flash. Acetic acid–water mixture is subsequently separated in a rectification column with ca. 40 theoretical stages; pure acetic acid is withdrawn as a bottom product. The gaseous stream consists of non-reacted ethane, ethylene and CO2 . A portion of produced CO2 is separated in an absorber; ethane and ethylene are recycled to the feedgas.
separation effort based on EDO; (ii) demonstrate the features and benefits of direct ethane oxidation to acetic acid compared to the industrial methanol carbonylation process; (iii) predict the investment and production costs.
2. Process description 2.1. Ethane oxidation 2.1.1. Reactor design A multi-channel fixed-bed reactor of 30000 tubes (length 3 m, i.d. 25 mm) was described by a pseudo-homogeneous one-dimensional model for acetic acid production of 50 kt/year. The mass and energy balances for the steady-state are given elsewhere [7,8]. The heat transfer coefficient between the catalyst bed and the wall was derived by the correlation of Hennecke and Schl¨under [9] and Tsotas and Schl¨under [10]. Depending on the reaction conditions (gas velocity, composition of gas feed, pellet size), the heat conductivity of the catalyst bed varied between 0.43 and 0.65 W m−1 K−1 . The respective internal heat transfer coefficients varied be-
CO2 Feed
Recycle
P-1
S-7 R-3
H-4
H-2 H-4a
Water S-5
C-6
Acetic acid
Fig. 1. The scheme of acetic acid processes: ethane oxidation; 1: compressor, 2: pre-heater, 3: multi-tubular fixed-bed reactor cooled by molten salt, 4, 4a: cooler, 5: flash, 6: rectification of water–acetic acid (purification of acetic acid), 7: CO2 separation, cooling 4, 4a: from reaction temperature to 303 K in two steps, p = 16 bar, flash: 303 K, p = 16 bar. Code letters used in Fig. 1: C: rectification column, H: heat exchanger, P: compressor, R: reactor, S: absorber, separator, flash.
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Table 2 Constants of NRTL equation of state used in simulation of acetic acid–water separation
CO2 ). In case of variant F, only a portion of produced CO2 is separated from the recycle stream.
Parameter
Value
2.2. Methanol carbonylation
ai, j aj , i bi, j bj , i ci, j
−1.9763 3.3293 609.8886 −723.8881 0.3
The phase equilibrium acetic acid–water was described using NRTL equation of state. The parameters of the NRTL equation are listed in Table 2 and were taken from ASPEN Plus databank (databank VLE-HOC). 2.1.3. Mode of process simulation for ethane oxidation The calculations were performed for a constant plant scale production 50 kt/year for each process variant discussed further below. The feed composition, conversion of ethane and selectivity of acetic acid were obtained from the simulation of a polytropic fixed-bed reactor and were used for the ASPEN PLUSTM simulation. Therefore, the optimal temperature range with respect to thermally stable reactor operation was derived, that is feed temperature 515 K and temperature of the cooling medium 515 K. The maximal tolerable reactor temperature in the hot spot was set to 550 K. The amount of recycled ethane was obtained by fitting the predicted product composition obtained by polytropic fixed-bed reactor simulation to the results of process simulation for which a so-called black-box reactor unit was used. The variants of the ethane oxidation are summarised in Table 3. Variants A and B refer to a water-free feed. Here, the influence of oxygen concentration was studied. Since all fixed-bed reactor simulations with coxygen > 14 mole % led to runaway of the reactor the oxygen concentration was limited to 13 mole % in the feed. The comparison of variants A, C and D shows the influence of water on the energy balance of the process and, especially, on the separation effort. A positive effect of water addition to the feed on the selectivity to acetic acid is well known for the ethane oxidation reaction [7]. In variant E, the influence of air as an oxidising agent in the presence of water in the feed is considered. Variants A, B, E and F deal with the comparison of different dilution agents (ethane, nitrogen and Table 3 Description of the EDO to acetic acid process variants, GHSV = 1200 h−1 , Tfeed = 515 K, p = 16 bar, calculated for 50 kt/year production of acetic acid Variant
A B C D Ea F a
Molar fraction Ethane
Oxygen
Water
Nitrogena
CO2
0.87 0.95 0.8 0.85 0.85 0.73
0.13 0.05 0.08 0.1 0.02 0.12
0 0 0.12 0.05 0.05 0
0 0 0 0 0.08 0
0 0 0 0 0 0.15
When using nitrogen as dilution agent.
For methanol carbonylation, the liquid methanol and homogeneous catalyst in form of methyl iodide complex (Rh(CO)x ) is mixed (Fig. 2) with a gaseous feed (compressed CO) and pre-heated to the reaction conditions [3,14] – 462 K, 30 bar, two phase system. The carbonylation in the stirred tank reactor is followed by separation of gas and liquid phases in a flash (5). From the liquid phase, the homogeneous catalyst is separated by rectification (7) and recycled back to the feed. In a small rectification column (8), a portion of CO is separated and consequently the water and acetic acid mixture is distilled in the rectification column (6) with efficiency 47 theoretical stages. Gas phase from flash consists of CO and CO2 ; after purification (9: CO2 selective absorption) CO is recycled back to the feedgas. The process specification is as follows: 33.4 wt.% methanol and 31 wt.% carbon monoxide were fed continuously into the reactor held at temperature of 462 K. The average composition of the reactor contents was: 2.6 wt.% methyl acetate, 5.6 wt.% water, 14.0 wt.% methyl iodide, 61.9 wt.% acetic acid, 0.55 wt.% lithium (present at least in part as iodide salt), 550 ppm rhodium and 11.6 wt.% iodide. The lithium iodide in the reactor composition functioned both as a carbonylation catalyst stabiliser in the reactor and a water relative volatility suppressant in the flash tank. In the simulation, the homogeneous catalyst was recycled back to the reaction feed after separation from the reaction mixture (8) [14]. Due to the highly corrosive catalyst mixture used in the process parts of the unit (at least pre-heater, reactor, cooler and catalyst separator) have to be made from Hastelloy alloy [15,16].
3. Results and discussion 3.1. Process simulation results The results of the process simulation for the different process variants (see Table 2) for the ethane oxidation process are summarised in Tables 3 and 4. The energy consumed and produced in the process is shown in Table 2. Generally, the higher the heat consumption of the process variant, the higher are operating and utility costs of the process. The heat produced by the reaction or the HD steam formed in cooler H-4 can be reused in other process operations, i.e. for pre-heating of the feed or for heating the rectification column. 3.1.1. Influence of the mode of dilution on the energy balance and ethane consumption per mole acetic acid For analysis of the influence of dilution mode on the energy balance, variants A, B, E and F were considered (conditions see Table 3). The advantage of the process variant A
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Feed CO2 Recycle P-1
Water
S-9
H-2
C-6 S-5
H-4
CST-3
Acetic acid Catalyst recycle C-7 C-8
Fig. 2. The scheme of acetic acid processes: methanol carbonylation by BP; 1: compressor, 2: heating, 3: CST reactor, 4: cooler, 5: flash, 6: rectification of water–acetic acid (purification of acetic acid), 7: separation of homogeneous catalyst, 8: pre-separation of reaction mixture, 9: CO2 separation, cooling: from reaction temperature to 303 K in one step, p = 30 bar, flash: 303 K, p = 30 bar. Code letters used in Fig. 1: C: rectification column, H: heat exchanger, P: compressor, CST: reactor, S: absorber, separator, flash.
and B (where only ethane and oxygen are fed) compared to variant E with air used as an oxidising agent can be deduced from Table 3. Air as an oxidising agent is not favourable due to the extended effort for compression and pre-heating of the reaction mixture (variant E). For variant A and B, lower compression and separation energy consumption is predicted for variant A with higher oxygen concentration in the feed. Variant B corresponds to the variant with 95% ethane and 5% oxygen in the feed; i.e. the oxygen content was decreased compared to variant A. Obviously, this leads to an increase of energy consumption. The variant F of EDO process, where the dilution of the reaction mixture is realised by CO2 is the most promising variant among all process variants discussed above. The com-
pression and heating/cooling effort for EDO variant F with CO2 is lower compared to variant A with only ethane and oxygen in the feed due to the lower heat capacity of carbon dioxide. Moreover, for nearly the same concentration of oxygen in the feed, the CO2 seems to be better dilution agent than ethane (compare variants A and F). 3.1.2. Influence of the water concentration in the feed on the energy balance and ethane consumption per mole acetic acid Variant C represents the highest concentration of water in the feed (13 mole %). The positive influence of water on the acetic acid selectivity is obvious from Table 5. The comparison of variant C with variant A, D and F shows, that
Table 4 Summary results of the computer simulation of variants A–F by EDO process, energy balance of ethane oxidation to acetic acid and BP variant based on methanol carbonylation (see Chapter 3.1.3), 50 kt/year production of acetic acid (HD: high density steam) Variant
A B C D E F BP
Heat consumption (kW)
Heat production (kW)
Compressor
Heater
Cooler
Rector
Total
Reactor
HD steam
Total
6120 14000 10000 8500 42000 5100
1400 3630 7800 2300 16000 600
7340 14050 16600 12100 45000 6600
7400 7200 21000 11250 27000 7400
22300 38900 55400 34150 130000 19700
29000 25200 26500 29300 22400 26600
2300 5600 4100 3500 17000 1800
31300 30800 30600 32800 39400 28400
670
3600
2900
13000
20500
7600
0
7600
Q. Smejkal et al. / Chemical Engineering and Processing 44 (2005) 421–428 Table 5 Selectivity and ethane consumption of variants A–F, EDO process, TFeed = 515 K, p = 16 bar, 50 kt/year production of acetic acid, 100% conversion of oxygen Variant
Selectivity of acetic acid (%)
Consumption of ethane (mol)/mol acetic acid
A B C D E F
76.3 77.7 81.0 79.6 80.7 76.1
1.31 1.29 1.23 1.25 1.24 1.31
the increase of the selectivity to acetic acid results to lower consumption of ethane per formed unit of ethylene (see Table 5). On the contrary, the separation effort is increasing with increasing amount of water. Taking into account the antagonistic requirements – high water content to improve the selectivity to acetic acid on one hand, and low water concentration to minimise the separation effort on the other hand – variant F (73 mole % ethane, 12 mole % oxygen, 15% CO2 ) seems to be a best variant of EDO (see Table 4).
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Table 6 Design assumptions for acetic acid production calculated for year 2002 Parameter
Value/type
Capacity Procedure
10–200 kt/year acetic acid (a) EDO (b) Methanol carbonylation
Process location Process control Type of process Soil condition Pressure vessel design code
Europe Digital Grass roots/clear field Soft clay DIN
acid process was used located in Europe and designed according to DIN. The process flowsheet for ethane oxidation and methanol carbonylation was adopted for the cost analysis. The plant unit results discussed above show the advantage of variant F of EDO, where CO2 is used as a dilution agent. The results from the flowsheet simulation were used as an input file for cost prediction routine ICARUS by ASPEN PLUSTM . 4.1. Investment cost
3.1.3. Comparison of the EDO and BP process The comparison of variant F of selective oxidation of ethane with the commercial BP methanol carbonylation processes is documented in Table 4 for an acetic acid production plant of 50 kt/year. Among the ethane oxidation processes, the set-up with the highest economic potential is represented by variant F with a feed composition 73 mole % ethane, 12 mole % oxygen and 15 mole % CO2 a cooling temperature and feed temperature of 515 K. The energy consumption for compression and product cooling is the lowest in the case of methanol carbonylation. On the contrary, the effort for the pre-heating and product purification (rectification) is lower in case of EDO variant F; the heat produced by the chemical reaction is higher in the EDO process compared to methanol carbonylation. Moreover, the high-density steam can be produced only in case of the EDO process (see Table 4).
4. Cost analysis The objective of this part of the study is to evaluate the economics of ethane direct oxidation compared to industry relevant methanol carbonylation process by BP [3,14]. The total investment and production costs of acetic acid consist of: (1) fixed (investment) costs, including apparatuses, project, buildings, labor, analysis and control [17,18] (2) production costs, i.e. variable costs, including the raw material ethane and oxygen (EDO process) and methanol and CO (BP process) price and the utilities – steam, gases, catalyst, cooling water, fuel, energy, etc. The basic assumptions used in process cost estimation are given in Table 6. A low to high-scale 10–200 kt/year acetic
The block flowseet for ethane direct oxidation and methanol carbonylation is identical with process schemes shown in Figs. 1 and 2. The total investment costs are based on total equipment costs. Therefore, the first step of cost estimation comprises the calculation of the purchase costs of apparatuses included in the process simulation. The results are summarised in Table 7. The price of the reactors was roughly estimated. Here, the material costs could play an important role. Due to the highly corrosive catalyst mixture used in methanol carbonylation process, the reactor and also pre-heater, cooler and flash have to be made from expensive Hastelloy alloy (see Chapter 2.2). These specific aspects have been included in the calculation, the catalyst reactivation has not. The equipment cost summarised in Table 7 shows an advantage of EDO process compare to methanol carbonylation set-up by BP for all production capacities of acetic acid. The main reduction on the apparatus cost can be seen on rectification and pre-heating, the EDO process is more simple. Also the reactor price is lower for EDO process mainly due to the special and expensive Hastelloy alloy used in methanol carbonylation set-up. The costs for a special rectification column for catalyst separation (see also Chapter 2.2.) overcomes all other costs and shows a benefit of simple and by material used non-limited EDO process. On the other hand, the price of compressor for EDO process increases dramatically with production amount and for 200 kt/year is the compressor a major contributor to the total equipment costs in EDO process (almost 70% of equipment costs). The investment costs have been calculated by two different methods. In ICARUS calculation, the investment costs depend on a number of factors like location, type of controlling, process and coil type (the so-called description method was
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Table 7 Investment costs of major equipment and total investment costs (103 US$) for acetic acid production equipment by EDO and BP methanol carbonylation for 10, 50 and 200 kt/year process Equipment
Description
P-6 H-2 R-1 H-4 S-5 C-6 C-8 S-9
Compressor Preheater Reactor Cooler Flash Rectification Catalyst separator Absorber
10 kt/year EDO
Equipment cost Investment cost [19] Investment cost [20]
50 kt/year BP
EDO
200 kt/year BP
EDO
BP
1150 45 114 70 50 540 – 46
1250 70 250 – 290 470 1620 50
2170 60 130 130 60 1150 – 46
1410 110 383 – 560 1050 4760 60
5540 80 140 320 85 1900 – 50
2330 120 450 – 1400 1250 8930 60
2020 7020 6000
4000 14000 9112
3740 13020 9170
8340 29030 17300
8100 28200 16400
14500 50630 24500
applied). In the “module method” by Guthrie [19] the investment costs have been obtained multiplying apparatuses prices by factor 3482. The difference of both calculation methods comprises the specification of non-direct costs, mainly the price of special materials used in methanol carbonylation process. The investment costs calculated by ICARUS [20] are from 12 to 50% lower compared to the costs resulted from the module method. The ICARUS method processes more comprehensive information and will be, therefore, used in the subsequent calculation of production costs. 4.2. Production costs The input file for ICARUS simulation is illustrated in Table 8. Five main parameter were compared by production cost calculation: (i) Project cost (ii) Operating cost
(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 calculated per year and specified production of acetic acid. The project cost (i) means the investment to the plant unit, price of ground, project, buildings, etc. Operating cost (ii) represents the lab, staff, analysis and control spendings. Raw material (iii) is here defined as a price of a feed components used (ethane and oxygen, methanol and CO). Utility cost (iv) means the investment to the steam, cooling water, electricity and fuel. The assumptions used in calculation are related to the investment analysis and description of the parameters [21] and are summarised in Table 9. Under assumptions from Table 8 the production cost calculations were performed. The results for 10–200 kt/year plant are listed in Table 9. The prices of raw material were adopted from literature [22–24] and added into the calcula-
Table 8 Input data for cost estimation Parameter
Data
Parameter description
Number of hours Desired rate of return (%) Number of periods for analysis Tax rate (%) Project capital escalation (%) Product escalation (%) Raw material escalation (%) Operating and maintenance lab escalation (%) Utilities escalation (%) Working capital percentage (%)
8000 20 20 40 5 5 3.5 3 3 5
Term used for investment analysis Rate of return on investment in percent per year Number of periods included in the cashflow and other project totals and calculations Percentage of earnings before taxes per year that must be paid to the government Rate at which project capital expenses may increase expressed in percent per year Rate at which the sales revenue escalates per year Rate at which the ethane price escalates per year Rate at which the operating and maintenance costs escalate per year Rates at which the utilities costs escalate per year Amount required to operate the facility until the revenue from product sales is sufficient to cover costs expressed as a percentage of total capital expense per year Supplies and laboratory charges expressed as a percentage of operating labor costs per year Charges during production for services, facilities, payroll overhead etc. expressed as a percentage of operating labor and maintenance costs per year General and administrative costs incurred during production such as administrative salaries/expenses, product distribution, research and development etc. expressed as a percentage of operating costs per year
Operating charges (%)
25
Plant overhead (%)
50
G and A expenses (%)
8
Q. Smejkal et al. / Chemical Engineering and Processing 44 (2005) 421–428 Table 9 The cost estimation (103 US$) of EDO set-up (F) and commercial BP process for capacity of the unit 10–200 kt/year acetic acid, the conditions see Tables 6–9 F
BP
10 kt/year Investment cost Operating cost Raw materials cost Utilities cost Acetic acid price ($/kg)
6000 3500 1080 470 0.51
8960 3530 1200 210 0.49
50 kt/year Investment cost Operating cost Raw materials cost Utilities cost Acetic acid price ($/kg)
9170 10400 5300 2500 0.36
17300 10400 6000 1200 0.35
200 kt/year Investment cost Operating cost Raw materials cost Utilities cost Acetic acid price ($/kg)
16400 35700 21000 9830 0.33
23500 31700 24000 2650 0.29
tion: ethane price: 72 US$/m.t., methanol: 166 US$/m.t. The price of CO and oxygen is almost similar, approximately 50 US$/m.t. The results in Table 9 show the relation between investment, utilities, raw material and operating costs. In the case of 10–200 kt/year unit, the advantage of process set-up F (direct oxidation of ethane to acetic acid) is shown. The reaction and separation unit is more simple compare to the methanol carbonylation and therefore the operating cost is lower. Moreover, the ethane oxidation unit is made from normal stainless steel and the investment cost is thus lower compare to the same unit capacity based on methanol carbonylation made from expensive Hastelloy (for capacity 10 kt/year). Due to the cheaper feedstock, the raw material cost is lower than
the same of methanol carbonylation set-up (all compared variants). The variant F has a disadvantage, the non-reacted ethane and CO2 must be recycled and reused as a feedgas. Thus, the utility cost is for the variant F in all production capacity higher than for methanol carbonylation. On contrary, the utility cost plays not the most important role in total costs. Depending on the unit capacity, the utility cost represents from 10% rel. for capacity 10 kt/year up to 15% rel. for capacity 200 kt acetic acid per year of total annual costs. The investment and production cost calculation results in an estimation of the acetic acid price in all variants and productions discussed above. The price of acetic acid consists of utility, raw material and operating costs and decreases with the capacity of the unit. In case of both EDO and methanol carbonylation process (50 kt/year) represents nearly 0.36 $/kg. This price is two times lower then the actual market price 0.7 $/kg [6]. For the highest acetic acid production 200 kt/year mentioned in this study, the price of acetic acid is decreasing to 0.33 $/kg, the price for acetic acid calculated for methanol carbonylation is even lower, 0.3 $/kg. The price of acetic acid must be additionally related to the investment costs, which are in all cases higher for traditional acetic acid technology. The improved selectivity of acetic acid in EDO process (due to the catalyst optimisation) can bring further advantages of process economy to EDO process; i.e. the selectivity increase of 90% calculated for EDO production of 50 kt/year results to the acetic acid price of 0.3 $/kg and EDO process becomes even more interesting for industrial application. Moreover, the acetic acid prices seems to stay at about 0.70 $/kg in next quarters. The acetic acid purchasing prices for the period from 2000 to the end 2002 with the forecast for year 2003 are listed in Fig. 3. The purchasing price of acetic acid decreases in last three quarters from its maximum of 0.79 $/kg in third quarter 2001 to actual 0.68–0.70 $/kg. For year 2003, a slight price rise could be expected to 0.72–0.73 $/kg of acetic acid.
0,9 outlook
0,8 0,7 0,6 0,5 0,4 0,3 0,2 0,1
0 II/ 0 0 20 III 00 /2 0 IV 0 0 /2 00 I/2 0 0 II/ 0 1 20 III 01 /2 0 IV 0 1 /2 00 I/2 1 00 II/ 2 20 III 02 /2 0 IV 0 2 /2 00 I/2 2 0 II/ 0 3 20 III 03 /2 0 IV 0 3 /2 00 3
0 I/2
AA Price ($/kg)
427
Quarter/year Fig. 3. Purchasing prices of acetic acid [6,23].
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5. Conclusions
References
The most attractive variant of the EDO reaction and separation set-up from energy point of view seems to be a variant F, where the water-free reaction mixture consists of ethane, oxygen and CO2 . Considering the ethane consumption per mole acetic acid as the major objective, the variant C with water in the feed is probably the best one. In general, the addition of high amount of the water to the feed is not recommended, since the separation effort increases significantly. The comparison of variant F of ethane oxidation to acetic acid with the commercial BP carbonylation of methanol shows that the new technology based on selective oxidation of ethane can compete with the existing processes for acetic acid production. From the cost analysis the following aspects can be concluded: acetic acid price is lower in methanol carbonylation compared to EDO in all cases, but the investment costs in state-of-the-art methanol carbonylation process at the same acetic acid quality are much higher due to the usage of a special resistant materials for construction of the plant. If the selectivity to acetic acid can be increased in EDO from 80 to 90%, the ethane direct oxidation becomes to be even more attractive for industrial application.
[1] Ullmann’s Encyclopedy of Chemical Technology, vol. A1, Wiley, NY, 1985, pp. 47–50. [2] L.A. Key, D.J. Law, US 6472558 (BP), 2002. [3] S. Aubigne, G.B. Cooper, B.L. Williams, D.J. Waston, US 5416237 (BP), 1994. [4] B.L. Smith, et al., US 5144068 (Celanese), 1992. [5] Chemical Week, July 2002. [6] Chemical Week, 1999–2003. [7] D. Linke, et al., Chem. Eng. Sci. 52 (1) (2002) 32–43. [8] G.F. Froment, K.B. Bischoff, Chemical Reactor Analysis and Design, Wiley, 1979. [9] F.W. Hennecke, E.-U. Schl¨under, Chem. Ing. Tech. 45 (1974) 277. [10] E. Tsotsas, E.-U. Schl¨under, Chem. Eng. Sci. 45 (4) (1990) 819–837. [11] C.W. Gear, Numerical Initial Value Problems in Ordinary Differential Equations, Prentice-Hall International, New Jersey, 1971. [12] K. Karim, M.H. Al-Hazmi, E. Mamedov, US 5907056, Saudi Basis Industries, 1999. [13] N.C. Benkalowycz, et al., US 5300684, The Standard Oil Comp., 1994. [14] C.S. Gerland, M.F. Giles, J.G. Sunley, US 5672743 (BP), 1994. [15] Perry’s Chemical Engineers’ Handbook, McGraw-Hill, New York 23/16–58, 1984. [16] K.-I. Sano, H. Uchida, S. Wakabayashi, Catal. Surveys Jpn. 3 (1999) 55–60. [17] H.J. Sandler, Practical Process Engineering: A Working Approach to Plant Design, McGraw-Hill, 1987. [18] N. Yoneda, et al., Recent advances in process and catalyst design for the production of acetic acid, Appl. Cat. A General 221 (2001) 253–265. [19] G.D. Ulrich, Guide to Chemical Engineering Process Design and Economics, Wiley, NY, 1984. [20] Aspen ICARUS Process Evaluator User’s Manual, 2002. [21] Peters, Timmerhaus, Plant Design and Economics for Chemical Engineers, third ed., McGraw-Hill, London, 1980. [22] www.methanex.com, June 2002. [23] Chemical Week, February 2002. [24] www.purchasing.com, January 2003.
Acknowledgment The study was supported by the German Federal Ministry for Education and Research (Project FKZ03C3013) and the Department of Science of the State of Berlin.
Energetic and economic evaluation of the production of acetic acid via ethane oxidation Q. Smejkal∗ , D. Linke, M. Baerns Institut f¨ur Angewandte Chemie Berlin-Adlershof e.V., Richard-Willst¨atter-Str. 12, D-12489 Berlin, Germany Received 10 April 2003; received in revised form 15 September 2003; accepted 4 June 2004 Available online 19 August 2004
Abstract Acetic acid production via the selective oxidation of ethane was studied. The feed composition and mode of dilution was taken as a major parameter in reactor and process simulation. The concentration of water (as a component improving acetic acid selectivity) in the reaction feed was varied. Heat and mass balances were predicted. Finally, the ethane direct oxidation process was compared to acetic acid production by methanol carbonylation and the investment and production costs are discussed. © 2004 Elsevier B.V. All rights reserved. Keywords: Acetic acid; Ethane oxidation; Reactor simulation; Process flowsheet
1. Introduction The oxidative conversion of light hydrocarbons to more valuable products is of high industrial interest. Acetic acid is an important industrial product with world-wide production over eight million tons. The direct oxidation of ethane to acetic acid can be an alternative to methanol carbonylation process because of its high selectivity and because of the cheap feedstock. The main target of the novel acetic acid process studies is to find process conditions, where the process economy can compete with the state-of-the-art methanol carbonylation process at the same acetic acid quality. The methanol carbonylation process was developed in 1913 and realised (with pioneer cobalt iodide catalyst) by BASF in 1960 in Ludwigshafen, Germany. The high-pressure and high-temperature process was improved by Monsanto and in 1970 commercialised for a production of 135 kt/year in Texas City, Texas (with a new iodide-promoted rhodium catalyst). The operating conditions (30 bar and 453 K) were much milder then in the BASF set-up [1]. Nowadays, there are two leaders in acetic acid production, that is Celanese and ∗
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BP. The BP process uses novel irridium or rhodium catalyst [2,3]. The Celanese process is based on the rhodium catalyst [4]. The world leading acetic acid producers are listed in Table 1. In present study, the set-up of BP was simulated and compared to the ethane oxidation process due to the more detailed references and literature published. The BP process could be taken as a successor of the original Monsanto process. The ethane direct oxidation (EDO) process is not yet commercialised, although SABIC announced in July 2002 a 30 kt/year acetic acid plant based on the EDO process, which has been developed by their Research and Development Center in Riyadh, Saudi Arabia [5,6]. In the present study, acetic acid production by ethane direct oxidation with respect to feed composition and reactor performance was studied. The evaluation of the process flowsheet was based on a simulation by ASPEN PLUSTM . The data from previous kinetic modelling [7] were used as a basis for the simulation of a non-isothermal fixed-bed reactor for the catalytic ethane conversion. The prediction of reactor performance as well as cooling down of the product stream and its separation in a series of heat exchangers and rectification columns were implemented into the process simulations. The aim of this paper is to (i) analyse the influence of feed composition, dilution mode and reactor performance and the
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tween 27 and 48 W m−2 K−1 for a stainless steel tube of i.d. = 25 mm. The differential equations of mass and energy balances were solved applying a numerical integration by the GEAR-algorithm implemented in FORTRAN [11].
Table 1 The world leading acetic acid producers (in thousands of m.t./year, [6]) Company
Global capacity
Celanese BP Chemicals Millennium Chemicals Acetex
2065 1175 450 400
2.1.2. Process design The scheme of the reaction and separation set-up is illustrated in Fig. 1 and was based on known technologies for the production of acetic acid [12,13]. It includes cooling of the reaction mixture, flash of non-condensed components and rectification of water–acetic acid mixture with a limitation of water content in acetic acid <0.1 mass%. As shown in Fig. 1, the ethane and oxygen stream is mixed with water steam and compressed and pre-heated up to the reaction conditions (515 K, 16 bar). The chemical reaction occurs in a multi-tubular reactor. Using thicker tubes than i.d. = 25 mm leads to the formation of severe hot spots and runaway of the reactor because of the non-sufficient heat removal from the fixed-bed reactor to the cooling media. The reaction temperature in the reactor is maintained by an additional heat-exchanger for cooling down the molten salt. The product from the reactor is cooled to 303 K in two steps, highdensity (HD) steam is produced in the first step. Formed gas and liquid mixture is divided in a flash. Acetic acid–water mixture is subsequently separated in a rectification column with ca. 40 theoretical stages; pure acetic acid is withdrawn as a bottom product. The gaseous stream consists of non-reacted ethane, ethylene and CO2 . A portion of produced CO2 is separated in an absorber; ethane and ethylene are recycled to the feedgas.
separation effort based on EDO; (ii) demonstrate the features and benefits of direct ethane oxidation to acetic acid compared to the industrial methanol carbonylation process; (iii) predict the investment and production costs.
2. Process description 2.1. Ethane oxidation 2.1.1. Reactor design A multi-channel fixed-bed reactor of 30000 tubes (length 3 m, i.d. 25 mm) was described by a pseudo-homogeneous one-dimensional model for acetic acid production of 50 kt/year. The mass and energy balances for the steady-state are given elsewhere [7,8]. The heat transfer coefficient between the catalyst bed and the wall was derived by the correlation of Hennecke and Schl¨under [9] and Tsotas and Schl¨under [10]. Depending on the reaction conditions (gas velocity, composition of gas feed, pellet size), the heat conductivity of the catalyst bed varied between 0.43 and 0.65 W m−1 K−1 . The respective internal heat transfer coefficients varied be-
CO2 Feed
Recycle
P-1
S-7 R-3
H-4
H-2 H-4a
Water S-5
C-6
Acetic acid
Fig. 1. The scheme of acetic acid processes: ethane oxidation; 1: compressor, 2: pre-heater, 3: multi-tubular fixed-bed reactor cooled by molten salt, 4, 4a: cooler, 5: flash, 6: rectification of water–acetic acid (purification of acetic acid), 7: CO2 separation, cooling 4, 4a: from reaction temperature to 303 K in two steps, p = 16 bar, flash: 303 K, p = 16 bar. Code letters used in Fig. 1: C: rectification column, H: heat exchanger, P: compressor, R: reactor, S: absorber, separator, flash.
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Table 2 Constants of NRTL equation of state used in simulation of acetic acid–water separation
CO2 ). In case of variant F, only a portion of produced CO2 is separated from the recycle stream.
Parameter
Value
2.2. Methanol carbonylation
ai, j aj , i bi, j bj , i ci, j
−1.9763 3.3293 609.8886 −723.8881 0.3
The phase equilibrium acetic acid–water was described using NRTL equation of state. The parameters of the NRTL equation are listed in Table 2 and were taken from ASPEN Plus databank (databank VLE-HOC). 2.1.3. Mode of process simulation for ethane oxidation The calculations were performed for a constant plant scale production 50 kt/year for each process variant discussed further below. The feed composition, conversion of ethane and selectivity of acetic acid were obtained from the simulation of a polytropic fixed-bed reactor and were used for the ASPEN PLUSTM simulation. Therefore, the optimal temperature range with respect to thermally stable reactor operation was derived, that is feed temperature 515 K and temperature of the cooling medium 515 K. The maximal tolerable reactor temperature in the hot spot was set to 550 K. The amount of recycled ethane was obtained by fitting the predicted product composition obtained by polytropic fixed-bed reactor simulation to the results of process simulation for which a so-called black-box reactor unit was used. The variants of the ethane oxidation are summarised in Table 3. Variants A and B refer to a water-free feed. Here, the influence of oxygen concentration was studied. Since all fixed-bed reactor simulations with coxygen > 14 mole % led to runaway of the reactor the oxygen concentration was limited to 13 mole % in the feed. The comparison of variants A, C and D shows the influence of water on the energy balance of the process and, especially, on the separation effort. A positive effect of water addition to the feed on the selectivity to acetic acid is well known for the ethane oxidation reaction [7]. In variant E, the influence of air as an oxidising agent in the presence of water in the feed is considered. Variants A, B, E and F deal with the comparison of different dilution agents (ethane, nitrogen and Table 3 Description of the EDO to acetic acid process variants, GHSV = 1200 h−1 , Tfeed = 515 K, p = 16 bar, calculated for 50 kt/year production of acetic acid Variant
A B C D Ea F a
Molar fraction Ethane
Oxygen
Water
Nitrogena
CO2
0.87 0.95 0.8 0.85 0.85 0.73
0.13 0.05 0.08 0.1 0.02 0.12
0 0 0.12 0.05 0.05 0
0 0 0 0 0.08 0
0 0 0 0 0 0.15
When using nitrogen as dilution agent.
For methanol carbonylation, the liquid methanol and homogeneous catalyst in form of methyl iodide complex (Rh(CO)x ) is mixed (Fig. 2) with a gaseous feed (compressed CO) and pre-heated to the reaction conditions [3,14] – 462 K, 30 bar, two phase system. The carbonylation in the stirred tank reactor is followed by separation of gas and liquid phases in a flash (5). From the liquid phase, the homogeneous catalyst is separated by rectification (7) and recycled back to the feed. In a small rectification column (8), a portion of CO is separated and consequently the water and acetic acid mixture is distilled in the rectification column (6) with efficiency 47 theoretical stages. Gas phase from flash consists of CO and CO2 ; after purification (9: CO2 selective absorption) CO is recycled back to the feedgas. The process specification is as follows: 33.4 wt.% methanol and 31 wt.% carbon monoxide were fed continuously into the reactor held at temperature of 462 K. The average composition of the reactor contents was: 2.6 wt.% methyl acetate, 5.6 wt.% water, 14.0 wt.% methyl iodide, 61.9 wt.% acetic acid, 0.55 wt.% lithium (present at least in part as iodide salt), 550 ppm rhodium and 11.6 wt.% iodide. The lithium iodide in the reactor composition functioned both as a carbonylation catalyst stabiliser in the reactor and a water relative volatility suppressant in the flash tank. In the simulation, the homogeneous catalyst was recycled back to the reaction feed after separation from the reaction mixture (8) [14]. Due to the highly corrosive catalyst mixture used in the process parts of the unit (at least pre-heater, reactor, cooler and catalyst separator) have to be made from Hastelloy alloy [15,16].
3. Results and discussion 3.1. Process simulation results The results of the process simulation for the different process variants (see Table 2) for the ethane oxidation process are summarised in Tables 3 and 4. The energy consumed and produced in the process is shown in Table 2. Generally, the higher the heat consumption of the process variant, the higher are operating and utility costs of the process. The heat produced by the reaction or the HD steam formed in cooler H-4 can be reused in other process operations, i.e. for pre-heating of the feed or for heating the rectification column. 3.1.1. Influence of the mode of dilution on the energy balance and ethane consumption per mole acetic acid For analysis of the influence of dilution mode on the energy balance, variants A, B, E and F were considered (conditions see Table 3). The advantage of the process variant A
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Feed CO2 Recycle P-1
Water
S-9
H-2
C-6 S-5
H-4
CST-3
Acetic acid Catalyst recycle C-7 C-8
Fig. 2. The scheme of acetic acid processes: methanol carbonylation by BP; 1: compressor, 2: heating, 3: CST reactor, 4: cooler, 5: flash, 6: rectification of water–acetic acid (purification of acetic acid), 7: separation of homogeneous catalyst, 8: pre-separation of reaction mixture, 9: CO2 separation, cooling: from reaction temperature to 303 K in one step, p = 30 bar, flash: 303 K, p = 30 bar. Code letters used in Fig. 1: C: rectification column, H: heat exchanger, P: compressor, CST: reactor, S: absorber, separator, flash.
and B (where only ethane and oxygen are fed) compared to variant E with air used as an oxidising agent can be deduced from Table 3. Air as an oxidising agent is not favourable due to the extended effort for compression and pre-heating of the reaction mixture (variant E). For variant A and B, lower compression and separation energy consumption is predicted for variant A with higher oxygen concentration in the feed. Variant B corresponds to the variant with 95% ethane and 5% oxygen in the feed; i.e. the oxygen content was decreased compared to variant A. Obviously, this leads to an increase of energy consumption. The variant F of EDO process, where the dilution of the reaction mixture is realised by CO2 is the most promising variant among all process variants discussed above. The com-
pression and heating/cooling effort for EDO variant F with CO2 is lower compared to variant A with only ethane and oxygen in the feed due to the lower heat capacity of carbon dioxide. Moreover, for nearly the same concentration of oxygen in the feed, the CO2 seems to be better dilution agent than ethane (compare variants A and F). 3.1.2. Influence of the water concentration in the feed on the energy balance and ethane consumption per mole acetic acid Variant C represents the highest concentration of water in the feed (13 mole %). The positive influence of water on the acetic acid selectivity is obvious from Table 5. The comparison of variant C with variant A, D and F shows, that
Table 4 Summary results of the computer simulation of variants A–F by EDO process, energy balance of ethane oxidation to acetic acid and BP variant based on methanol carbonylation (see Chapter 3.1.3), 50 kt/year production of acetic acid (HD: high density steam) Variant
A B C D E F BP
Heat consumption (kW)
Heat production (kW)
Compressor
Heater
Cooler
Rector
Total
Reactor
HD steam
Total
6120 14000 10000 8500 42000 5100
1400 3630 7800 2300 16000 600
7340 14050 16600 12100 45000 6600
7400 7200 21000 11250 27000 7400
22300 38900 55400 34150 130000 19700
29000 25200 26500 29300 22400 26600
2300 5600 4100 3500 17000 1800
31300 30800 30600 32800 39400 28400
670
3600
2900
13000
20500
7600
0
7600
Q. Smejkal et al. / Chemical Engineering and Processing 44 (2005) 421–428 Table 5 Selectivity and ethane consumption of variants A–F, EDO process, TFeed = 515 K, p = 16 bar, 50 kt/year production of acetic acid, 100% conversion of oxygen Variant
Selectivity of acetic acid (%)
Consumption of ethane (mol)/mol acetic acid
A B C D E F
76.3 77.7 81.0 79.6 80.7 76.1
1.31 1.29 1.23 1.25 1.24 1.31
the increase of the selectivity to acetic acid results to lower consumption of ethane per formed unit of ethylene (see Table 5). On the contrary, the separation effort is increasing with increasing amount of water. Taking into account the antagonistic requirements – high water content to improve the selectivity to acetic acid on one hand, and low water concentration to minimise the separation effort on the other hand – variant F (73 mole % ethane, 12 mole % oxygen, 15% CO2 ) seems to be a best variant of EDO (see Table 4).
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Table 6 Design assumptions for acetic acid production calculated for year 2002 Parameter
Value/type
Capacity Procedure
10–200 kt/year acetic acid (a) EDO (b) Methanol carbonylation
Process location Process control Type of process Soil condition Pressure vessel design code
Europe Digital Grass roots/clear field Soft clay DIN
acid process was used located in Europe and designed according to DIN. The process flowsheet for ethane oxidation and methanol carbonylation was adopted for the cost analysis. The plant unit results discussed above show the advantage of variant F of EDO, where CO2 is used as a dilution agent. The results from the flowsheet simulation were used as an input file for cost prediction routine ICARUS by ASPEN PLUSTM . 4.1. Investment cost
3.1.3. Comparison of the EDO and BP process The comparison of variant F of selective oxidation of ethane with the commercial BP methanol carbonylation processes is documented in Table 4 for an acetic acid production plant of 50 kt/year. Among the ethane oxidation processes, the set-up with the highest economic potential is represented by variant F with a feed composition 73 mole % ethane, 12 mole % oxygen and 15 mole % CO2 a cooling temperature and feed temperature of 515 K. The energy consumption for compression and product cooling is the lowest in the case of methanol carbonylation. On the contrary, the effort for the pre-heating and product purification (rectification) is lower in case of EDO variant F; the heat produced by the chemical reaction is higher in the EDO process compared to methanol carbonylation. Moreover, the high-density steam can be produced only in case of the EDO process (see Table 4).
4. Cost analysis The objective of this part of the study is to evaluate the economics of ethane direct oxidation compared to industry relevant methanol carbonylation process by BP [3,14]. The total investment and production costs of acetic acid consist of: (1) fixed (investment) costs, including apparatuses, project, buildings, labor, analysis and control [17,18] (2) production costs, i.e. variable costs, including the raw material ethane and oxygen (EDO process) and methanol and CO (BP process) price and the utilities – steam, gases, catalyst, cooling water, fuel, energy, etc. The basic assumptions used in process cost estimation are given in Table 6. A low to high-scale 10–200 kt/year acetic
The block flowseet for ethane direct oxidation and methanol carbonylation is identical with process schemes shown in Figs. 1 and 2. The total investment costs are based on total equipment costs. Therefore, the first step of cost estimation comprises the calculation of the purchase costs of apparatuses included in the process simulation. The results are summarised in Table 7. The price of the reactors was roughly estimated. Here, the material costs could play an important role. Due to the highly corrosive catalyst mixture used in methanol carbonylation process, the reactor and also pre-heater, cooler and flash have to be made from expensive Hastelloy alloy (see Chapter 2.2). These specific aspects have been included in the calculation, the catalyst reactivation has not. The equipment cost summarised in Table 7 shows an advantage of EDO process compare to methanol carbonylation set-up by BP for all production capacities of acetic acid. The main reduction on the apparatus cost can be seen on rectification and pre-heating, the EDO process is more simple. Also the reactor price is lower for EDO process mainly due to the special and expensive Hastelloy alloy used in methanol carbonylation set-up. The costs for a special rectification column for catalyst separation (see also Chapter 2.2.) overcomes all other costs and shows a benefit of simple and by material used non-limited EDO process. On the other hand, the price of compressor for EDO process increases dramatically with production amount and for 200 kt/year is the compressor a major contributor to the total equipment costs in EDO process (almost 70% of equipment costs). The investment costs have been calculated by two different methods. In ICARUS calculation, the investment costs depend on a number of factors like location, type of controlling, process and coil type (the so-called description method was
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Table 7 Investment costs of major equipment and total investment costs (103 US$) for acetic acid production equipment by EDO and BP methanol carbonylation for 10, 50 and 200 kt/year process Equipment
Description
P-6 H-2 R-1 H-4 S-5 C-6 C-8 S-9
Compressor Preheater Reactor Cooler Flash Rectification Catalyst separator Absorber
10 kt/year EDO
Equipment cost Investment cost [19] Investment cost [20]
50 kt/year BP
EDO
200 kt/year BP
EDO
BP
1150 45 114 70 50 540 – 46
1250 70 250 – 290 470 1620 50
2170 60 130 130 60 1150 – 46
1410 110 383 – 560 1050 4760 60
5540 80 140 320 85 1900 – 50
2330 120 450 – 1400 1250 8930 60
2020 7020 6000
4000 14000 9112
3740 13020 9170
8340 29030 17300
8100 28200 16400
14500 50630 24500
applied). In the “module method” by Guthrie [19] the investment costs have been obtained multiplying apparatuses prices by factor 3482. The difference of both calculation methods comprises the specification of non-direct costs, mainly the price of special materials used in methanol carbonylation process. The investment costs calculated by ICARUS [20] are from 12 to 50% lower compared to the costs resulted from the module method. The ICARUS method processes more comprehensive information and will be, therefore, used in the subsequent calculation of production costs. 4.2. Production costs The input file for ICARUS simulation is illustrated in Table 8. Five main parameter were compared by production cost calculation: (i) Project cost (ii) Operating cost
(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 calculated per year and specified production of acetic acid. The project cost (i) means the investment to the plant unit, price of ground, project, buildings, etc. Operating cost (ii) represents the lab, staff, analysis and control spendings. Raw material (iii) is here defined as a price of a feed components used (ethane and oxygen, methanol and CO). Utility cost (iv) means the investment to the steam, cooling water, electricity and fuel. The assumptions used in calculation are related to the investment analysis and description of the parameters [21] and are summarised in Table 9. Under assumptions from Table 8 the production cost calculations were performed. The results for 10–200 kt/year plant are listed in Table 9. The prices of raw material were adopted from literature [22–24] and added into the calcula-
Table 8 Input data for cost estimation Parameter
Data
Parameter description
Number of hours Desired rate of return (%) Number of periods for analysis Tax rate (%) Project capital escalation (%) Product escalation (%) Raw material escalation (%) Operating and maintenance lab escalation (%) Utilities escalation (%) Working capital percentage (%)
8000 20 20 40 5 5 3.5 3 3 5
Term used for investment analysis Rate of return on investment in percent per year Number of periods included in the cashflow and other project totals and calculations Percentage of earnings before taxes per year that must be paid to the government Rate at which project capital expenses may increase expressed in percent per year Rate at which the sales revenue escalates per year Rate at which the ethane price escalates per year Rate at which the operating and maintenance costs escalate per year Rates at which the utilities costs escalate per year Amount required to operate the facility until the revenue from product sales is sufficient to cover costs expressed as a percentage of total capital expense per year Supplies and laboratory charges expressed as a percentage of operating labor costs per year Charges during production for services, facilities, payroll overhead etc. expressed as a percentage of operating labor and maintenance costs per year General and administrative costs incurred during production such as administrative salaries/expenses, product distribution, research and development etc. expressed as a percentage of operating costs per year
Operating charges (%)
25
Plant overhead (%)
50
G and A expenses (%)
8
Q. Smejkal et al. / Chemical Engineering and Processing 44 (2005) 421–428 Table 9 The cost estimation (103 US$) of EDO set-up (F) and commercial BP process for capacity of the unit 10–200 kt/year acetic acid, the conditions see Tables 6–9 F
BP
10 kt/year Investment cost Operating cost Raw materials cost Utilities cost Acetic acid price ($/kg)
6000 3500 1080 470 0.51
8960 3530 1200 210 0.49
50 kt/year Investment cost Operating cost Raw materials cost Utilities cost Acetic acid price ($/kg)
9170 10400 5300 2500 0.36
17300 10400 6000 1200 0.35
200 kt/year Investment cost Operating cost Raw materials cost Utilities cost Acetic acid price ($/kg)
16400 35700 21000 9830 0.33
23500 31700 24000 2650 0.29
tion: ethane price: 72 US$/m.t., methanol: 166 US$/m.t. The price of CO and oxygen is almost similar, approximately 50 US$/m.t. The results in Table 9 show the relation between investment, utilities, raw material and operating costs. In the case of 10–200 kt/year unit, the advantage of process set-up F (direct oxidation of ethane to acetic acid) is shown. The reaction and separation unit is more simple compare to the methanol carbonylation and therefore the operating cost is lower. Moreover, the ethane oxidation unit is made from normal stainless steel and the investment cost is thus lower compare to the same unit capacity based on methanol carbonylation made from expensive Hastelloy (for capacity 10 kt/year). Due to the cheaper feedstock, the raw material cost is lower than
the same of methanol carbonylation set-up (all compared variants). The variant F has a disadvantage, the non-reacted ethane and CO2 must be recycled and reused as a feedgas. Thus, the utility cost is for the variant F in all production capacity higher than for methanol carbonylation. On contrary, the utility cost plays not the most important role in total costs. Depending on the unit capacity, the utility cost represents from 10% rel. for capacity 10 kt/year up to 15% rel. for capacity 200 kt acetic acid per year of total annual costs. The investment and production cost calculation results in an estimation of the acetic acid price in all variants and productions discussed above. The price of acetic acid consists of utility, raw material and operating costs and decreases with the capacity of the unit. In case of both EDO and methanol carbonylation process (50 kt/year) represents nearly 0.36 $/kg. This price is two times lower then the actual market price 0.7 $/kg [6]. For the highest acetic acid production 200 kt/year mentioned in this study, the price of acetic acid is decreasing to 0.33 $/kg, the price for acetic acid calculated for methanol carbonylation is even lower, 0.3 $/kg. The price of acetic acid must be additionally related to the investment costs, which are in all cases higher for traditional acetic acid technology. The improved selectivity of acetic acid in EDO process (due to the catalyst optimisation) can bring further advantages of process economy to EDO process; i.e. the selectivity increase of 90% calculated for EDO production of 50 kt/year results to the acetic acid price of 0.3 $/kg and EDO process becomes even more interesting for industrial application. Moreover, the acetic acid prices seems to stay at about 0.70 $/kg in next quarters. The acetic acid purchasing prices for the period from 2000 to the end 2002 with the forecast for year 2003 are listed in Fig. 3. The purchasing price of acetic acid decreases in last three quarters from its maximum of 0.79 $/kg in third quarter 2001 to actual 0.68–0.70 $/kg. For year 2003, a slight price rise could be expected to 0.72–0.73 $/kg of acetic acid.
0,9 outlook
0,8 0,7 0,6 0,5 0,4 0,3 0,2 0,1
0 II/ 0 0 20 III 00 /2 0 IV 0 0 /2 00 I/2 0 0 II/ 0 1 20 III 01 /2 0 IV 0 1 /2 00 I/2 1 00 II/ 2 20 III 02 /2 0 IV 0 2 /2 00 I/2 2 0 II/ 0 3 20 III 03 /2 0 IV 0 3 /2 00 3
0 I/2
AA Price ($/kg)
427
Quarter/year Fig. 3. Purchasing prices of acetic acid [6,23].
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5. Conclusions
References
The most attractive variant of the EDO reaction and separation set-up from energy point of view seems to be a variant F, where the water-free reaction mixture consists of ethane, oxygen and CO2 . Considering the ethane consumption per mole acetic acid as the major objective, the variant C with water in the feed is probably the best one. In general, the addition of high amount of the water to the feed is not recommended, since the separation effort increases significantly. The comparison of variant F of ethane oxidation to acetic acid with the commercial BP carbonylation of methanol shows that the new technology based on selective oxidation of ethane can compete with the existing processes for acetic acid production. From the cost analysis the following aspects can be concluded: acetic acid price is lower in methanol carbonylation compared to EDO in all cases, but the investment costs in state-of-the-art methanol carbonylation process at the same acetic acid quality are much higher due to the usage of a special resistant materials for construction of the plant. If the selectivity to acetic acid can be increased in EDO from 80 to 90%, the ethane direct oxidation becomes to be even more attractive for industrial application.
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Acknowledgment The study was supported by the German Federal Ministry for Education and Research (Project FKZ03C3013) and the Department of Science of the State of Berlin.