Exergy analysis with a flowsheeting simulator—II. Application; synthesis gas production from natural gas

Pergamon

Chemical En#ineerin~4 Science, Vol. 51, No. 20, pp. 4701-4715, 1996 Copyright ~ 1996 Elsevier Science Ltd Printed in Great Britain. All rights reserved P I I : S0009-2509(96)00221-7 ooo9 2509/96 $15.00 + 0.00

EXERGY ANALYSIS W I T H A F L O W S H E E T I N G S I M U L A T O R - - I I . A P P L I C A T I O N ; SYNTHESIS GAS P R O D U C T I O N F R O M N A T U R A L GAS A. P. H I N D E R I N K , * F. P. J. M. K E R K H O F , *t A. B. K. LIE,* J. D E S W A A N ARONS* and H. J. V A N D E R K O O I * *Stork Comprimo B.V., P.O. Box 58026, 1040 HA Amsterdam, The Netherlands *Laboratory of Applied Thermodynamics and Phase Equilibria, Delft University of Technology, The Netherlands (First received 17 November 1994; revised manuscript received and accepted l0 January 1996)

Abstract--Several processes, producing synthesis gas from natural gas, have been analyzed by the exergy method, showing exergy analysis as a valuable diagnostic tool. In addition, a generally applicable and systematic way of performing exergy analyses and dealing with their results is illustrated. Exergy calculations have been carried out by user-defined subroutines, which are integrated with a flowsheeting simulator. The method of calculating exergies has been described in part 1 of this article. First, in order to systematically perform exergy analyses, the overall exergy loss for each process is determined. Absolute exergy losses based on final product yield, which is chosen to be methanol, are used for process comparison and diagnosis rather than exergetic efficiencies. Compared to the conventional steam reforming process, giving an overall exergy loss of approximately 8.5 GJ/t methanol, the exergy loss can be reduced to about 4.9 GJ/t methanol by application of the convective reforming option in combination with partial oxidation. Secondly, by considering progressively smaller subsystems within the overall process, locations of major exergy loss are revealed and their potential for improvement can be indicated. Finally, the minus value of the standard Gibbs energy of the overall reaction of each process, denoted as available reaction exergy, is compared to the exergy loss associated with this overall reaction. This comparison demonstrates that available reaction exergies should always be minimized to reduce exergy losses associated with chemical reactions. They cannot however, be eliminated completely when reactions are only thermally coupled. Further improvement can be attained by direct coupling of chemical reactions such that the overall Gibbs energy of reaction is still reasonably negative. This latter conclusion will hopefully result in a reconsideration of the chemical paths along which important chemicals are produced nowadays. Copyright .(7 1996 Elsevier Science Ltd

1. INTRODUCTION In order to illustrate the applicability of the exergy concept and a way of presenting results of exergy analyses of chemical processes, computer-aided exergy analyses have been carried out for several process routes for the production of synthesis gas. Synthesis gas processes are selected, since these are large energy consumers. Since methanol synthesis is one of the largest synthesis gas based processes, it is worthwhile to approach the extent of exergy-utilization of synthesis gas production with the specific case of methanol synthesis. The aim of these exergy analyses is matching the revealed differences in exergy losses with differences in process configuration and chemistry.

Synthesis gas is generally produced from natural gas, mainly consisting of methane, by either catalytic steam reforming or partial oxidation (see Appendix

A). Principal reactions during steam reforming: CH4+H20

~-~CO+3H2,

+ 144kJ/mol

{1) CH4+CO2

~2CO+2H2,

A,G ° =

+173kJ/mol

(2) C O + H 2 0 ~_~ CO2 + H2,

ArG o = - 29 k J / m o l (3)

Principal partial oxidation reaction: CH4 + ½02 ~ C O + 2H2,

2. S Y N T H E S I S GAS P R O D U C T I O N ; T H E O R Y

Mixtures of hydrogen and carbon oxides are usually referred to as synthesis gas. Today, methanol is almost exclusively produced by the catalytic conversion of C O and COz with HE, i.e. from synthesis gas.

A,G ° =

ArG o = - 85 kJ/mol.

(4) The ideal synthesis gas composition for methanol synthesis is best described by the so-called stoichiometric number (SN): SN

*Corresponding author.

nn2 - nco2 ~CO + ~C02"

4701

(1)

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A. P. HINDERINK et al.

Steam reforming of natural gas yields a synthesis gas with a SN of approximately 3, whereas partial oxidation of natural gas produces synthesis gas with a SN below 2. The final composition varies to some extent depending on hydrocarbon feedstock characteristics and operating conditions, but both processes require ratio adjustment, since the ideal synthesis gas for methanol synthesis must have a SN of 2. More detailed information on the theory of synthesis gas production is given for example by Supp (1990). 3. PROCESS DESCRIPTIONS

Several (commercial) processes exist for the production of methanol, mainly differing in the way the intermediate synthesis gas is produced. These processes are described in literature and they all use either or both the steam reforming and the partial oxidation process. In this article, all processes are given a code by which they are indicated in tables and graphs. Processes selected for exergy analysis are the following: • Conventional steam reforming; designated as process A1.

• Conventional steam reforming + pre-reformer; designated as process A2. • Combined reforming with a natural gas bypass; designated as process B. • Convective reforming parallel to non-catalytic partial oxidation; designated as process C. In fact several more processes have been analyzed. They will however not be discussed for reasons of limitations in the length of this article. These processes are: • Steam reforming with CO2 addition to the reformer feed. • Steam reforming with CO2 addition to the synthesis gas. • Combined reforming without a natural gas bypass. • Convective reforming parallel/in series to catalytic partial oxidation. Conventional steam reforming will be considered as the reference process during the exergy analyses. This process produces excess hydrogen, while the more sophisticated process configurations are believed to be able to directly produce stoichiometric synthesis gas. Operating conditions of each of the processes A 1, A2, B and C are presented in Table 1. These data are taken from literature or are results from simulation calculations. Process A 1: conventional steam reforming

Figure l(a) shows the major process steps of the conventional steam reforming process in a block diagram. Natural gas, which is actually a high caloric gas, is depressurized from 6000 kPa to the operating pressure of 2000 kPa and then mixed with steam giving a steam-to-natural gas ratio of 3.5 on a molar basis.

The mixture is preheated to 500°C before entering the steam reformer. In this reformer, the mixture catalytically reacts to hydrogen and carbon oxides at temperatures up to 900°C. Hot combustion gases leaving the reformer furnace are cooled to about 150°C before disposal to the environment. Hot product gases are cooled to 50°C, whereafter condensate is separated. The resulting dry synthesis gas is then compressed and heated to methanol synthesis conditions, which are set to 300°C and 6000 kPa. Excess hydrogen (and unconverted methane) is separated from the methanol synthesis loop and directed to the reformer furnace where it is combusted together with fresh fuel. Process A2: conventional steam reforming including a pre-reformer

The conventional steam reforming process involving a pre-reformer is only a minor change of the process described above. In the pre-reformer, which operates adiabatically, all higher hydrocarbons and some of the methane are converted to an equilibrium mixture at about 400°C by using medium temperature level process heat (300-600°C). After the pre-reformer, the product gases are heated again to about 500°C, whereafter they enter the main reformer. All other process steps are identical to those of the conventional steam reforming process. Further information on conventional steam reforming and the pre-reforming option is given by Goff and Wang (1987), Foreman (1990), Farina and Supp (1992), Giacobbe et al. (1992), Johansen et al. (1992), Schneider and LeBlanc (1992) and Baade et al. (1993). Process B: combined reforming

Figure l(b) shows the process steps of the combined reforming process in a block diagram. This process option combines the catalytic steam reforming process, which produces excess hydrogen, in series with a catalytic partial oxidation process (oxygen reforming), which produces excess carbon oxides. By properly selecting operating conditions, i.e. pressure and temperature, both processes can be combined, such that the desired synthesis gas composition is directly produced. Only half of the natural gas feedstock is directed to the primary steam reformer (also called primary reformer), now operating typically at a temperature of about 780°C. The steam-to-natural gas ratio of the feed directed to this primary reformer is kept equal to the ratio used in the previous processes. The less severe conditions of this reformer allow for operation at a pressure up to 4000 kPa, instead of a maximum pressure of about 2000 kPa in case of conventional steam reforming. As a result, synthesis gas compression downstream is lowered. After primary reforming, product gases are mixed with fresh feedstock and pure oxygen, generated by an air separation unit, and subsequently directed to a catalytic partial oxidation unit (secondary reformer),

4703

Exergy analysis with a flowsheeting simulator--II. FUEL

AIR

HEAT --

FEEDSTOCK

WATER

--

FEED PREPARATION

FURNACE

-

- )

REFORMER

HEAT

r LU (3

RECOVERY AND TRANSFER

HEAT

¢k

--

--'~

USEFUL

HEAT

--

---~

HEAT TO CW CONDENSATE

"I"

EXHAUST TO STACK HEAT

SYNGAS

PREPARATION

COMPR.

POWER

~E_

SN = 2.8

t i I

--

--

METHANOL SYNTHESIS LOOP

I_

(a)

7

I J

I

METHANOL

Fig. 1. (a) Block diagram of the conventional catalytic steam reforming process (A1).

where further conversion is accomplished at temperatures near 1000°C. No additional steam supply is needed if this type of reformer is used as a secondary reformer. Besides the literature references already given for the previous process options, additional information on combined reforming is presented by Supp (1985), Vannby and Winter-Madsen (1992) and Christensen and Primdahl (1994).

Process C: convective reformin9 parallel to partial oxidation Figure l(c) illustrates the general block diagram of the convective reforming process parallel to noncatalytic partial oxidation. In conventional steam reforming, considerable amounts of sensible heat are produced and only approximately 60% of the heat of combustion is used for driving the endothermic reforming reaction. The remainder is used

for preheating feeds and generating high pressure steam. In contrast, the convective reforming process option in combination with partial oxidation uses high temperature level process heat ( > 600°C) for converting a part of the reformer feed in a heat exchanger reactor, in which the tube-side is filled with catalyst (for convenience, this type of reformer is also designated as primary reformer in Table 1). The tube-side effluent is mixed with the shell-inlet stream, originating from the partial oxidation unit and having a temperature of about 1300°C. As the resulting mixture flows through the shell it serves as a heating medium for the endothermic reforming reaction occurring inside the tubes. Operating pressures can be twice as high as for conventional steam reforming. According to Schneider and LeBlanc (1992), this process option is believed to be able to directly produce stoichiometric

4704

A.P. HINDERINKet al. FUEL FEEDSTOCK WATER

AIR

HEAT -- --

t

AIR

AIR

FEED PREPARATION

l

SEPARATION

• ~-~

COMPR. POWER

TM

t OXYGEN

L CATALYTIC PARTIAL O X I D A T I O N

1 HEAT RECOVERY AND TRANSFER

t

--'1~ USEFUL HEAT --tk HEAT TO CW

r

CONDENSATE

EXHAUST TO STACK HEAT/l

SYNGAS PREPARATION

COMPR. "1 POWER

(b)

1 SYNGAS SN = 2.0

Fig. 1. (b) Block diagram of the combined reforming process (B).

synthesis gas, without the need for a reformer furnace, if operating conditions are selected properly. However, with a high caloric gas taken as feedstock, stoichiometric synthesis gas can only be obtained by installing a pre-reformer prior to the convective reformer. Also, the process is found to be energy deficient, making external heat supply necessary. Besides the literature references already given for the previous process options, additional information on convective reforming is presented by English et al. (1990). 4. SIMULATION AND ANALYSISPROCEDURE The method of systematically performing exergy analyses is partly described by, for example, Gaggioli

(1980), Maloney and Burton (1980), Moran (1982) and Kotas (1985), and consists of the following steps: • Define a suitable and equal system boundary for all processes to be analyzed and define (as far as possible) equal process transformations using equal starting materials--preferably natural re-

sources. • Model all processes (and equal unit-operations) in a similar manner and carefully make assumptions with respect to this modelling, since they directly affect the exergy analysis results. • Choose a convenient heat recovery and heat integration strategy applicable to all processes being analyzed.

Exergy analysis with a flowsheeting simulator--II.

AIR

FEEDSTOCK WATER

LI

(IMPORT) HEAT-- --b

4705

FEED PREPARATION

NATURALGASBYPASS

kk I

I AIR -¢ ~ SEPARATION

COMPR. POWER

OXYGEN

PRE-REFORMER

PARTIAL OXIDATION

TUBE-SIDE A

CONVECTIVE REFORMER

I

HEAT

~.

SHELL-SIDE

........

i

......

:

HEAT RECOVERY AND TRANSFER

USEFUL HEAT HEAT TO CW CONDENSATE

,k SYNGAS PREPARATION

COMPR, POWER (c)

SYNGAS SN = 2.0 Fig. 1. (c) Block diagram of the convective reforming process parallel to partial oxidation (C).

• Obtain consistent energy and mass balances (e.g. by flowsheeting simulation). • Calculate absolute exergy flow rates of all material and heat streams. • Define useless and useful streams leaving the system boundary. • Determine the overall exergy loss by accounting and define a basis on which results are presented. • Determine the exergy loss of important subsystems by accounting and determine the nature of the revealed losses. • Prioritize areas for improvement.

4.1. Method Simulation of the various synthesis gas processes has been carried out by using a flowsheeting simulator (Aspen Plus licensed by Aspen Technology, Inc.) and only covered those process items falling within system boundary definition I, as indicated in Fig. 2. A short summary of the simulation and modelling method is given in Appendix B. With user-defined subroutines (ExerCom licensed by Stork Comprimo B.V.), described in part 1 of this article, the exergy flow rates of all material streams were calculated by the flowsheeting simulator simultaneously with energy and massbalance calculations. For these exergy calculations,

4706

A. P.

HINDERINK et al.

Table 1. Characteristics of the synthesis gas processes; simulation results

Combined ref. B

Convect. ref. non-cat. POX Parallel C

Unit

Steam ref. A1

Pre-ref. Steam ref. A2

Feeds Nat. Gas [NG]* Steam

Nm3/t MeOH t/t MeOH

884 1.9

863 1.9

801 1.0

669 1.1

Compression power Total

GJ/t MeOH

0.61

0.61

0.86

0.83

Pre-reformer NG feed In/out temp. Pressure

N m3/t MeOH °C kPa

N.A.

668 500/430 2000

N.A.

261 525/465 4000

Primary reformer Type NG feed In/out temp. Pressure Absorbed duty

Nm3/t MeOH °C kPa GJ/t MeOH

SR 688 500/900 2000 9.1

500/900 2000 8.7

Reformer furnace Fuel feed H2-purge feed Out temp. Fired duty [LHV]

Nm~/t MeOH Nm3/t MeOH '~C GJ/t MeOH

196 610 920 15.1

175 611 920 14.4

Autothermal reformer Type Fresh NG feed Oe supply In/out temp. Pressure

N m3/t MeOH N m3/t MeOH °C kPa

Item

Syngas SN* % CH4 Condensate

SR 342 500/780 3500 2.8

Convective tube: 465/940 4000 3.5

117

2.82 1.59

2.82 1.59

1.2

1.2

Purge of excess hydrogen

Purge of excess hydrogen

mol% t/t MeOH

Remarks

SR

800 4.4

N.A

Cat. POX 342 281 670/980 3500

POX 408 288 500/1325 4000

2.02 1.79 0.7

2.00 1.21 0.9

50% NG bypass Additional heat primary reformer supply necessary

*Reformer feedstock and furnace fuel available at 60 bar; *including excess hydrogen for processes A1 and A2.

SYSTEM

BOUNDARY

DEFINITION

SYSTEM BOUNDARY

IMPORT EXERGY

- -~-

NAT. GAS EQUIVALENTS

~

-k

-i

~

. . . . . .

--'~

HEAT INTEGRATION

A'R sEP.

,-

,

- - ~ :

/

so..

FEEDS

%,!!]

__~

~

I

[HEAT]

i NAT. GAS EQUIVALENTS

DEFINITION

. . . . . . .

~

50%

A,. -- ]

II

A I

.I ~

I J '1.

I

WASTE STREAMS

EXPORT EXERGY

S Y.T.ESIS

~AS PROCESS -- - H2 PURGE

,.E..,

i

I'

SYNTHESIS GAS [H~ -- CO,}

I

-

-

=

2

ICO + C02]

INPUT

P

PROCESS

b

OUTPUT

t

Fig. 2. System boundary of a general synthesis gas process used for performing exergy analyses.

4707

Exergy analysis with a flowsheeting simulator--II. the reference environment model proposed by Szargut et al. (1988) has been adopted. All simulations satisfy the following process transformation: Process transformation: The production of a mixture of hydrogen and carbon oxides with a stoichiometric number of 2.0 at methanol synthesis conditions (i.e. 300°C and 6000 kPa)from the commonly available starting materials natural gas, water and air at well specified conditions. 4.2. Assumptions The processes require different amounts of electricity for driving compressors (power) and may also need additional exergy supply in the form of low temperature level heat (import exergy). These external process inputs have to be generated from natural gas and must therefore be given a generation efficiency for obtaining fair exergy analysis results. External assumptions, with respect to the efficiency by which power and import exergy is provided to the system, are therefore essential, since they are not covered by system boundary I. A new system boundary, denoted as system boundary definition II and which is also indicated in Fig. 2, incorporates these external assumptions which are the following: • Power (electricity) for driving compressors is generated with an exergetic efficiency (qex,e) of 50%. • Import exergy (heat) is generated with an exergetic efficiency (r/ex,e) of 50%. These efficiency values are accounted for by Kerkhof and van Steenderen (1993). Further, unconverted methane present in the synthesis gas, is not excluded from the useful exergy output, since it is assumed to be reusable after separation from the methanol synthesis loop. 4.3. Heat recovery method During simulation and analysis, close attention has been paid to heat recovery. Since no detailed heat recovery and heat integration data are available in literature for most of the processes being analyzed (and the methanol synthesis itself is not considered), a general method has been applied. Instead of direct heat transfer between actual process streams, the maximum transferrable amount of exergy on cooling a hot stream is estimated by means of exergetic efficiencies of heat transfer (r) as introduced by Kotas (1985) and Szargut et al. (1988): AEXheated medium rex

AEXheating medium .

(2)

Three temperature levels above 150°C are defined at which heat transfer might occur. For each of these levels, typical values of r are determined, which are found to be 0.82, 0.95 and 0.78 for heat transfer at low,

medium and high temperature levels, respectively. Heat streams below 150°C are not submitted to heat recovery (z = 0). The z value of heat transfer at high temperature level ( > 600°C) is poor, since practically no heat can be transferred at such a high temperature level without large temperature gradients. For heat exchange at the other temperature levels, a temperature difference of 50°C has been assumed for determining the r value (Kotas, 1985). By using these efficiency factors, the maximum transferrable amount of exergy on cooling a hot process stream can be determined easily, which in turn can directly be used for heating purposes. 4.4. Determination of exergy losses For the determination of the overall exergy loss, it is necessary to make a subdivision into internal and external losses. External exergy losses are associated with useless heat and material streams leaving the system boundary and rejected to the environment, while internal exergy losses are caused by process irreversibilities. In this study, losses accompanied by generating power and import exergy are incorporated into the internal exergy loss. Useful streams leaving the system boundary are assumed to consist only of the synthesis gas product and the exergy value of excess heat (designated as export exergy). No penalty is given for the production of export exergy; however, it actually is an undesired product. An exergy analysis of a real process compares the exergy entering this process to the exergy leaving this process and shows that: ~Ex + > ~Ex-

(3)

which is in agreement with the second law of thermodynamics. The overall exergy loss of a synthesis gas process according to definition II of the system boundary and considering only useful exergy output is calculated by the following expression:

yp+

~Ex~

~e×,P

~ex,Q

Exloss = ~Ex~*~ + -

EEx~.

(4)

Results of the exergy analyses will be mainly presented and discussed on the basis of the exergy loss per unit product yield, i.e. methanol yield. This methanol yield is obtained by assuming 100% conversion of the carbon oxides present in the synthesis gas to methanol. The advantage of presenting exergy analysis results in absolute values is that the importance and cost of the determined exergy losses can easily be judged, since absolute exergy losses can be readily converted into economic terms, such as primary fuel consumption in the form of e.g. natural gas equivalents (1 N m 3 natural gas ~ 3.9 MJex), a term process engineers are familiar with. In order to reveal locations of "major" exergy loss, the overall process has been divided into several segments containing specific process steps. For an equal

4708

A.P. HINDERINKet al.

basis, it is important that equal process steps (e.g. mixing of reactor feeds) in different processes always appear in the same process segment. The following process segments have been distinguished for performing detailed exergy analyses: 1. Feed~product preparation (excluding losses associated with heat transfer) • depressurization of natural gas feedstock/fuel • pressurization of water • mixing of reactor feeds (except for autothermal reformer feeds) • feed preheating and generation of process steam • compression and heating of (dry) synthesis gas to methanol synthesis conditions 2. Reaction section: further divided into separate units • pre-reformer • steam reformer tubes (including heat transfer losses associated with external heat supply by a furnace) • reformer furnace • autothermal reformer (including mixing feeds) • convective reformer 3. Heat recovery (including losses associated with heat transfer) • cooling hot combustion gases and hot product gases 4. Air separation • compression of air • (black box) separation of air into a useless nitrogenrich stream and a desired oxygen-rich product • compression of the oxygen-rich product to process pressure by an interstage cooled compressor (including cooling losses) 5. Generation of power for drivin9 compressors 6. Generation of import exergy (heat).

O

O

.= ;A

.e

I

2 8

5. R E S U L T S

AND

oo

°2

~D

,fi ~

The exergy loss of each subprocess falling within system boundary I is determined by accounting. Generally, this method is straightforward, except for heat streams being involved. It must be remarked that the exergy loss of each process segment only involves internal losses, external exergy losses are considered separately and consist of the following: External exergy losses • Stack gases at 150~'C • Process condensate at 50°C • Exergy taken up by cooling water on cooling synthesis gas below 150°C.

O

.e~

g

~Z

~Z *©

5

~m

z zg r,i II

INTERPRETATION

Several authors already have published their experiences with (computer-aided) exergy analysis and present their results in various ways (e.g. Hedman et al., 1980; Yang et al., 1989; Rosen, 1991; Tsatsaronis et al., 1991; Kerkhof and van Steenderen, 1993). In this article, the results of the exergy analyses are presented in a generally applicable way. 5.1. Overall processes Table 2 presents energy and exergy figures of the various synthesis gas process configurations on an

~Z ~m Z~ *

2;

4709

Exergy analysis with a flowsheetingsimulator--II. absolute basis per unit of methanol yield according to the system boundary definition II. Since absolute exergy losses are invariant and directly reflect the amount of work-potential being lost, regardless of the total amount of incoming exergy, these are recommended by us rather than exergetic efficiencies, when reporting on exergy analyses. Also, in our opinion, exergetic efficiencies are a more ambiguous measure, since they can be defined in numerous ways. With the exception of the process inputs water and air, which do not significantly contribute to the overall incoming energy and exergy, all inputs are associated with the application of natural gas, when generation efficiencies for power and import exergy are accounted for. Approximately, 68-82% of the incoming exergy is retained in the synthesis gas product. Apart from export exergy, the remaining part of the incoming exergy is lost either internally or externally. Internal exergy losses as defined in the previous section account for a major part of the overall loss. About 16-21% of the exergy input is internally consumed, in contrast with external losses, which are at the most 2% of the incoming exergy. The results presented in Table 2 indicate a decrease of the overall exergy loss from 8.5 GJ/t methanol for the conventional steam reforming option A1 to 4.9 GJ/t methanol for the convective reforming option C, which equals a reduction of over 40%. It is also illustrative to compare the actual input of exergy to the exergy stored in the final product, which we denote as 'target exergy'. In theory, a quantity of exergy equal to this target exergy, i.e. 22.4 GJ/t for the case of methanol synthesis, must be supplied in the form of natural resources. In practice, exergy supply is larger and ranges from 30 GJ/t to about 36 G J/t, as indicated in Table 2.

The indicated ability of decreasing the overall exergy loss by changing the process configuration illustrates clearly the existing potential for process improvement. However, it should be kept in mind that exergy analysis alone is not sufficient for selecting and judging processes; it remains an additional diagnostic tool besides the other kinds of process analysis. From Table 2, it can be seen that the convective reforming process option C does not produce sufficient process heat for preheating feeds and generating process steam and requires therefore additional heat supply, which is called import exergy. Taking losses due to irreversible heat transfer into account and giving the required import exergy a generation efficiency of 50%, about 2.4 GJ/t methanol of exergy in the form of natural gas equivalents must additionally be supplied to this process. Despite this additional source of exergy loss, the convective reforming process option is found to be superior to the other processes which have been analyzed. In the following subsection, results of a more detailed exergy analysis per process are presented along with a match of the indicated exergy losses with differences in process configuration. 5.2. Process segments Table 3 presents the results of the detailed exergy analyses. Before describing these results for each process separately, first, some general findings are given below. • Two major chemical conversion steps, i.e. combustion and partial oxidation, account for the largest consumptions of exergy. Improving the way of driving the endothermic reforming reaction, that is, reducing the natural gas

Table 3. Subdivision of the overall exergy loss (GJ/t MeOH) per process section Process Item

A1

A2

B

C

Feed/product preparation Reaction section Pre-reformer Steam reformer Reformer furnace Convective reformer Autothermal reformer* Heat recovery/transfer Air separation * External loss Stack loss

0.84 4.8 N.A. 0.60 4.22 N.A. N.A. 1.38 N.A. 0.80 0.50

0.81 4.5 0.07 0.48 3.98 N.A. N.A. 1.35 N.A. 0.77 0.47

0.26 3.2 N.A. 0.19 1.47 N.A. 1.54 0.74 0.49 0.31 0.16

0.056 1.7 0.09 N.A. N.A. 0.09 1.52 0.51 0.50 0.15 N.A.

Power generation Generation import exergy

0.61 N.A.

0.61 N.A.

0.86 N.A.

0.83 1.18

8.45

8.07

5.86

4.93

Overall

*including mixing of the autothermal reformer feeds; ~including loss by Nz rejected to the environment and losses by the interstage cooled 02 compressor.

4710

A. P. H1NDERINKet al.

comsumption per amount of methanol being produced, mainly determines the improvement on the basis of exergy analysis; • Processes in which uncontrolled combustion is reduced (option B) or eliminated (option C) show large improvement with respect to exergy-utilization; • Besides losses due to chemical irreversibilities, heat recovery is a second large exergy consuming process step, mainly due to the low efficiency by which high temperature level heat can be recovered; • The extent of the external exergy loss depends on the role of the reformer furnace, since a major contribution to this kind of exergy loss are stack gases, which are inherent to the application of a furnace. In the remaining part of this section, results of the detailed exergy analyses are matched with the specific process configuration. General ways of reducing exergy losses are given by Sama et al. (1989) and de Swaan Arons and van der Kooi (1993). Process AI: It is obvious that a major part of the exergy loss associated with steam reforming--approximately 50%--is attributable to losses occurring in the furnace section. On the whole, losses of other process segments are to a greater or lesser extent related to the use of external heating by a furnace. A remarkable aspect is that a large part of losses within the feed/product preparation section is attributable to mixing of the hydrogen purge with fuel and air. This source of exergy loss can be avoided by eliminating the production of excess hydrogen (e.g. by supply of external CO2). Process A2: About 12% of the reforming duty is completed within the pre-reformer, resulting in a lowered load of the reformer and consequently a decrease of the fired duty of the furnace (see Table 1). All exergy savings are attributable to this reduction of the fired duty. The additional loss associated with prereforming is small compared to exergy savings elsewhere. Process B: As can be seen from Table 1, the combined reforming process accomplishes a lowering of external heat supply by a reformer furnace. The exergy savings

accompanied by this reduction of the furnace duty outweigh the supplemental exergy loss associated with oxygen reforming and air separation. The main reason why a partial replacement of steam reforming by oxygen reforming leads to reduction of the exergy loss lies in the following: • internal combustion with pure oxygen eliminates heat transfer through reformer pipe-walls and heating of ballast nitrogen, resulting in a lowering of the fuel requirement. • Oxygen reforming applied as secondary reforming does not need additional steam supply; thus steam requirement and consequently the reformer load is lowered. • combustion products of autothermal reforming serve as reactants for the steam reforming reaction also occurring in this reactor and are thus not disposed to the environment. Process C: The convective reforming process operated parallel to partial oxidation requires input of additional heat when stoichiometric synthesis gas must be produced. The advantage of eliminating the reformer furnace is therefore partly overshadowed by the need for generating this additional heat (import exergy). Besides the elimination of a reformer furnace, which also leads to a reduction of losses associated with feed preparation, a further improvement is obtained by the utilization of high temperature level process heat ( > 600°C) for converting a part of the hydrocarbon feedstock in a heat exchanger reactor. Since this way of utilization of process heat can be accomplished with smaller temperature gradients than when producing high pressure steam, process efficiency is improved with respect to heat recovery and heat transfer.

6. AVAILABLE REACTION EXERGY A N D EXERGY LOSS

In this section, the standard Gibbs energy of the overall reaction of the various processes is compared to the exergy loss associated with this overall reaction. The overall reaction is the net result of reactions with a negative A,G° and reactions with a positive A,G °, thus incorporating both combustion steps and

Table 4. Overall reactions Process

- A,G° (GJ/t MeOH)

(a) From natural gas to synthesis gas A1 1NG + 0.8902 ~ 1.86H2 + 0.53CO + 0.56CO2 + 0.17H20 A2 ING + 0.8602 ~ 1.92H2 + 0.54CO + 0.55CO2 + 0.12H20 B ING + 0.08H20 + 0.6502 ~ 2.00H2 + 0.65CO + 0.39CO2 + 0.05CH4 C 1NG + 0.39H20 + 0.4302 -~ 2.33H2 + 0.81CO + 0.24CO2 + 0.04CH4

10.8 10.0 6.1 1.9

(b) From natural gas to methanol A1 1.260NG+ 1.11702 --* IMeOH + 0.373CO2 + 0.549H20 A2 1.224NG+ 1.05302 --* 1MeOH + 0.334CO2 + 0.486H20 B 1.147NG+ 0.74502 ~ 1MeOH + 0.186CO2 + 0.167H20 + 0.03H2 + 0.06CH4 C 0.952NG+ 0.143H20 + 0.41002 ~ 1MeOH + 0.038CH4

11.4 10.5 6.7 2.4

4711

Exergy analysis with a flowsheeting simulator--II.

reforming steps. The first category of reactions can proceed spontaneously; hence they are denoted by us as 'downhill' reactions. The latter category of reactions cannot proceed without a driving force and are denoted by us as 'uphill reactions'. Material streams entering and leaving the system (i.e. the reaction section) are regarded as being demixed to eliminate mixing effects from the discussion below. Also, processes using pure oxygen have not been given a penalty. In Tables 4(a) and 4(b), the overall reactions and their standard Gibbs energy of reaction are presented for conversion of natural gas to synthesis gas and to methanol, respectively, for every synthesis gas process. Standard Gibbs energies of reaction can easily be determined, since they equal the change in the standard chemical exergy during a reaction. Values of standard chemical exergies are taken from Szargut et al. (1988). In the following, the minus of the standard Gibbs energy of the overall reaction is denoted as available reaction exeryy. It plays an important role in understanding exergy losses associated with chemical reactions. In Fig. 3, the available reaction exergies of the synthesis gas processes are graphically presented. As can be seen from this diagram, the various processes differ largely in their available reaction exergy. Since the available reaction exergy is the net result of reactions downhill and reactions uphill, its value is determined by the excess of the downhill reactions (combustion steps). Consequently, differences in available reaction exergy indicate differences in fuel requirement for driving the uphill reactions. The two reference reactions indicated in this figure are discussed at the end of this section. In order to gain more insight into the importance of the available reaction exergy for improving chemical processes with regard to exergy-utilization, this value

[kJ/tool MeOHI

12

320

10 ...................

,~ ~ 8 ~, ~, 6

'; <

4

256

....................... i ...........................

i

......................... ° I,i ....... ..................., k

C

2~- . . . . . . . . . . . . . .

~ .....

! .........................

Exlo.

Ex°cham .....

iiii~ii~

: .

REACTION

::::

: : : : : ::::: ::

:::::.

:"

:

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INPUT

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iiiii!iiiii;iiiiiiiiiii ~

4

is compared to the exergy loss associated with the overall reaction. Part of the chemical exergy entering the system is released, i.e. the available reaction exergy, and can only be converted into physical exergy in our case. This conversion, however, is not complete, but is accompanied by an exergy loss, see Fig. 4. The relation between this loss and the actual Gibbs energy of reaction is given by Denbigh (1956), who assumes only the heat effect of a chemical

1771 !i iii~!:iiii

120

Fig. 3. Available reaction exergy for various process routes to synthesis gas (indicated values are in GJe,/t MeOH).

- - - - EXpnys

ii ii!iiiiiiiiiiiiiiiiii ii iiiiii!iiiiiiiiiiiiii

!1,2

s

EXphv=

ii !ii~iiiiii~iiii:iiiii

384

: ........

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OUTPUT

Fig. 4. Changes in chemical and physical exergy during an overall chemical reaction.

A. P. HINDERINK et al.

4712

reaction, corrected by the Carnot-factor, is a useful form of exergy. The remaining part of the available reaction exergy is considered as unavoidable exergy loss, which can be expressed as follows: EXl°ss'unav°idable

=

--

To r~" ArG(Tr ~.

tical losses can be explained by the presence of inefficient heat transfer steps within the reaction section. However, it should also be remarked that the Denbigh losses are estimates, assuming amongst others a mean overall reaction temperature of about 900°C. The conventional steam reforming process lies at the right of the graph. It has a large available reaction exergy, which is partly converted into physical exergy, the remaining part equals the exergy loss and can primarily be explained by the Denbigh relation. It can be seen that more sophisticated synthesis gas processes have available reaction exergies lower than the conventional steam reforming process. Consequently, these processes produce less physical exergy and have lower Denbigh losses. The graph also indicates that these processes are improved with respect to inefficient heat transfer steps within the reaction section, since differences between practical exergy losses and Denbigh losses are reduced. An important aspect of this graph is the point where the practical exergy loss and the Denbigh loss equal the available reaction exergy (point of interception). At this point, no physical exergy is produced, i.e. no heat production at temperature levels different from To. Because the reactions uphill and reactions downhill are only coupled by their heat of reaction, i.e. thermal coupling, exergy losses are minimized at the point of intersection. Since this point does not coincide with

(5)

For estimating the actual Gibbs energy of a reaction, the following approximation is made, thereby assuming ideal gas behaviour and neglecting the pressure dependence of &H: ArG(Tr)

ArNO-

~

T,

/e,\~,\

A,S ° -

Rlnl~o)

)

(6)

where ArS° can be calculated from the available reaction exergy and the standard heat of reaction. Figure 5 presents the unavoidable exergy losses together with the practical exergy losses of the reaction section of each synthesis gas process obtained by simulation as a function of the available reaction exergy. The practical exergy losses are obtained by adding up all exergy losses within the reaction section, as presented in Table 3, corrected for mixing effects. These mixing effects during reaction range from 0.25 to 0.40 GJ/t methanol. As can be seen from Fig. 5, the practical exergy losses approach the unavoidable exergy losses, predicted by the Denbigh relation (5). Differences between the Denbigh losses and the prac-

Exergy loss reaction section [GJ/t MeOH] 12 J

I

./I fJ

10

. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .

7 j . . . . . . . . . . .

f f f I I I

i"

J

I

~'t-

f

f

6 --

jf y . . . . . . . . . . . . . . . . . . . . . . . . . . . .

. . . . . . . . . . . . . . . . . . . . . . . . . . . . / I" ~f

- -

L . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . /" / 7

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f

'

I 2

'

I 4

'

I 6

'

I 8

'

J

'

10

Available reaction exergy [Gdlt M e O H ]

Fig. 5. Available reaction exergy vs exergy loss (demixed material streams). Only processes indicated by a code are discussed in the text.

12

4713

Exergy analysis with a flowsheeting simulator II. the origin of the graph, it indicates that there is a certain minimum unavoidable exergy loss associated with the overall reaction to synthesis gas. While not changing starting materials to reduce entropy production, further improvement is only possible by direct coupling of chemical reactions by their Gibbs energy of reaction. Since synthesis gas production by direct coupling of reactions is not yet practically accomplished, the convective reforming option in combination with partial oxidation (process C) presents a 'state of the art' process considering recent technology and knowledge. It combines elimination of inefficient heat transfer steps within the reaction section with efficient utilization of process heat for driving uphill reactions. Further improvement of this synthesis gas process is therefore only possible outside the reaction section, if losses associated with thermal coupling are taken for granted. In Fig. 3, two reference reactions are given. Reference I is the direct partial oxidation of methane to methanol. Depending on the temperature level at which the heat of this exothermic reaction is released, part of the available reaction exergy of this reaction is being lost according to Denbigh. If we are in some way able to directly couple this reference reaction 1 to the decomposition of water, i.e. reference reaction 2, the available reaction exergy decreases by a factor 3, thereby reducing Denbigh losses, and, moreover, hydrogen is produced. Another interesting possibility is the direct formation of methanol and hydrogen by reacting methane with water. This reaction has a negative available reaction exergy ( = - A , G °) of - 3.5 G J / t methanol. This reaction is an 'uphill' reaction and therefore needs to be supplied with energy, e.g. electricity. Resumptively, it can be concluded that changing chemical paths by direct coupling of reactions, enabled for example by fuel cells, can be very fruitful from an exergetic point of view.

in combination with partial oxidation. A major part of the determined overall exergy loss is associated with internal exergy consumption due to process irreversibilities. The extent of the determined exergy loss for the various processes can be matched with differences in process configuration. Improvement of exergy-utilization is mainly accomplished by reducing inefficiencies in combustion steps and lowering net heat production. External heat supply by uncontrolled combustion of high quality fuel in a furnace is found to be responsible for about 50% to the overall exergy loss during conventional steam reforming. On the contrary, the autothermal reforming method, involving internal combustion with pure oxygen shows improved exergy-utilization, despite additional losses accompanied by the need for air separation for provision of oxygen. An efficient reduction of the net heat production is also accomplished by the convective reforming option, in which the sensible heat content of process streams is used for additional conversion of natural gas. For the chemical reactions during synthesis gas production, improvements can be attained by reducing the extent of downhill reactions relative to reactions uphill. As a result, the available reaction exergy, playing a crucial role in the extent of the unavoidable exergy loss, is reduced. The practical exergy loss associated with an overall reaction is mainly determined by this unavoidable exergy loss. It is therefore recommended to reduce the available reaction exergy of the overall reaction. Losses are minimized when no net physical exergy is produced, but cannot, however, be eliminated completely when reactions are only thermally coupled. Further improvement will only be possible by direct coupling of reactions by their complete Gibbs energy of reaction. Application of fuel cells with special catalysts, enabling such a complete coupling of reactions, is therefore recommended from an exergetic point of view.

Acknowledgements--The authors would like to thank 7. CONCLUSIONS

The main objective of this part of the article is the presentation of the results of a project on computeraided exergy analysis of several synthesis gas processes in which exergies have been calculated by subroutines, integrated with a flowsheeting simulator and developed by the present authors. Exergy analysis is found to be a useful (but supplemental) diagnostic tool for process analysis. It prorides for an objective measure for selecting and judging processes with regard to their energy (or exergy) utilization. Also, potentials for improving this utilization can be revealed. Reductions of the overall exergy loss can be accomplished by using more sophisticated process designs. It is found that the overall exergy loss can be reduced from about 8.5 GJ/t of methanol yield for conventional steam reforming to approximately 4.9 GJ/t for the convective reforming process option

P. van Steenderen for his support to the exergy analysis project. His clear vision on exergy analysis has been found to be very fruitful. Financial support is acknowledged from the Netherlands agency for Energy and Environment (Novem B.V.). NOTATION

Exlo~s Ex + Ex,,+, Ex6 ExEXQ

Ex~seful,ms MeOH n NG

exergy loss, kJ/mol total exergy input, kJ/mol exergy input associated with material streams, k J/tool required amount of import exergy, kJ/mol total exergy output, kJ/mol export exergy in the form of heat, kJ/mol exergy output associated with useful material streams, kJ/mol methanol number of moles natural gas feedstock/fuel

4714 p+

Po POX P, R SN SR To

T,

A. P. HINDERINKet al. power input based on product yield, kJ/mol reference pressure, kPa partial oxidation reactor/reaction pressure, kPa gas constant, kJ/(mol K) stoichiometric number steam reforming ambient temperature; standard temperature, K, °C reactor/reaction temperature, K, °C

Greek symbols AEx exergy difference across a system, k J / m o l A,G Gibbs energy of reaction, k J / m o l A,G ° standard Gibbs energy of reaction, kJ/mol A,H ° standard heat of reaction, k J / m o l A,S° standard entropy of reaction, kJ/(mol K) exergetic efficiency of generating com/']ex,P pression power exergetic efficiency of supplying import r/ex,Q exergy exergetic efficiency of heat transfer proTex cesses summation Z Zv sum of stoichiometric coefficients (mole production) REFERENCES

Baade, W. F., Snyder, G. D. and Abrardo, J. M., 1993, Generate hydrogen for reformulated gasoline and clean diesel requirements. Hydrocarbon process. 1, 77-85. Christensen, T. S. and Primdahl, I. I., 1994, Improve syngas production using autothermal reforming. Hydrocarbon process. 3, 39-46. Denbigh, K. G., 1956, The second law efficiency of chemical processes. Chem. Enong Sci. 6(1), 1-9. de Swaan Arons, J. and van der Kooi, H. J., 1993, Exergy analysis, adding insight and precision to experience and intuition, in Precision Process Technology (Edited by M. P. C. Weynen and A. A. H. Drinkenburg) pp. 89-113. Kluwer Academic Publishers, Dordrecht, the Netherlands. English, A., Forbes, I. A. and McKee, D., 1990/1991, Synthesis gas production--the reforming route. Hydrocarbon Technol. International (Edited by Harrison, P.), pp. 115-118. Sterling Publications International. Farina, G. L. and Supp, E., 1992, Produce syngas for methanol. Hydrocarbon Process. 3, 77-81. Foreman, J. M., 1990, Pre-reformer aids syngas units. Hydrocarbon Process. 12, 34B-34D. Gaggioli, R. A., 1980, Thermodynamics: second law analysis. ACS Syrup. Set. 122. The American Chemical Society, Washington. Giacobbe, F. G., laquaniello, G. and Loiacono, O., 1992, Increase hydrogen production. Hydrocarbon Process. 3, 69-72. Goff, S. P. and Wang, S. I., 1987, Syngas production by reforming. Chem. Engn9 Prog. 8, 4653. Hedman, B. A., Brown, H. L. and Hamel, B. B., 1980, Second law analysis of industrial processes. Energy 5, 905-914. Johansen, T., Raghuraman, K. S. and Hackett, L. A., 1992, Trends in hydrogen plant design. Hydrocarbon Process. 8, 119-127. Kerkhof, F. P. J. M. and van Steenderen, P., 1993, Integration of gas turbine and air separation unit for IGCC

power plants; study on efficiency of IGCC under aspect of exergy. Netherlands Agency for Energy and Environment (Novem), Sittard. Kotas, T. J., 1985, The Exergy Method of Thermal Plant Analysis. Buttersworths, London. Maloney, D. P. and Burton, J. R., 1980. Using second law analysis for energy conservation studies in the petrochemical industry. Energy 5, 925-930. Moran, M. J., 1982, Availability Analysis: A Guide to Efficient Energy Use. Prentice-Hall, Englewood Cliffs, New Jersey. Rosen, M. A., 1991, Thermodynamic investigation of hydrogen production by steam-methane reforming. Int. J. Hydrogen Energy 16(3), 207-217. Sama, D. A., Sanhong, Q. and Gaggioli, R., 1989. A common-sense second law approach for improving process efficiencies, in Proceeding of International Symposium on Thermodynamic analysis and Improvement of Energy Systems (Edited by C. Ruixian and M. J. Moran), TAIES '89, Beijing (China), pp. 520-531. Pergamon Press, Oxford. Schneider, R. V. and LeBlanc, J. R., 1992, Choose optimal syngas route. Hydrocarbon Process. 3, 51-57. Supp, E., 1990, How to produce methanol from coal, pp. 107 110. Springer, Berlin. Supp, E., 1995, Improved methanol production and conversion technologies. Energy Pro9. 5(3), 127-130. Szargut, J., Morris, D. R. and Steward, F. R., 1988, Exergy Analysis of Thermal, Chemical and Metallurgical Processes. Hemisphere Publishing, New York. Tsatsaronis, G., Tawfik, T., Lin, L. and Gallaspy, D. T., 1991, Exergetic comparison of two KRW-based IGCC power plants, in Second Law Analysis--Industrial and Environmental Applications (Edited by G. M. Reistad), AES-25, pp. 7-18, The American Society of Mechanical Engineers, New York. Vannby, R. and Winter-Madsen, S. E. L., 1992, New developments in synthesis gas production. Hydrocarbon Technol. International (Edited by Harrison, P.), pp. 105 ll 1. Sterling Publications International. Yang, Y., Wang, Q., Ling, W. and Zeng, S., 1989, Thermodynamic analysis and evaluation of several typical ammonia synthesis flowsheets, in Proceedings of International Symposium on Thermodynamic Analysis and Improvement of Energy Systems (Edited by C. Ruixian and M. J. Moran) TAIES '89 Beijing (China), pp. 512-519. Pergamon Press, Oxford. APPENDIX A

A.1. Synthesis gas production by catalytic steam reforming The most common and well known way of producing synthesis gas is by conventional steam reforming where hydrocarbons are catalytically converted to carbon oxides and hydrogen by reaction with steam at temperatures up to 900°C and pressures below approximately 2000 kPa. All higher hydrocarbons are completely converted and methane conversion will be close to equilibrium. During steam reforming, heat for the endothermic reactions is externally supplied by combustion of fuel with air in a furnace. Generally, the steam to carbon ratio in the reformer feed ranges from 3 to 6, depending on the purpose for which the synthesis gas is produced. This ratio may not be too low due to the risk of coke-formation at high reformer at high reformer temperatures. A.2. Synthesis gas production by partial oxidation (autotherreal reforming) In the partial oxidation process, hydrocarbons react with a less than stoichiometric amount of pure oxygen. Reaction takes place according to two major steps. First, a portion of the hydrocarbon feedstock is oxidized in a combustion zone, thereby producing CO2 and H20 together with a large amount of heat and accompanying temperature rise. Secondly, when the temperature level is increased to a certain

Exergy analysis with a flowsheeting simulator----If. degree, the CO2 and H 2 0 are reformed with the remaining methane to form synthesis gas. Since the endothermic reforming reaction shows a large activation energy, a high reactor temperature is required to react methane with steam without a catalyst. Since no external heat supply is involved with partial oxidation processes, the reaction can be carried out in an adiabatic reactor, allowing operating at much higher pressures than possible with steam reforming. For non-catalytic partial oxidation, the operating temperature lies around 1400°C and no additional steam supply is needed for obtaining sufficient hydrocarbon conversion. In contrast, for catalytic partial oxidation (also denoted as oxygen reforming), additional steam supply is needed, since the operating temperature is limited to about 100WC. The advantage of eliminating external heat supply is partly overshadowed by the need for oxygen production by a costly air separation unit.

• • •

•

•

• APPENDIX B. SUMMARY OF THE METHOD OF SIMULATING/MODELLING THE SYNTHESIS GAS PROCESS • The Peng Robinson equation of state is used with interaction parameters taken from the Aspen Plus databank. For steam properties, the steam tables are applied. • For feedstock and fuel, a high caloric gas, containing amongst others about 87 vol% methane, has been used. This raw material is available at 6000 kPa and 25cC. • In order to obtain the desired feed pressure, feedstock and fuel are depressurized isenthalpically, thereby de-

•

•

4715

stroying part of the exergetic potential of these process inputs. All reactors are modelled as a Gibbs reactor to calculate the equilibrium composition. Air is modelled as dry air containing 79% nitrogen and 21% oxygen. Combustion in the reformer furnace takes place ideally with an excess a m o u n t of air giving an oxygen content of 4 vol% in the exhaust on dry basis. The outlet temperature of the reformer furnace is set to be 2OC higher than the temperature of hot product gases leaving the reformer to model heat transfer resistance. Air separation is modelled as a black-box unit operation. Oxygen recovery and purity is set to 95% and 98%,, respectively. Before entering the air separation unit, air is compressed to 600 kPa. The temperature of stack gases is set to 1 5 0 C before disposal to environment. The temperature of process condensate is set to 5 0 C . Except for oxygen compression, compressors are modelled as single stage compressors having a polytropic efficiency of 80%. Oxygen compression is accomplished by a 4-stage compressor, with interstage cooling capacity to 25-C between stages one and three. The convective reformers are simulated as several heaters and Gibbs reactors in series to model the shell and tube-side, respectively, since no model of a heat exchanger reactor is available in Aspen Plus.