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A novel hybrid absorption–compression refrigeration cycle
International Journal of Refrigeration 24 (2001) 208±219
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A novel hybrid absorption±compression refrigeration cycle J. Swinney, W.E. Jones *, J.A. Wilson School of Chemical, Environmental and Mining Engineering, University of Nottingham, University Park, Nottingham NG7 2RD, UK Received 22 November 1999; received in revised form 28 February 2000; accepted 8 March 2000
Abstract When used in traditional pool-boiling type refrigeration cycles, non-azeotropic mixed refrigerants tend to result in a reduced eciency compared to pure refrigerants. This results from the composition shift eect, which distributes the mixture components: concentrating the more volatile component in the high pressure part of the cycle, and the less volatile component in the low pressure part. The obvious eect of this is to increase the compression ratio relative to a single component. This article investigates a way of manipulating the composition change of a refrigerant mixture, using two components of similar volatility, in order to reduce the compression ratio. Counter-current vapour±liquid contact is used in a ``refrigeration column'', which is combined with a distillation column. The cycle is able to exploit heat sources below 100 C as input to the distillation column and the designer is able to optimise the consumption of compressor power and distillation heat input. # 2001 Elsevier Science Ltd and IIR. All rights reserved. Keywords: Refrigerating system; Refrigerant; Zeotropic mixture; Compression system; Absorption system; Hybrid system; Design
Nouveau cycle frigori®que aÁ absorption aÁ compression hybride ReÂsume La performance obtenue lorsqu'on utilise les meÂlanges de frigorigeÁnes azeÂotropes dans les cycles frigori®ques classiques aÁ eÂbullition libre est en geÂneÂrale diminueÂe si on la compare avec celle obtenue avec les frigorigeÁnes purs. Ce pheÂnomeÁne est duà aux changements de composition avec une distribution des composants variable : le composant le plus volatil est concentre dans la partie haute pression du cycle alors que le composant le moins volatil se trouve concentre dans la partie basse pression du cycle. L'eet eÂvident est d'augmenter le taux de compression vis-aÁ-vis d'un seul composant. Cette communication eÂtudie une facËon d'ajuster les changements de composition d'un meÂlange de frigorigeÁnes aÁ l'aide de deux composants d'une volatilite comparable, permettant ainsi de diminuer le taux de compression. On assure le contact vapeur-liquide aÁ contre-courant dans une « colonne frigori®que» combineÂe avec une colonne de distillation. Ce cycle utilise des sources de chaleur en dessous de 100 C pour la colonne de distillation. Le concepteur peut donc optimiser la consommation eÂnergeÂtique du compresseur et la chaleur requise pour le processus de distillation. # 2001 Elsevier Science Ltd and IIR. All rights reserved. Mots cleÂs : SysteÁme frigori®que ; FrigorigeÁne ; MeÂlange zeÂotrope ; SysteÁme aÁ compression ; SysteÁme aÁ absorption ; SysteÁme mixte ; Conception
1. Introduction * Corresponding author. Tel.: +44-115-951-4172; fax: +44115-951-4181. E-mail address: [email protected] (W.E. Jones).
Non-azeotropic mixed refrigerants [1] may be used to advantage in counter-current ¯ow-boiling refrigeration systems, where they can be used to provide a temperature-
0140-7007/01/$20.00 # 2001 Elsevier Science Ltd and IIR. All rights reserved. PII: S0140-7007(00)00025-6
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Nomenclature a COP COP L LVC MVC n N Pc Pe Psat QE QR V x y
Relative volatility Coecient of performance for vapour compression refrigeration Coecient of performance for absorption refrigeration Liquid molar ¯ow rate Less volatile component More volatile component Stage number Total number of stages Condenser pressure Evaporator pressure Saturated vapour pressure Refrigeration duty Heat input to the regenerator Vapour molar ¯ow rate Liquid phase composition (mol fraction MVC) Vapour phase composition (mol fraction MVC).
glide during evaporation and condensation. By matching the refrigerant temperature glide to that of the process ¯uid or cooling water in the heat exchangers, irreversibilities in heat transfer can be reduced, leading to energy
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saving [2]. However, in pool-boiling systems, the use of mixtures imposes a distinct penalty due to the composition shift, which tends to concentrate the more volatile components (MVC) at the high temperature (condenser) end and the less volatile components (LVC) at the low temperature (evaporator) end (see Fig. 1). This has the eect of tending to lower the evaporator pressure and raise the condenser pressure, thus increasing the compression ratio and reducing the coecient of performance (COP), relative to that given by a comparable pure ¯uid [3]. This can be easily demonstrated for a binary mixed refrigerant, where it can be shown that the calculated COP passes through a minimum at some point in the composition range, as shown in Fig. 2 for a mixture of propane (MVC) and isobutane (LVC). The situation described above, where the LVC tends to become concentrated in the cold end of the cycle and the MVC at the hot end, is the opposite of the ideal situation, which would be to concentrate the LVC at the hot (condenser) end of the cycle, and the MVC at the cold (evaporator) end. This would tend to lower the condenser pressure and raise the evaporator pressure for given temperatures, thus reducing the compression ratio and increasing the COP. One way to achieve this would be to introduce a counter-current ¯ow between liquid and vapour in a series of contacting stages, as shown in Fig. 3. If the vapour feed stream is signi®cantly richer in MVC than the
Fig. 1. Illustration of the mechanism giving rise to a composition shift in a pool-boiling refrigeration system. (a) A single pool-boiling evaporator. Liquid from the condenser is pressure reduced to feed the evaporator, which produces a vapour stream that feeds the compressor. The vapour stream and the liquid pool in the evaporator are in equilibrium, hence the liquid pool is richer in the less volatile component and poorer in the more volatile component then the vapour. The liquid stream from the condenser and the vapour feed to the compressor have the same composition. (b) A single pool-boiling evaporator with the addition of an economiser ¯ash. The process described above is now repeated twice, so that the vapour removed from the intermediate pressure ¯ash is richer, so that the vapour removed from the intermediate pressure ¯ash is richer in MVC (and poorer in LVC) than the vapour from the evaporator. Similarly the liquid pool in the evaporator is poorer in MVC (and richer in LVC) than the liquid pool in the intermediate pressure ¯ash drum. Fig. 1. ScheÂma du meÂcanisme de changement de composition dans un systeÁme frigori®que aÁ eÂbullition libre. (a) Evaporateur aÁ eÂbullition libre simple. La pression du liquide du condenseur est reÂduite avant son passage par l'eÂvaporateur qui produit un ¯ux de vapeur alimentant le compresseur. Le ¯ux de vapeur et le liquide contenu dans l'eÂvaporateur sont en eÂquilibre ; le liquide a donc une concentration plus eÂleveÂe en composant moins volatil et de la meÃme manieÁre a une concentration moins eÂleveÂe en composant plus volatil que la vapeur. Le ¯ux liquide sortant du compresseur et le ¯ux de vapeur alimentant le compresseur ont la meÃme composition. (b) Un eÂvaporateur aÁ eÂbullition libre simple muni d'un eÂconomiseur. Le proceÂde deÂcrit ci-dessus est reÂpeÂte deux fois, de telle sorte que la vapeur enleve du ballon de deÂtente aÁ pression intermeÂdiaire est plus riche en MVC et moins riche en LVC que la vapeur en provenance de l'eÂvaporateur. De la meÃme manieÁre, le lquide dans l'eÂvaporateur a une concentration en MVC moins eÂleveÂe et une concentration en LVC plus eÂleveÂe que le liquide dans le ballon de deÂtente aÁ pression intermeÂdiaire.
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Fig. 2. Variation of calculated COP with refrigerant charge composition for a single stage pool-boiling refrigeration system, using a binary propane/isobutane refrigerant. Fig. 2. Variation du COP calcule selon la composition de la charge pour un systeÁme frigori®que monoeÂtage aÁ eÂbullition libre, pour un frigorigeÁne binaire propane/isobutane.
liquid feed stream, then this will tend to impose the desired composition and temperature pro®le on the stages. This assumes good mixing between liquid and vapour in each stage to promote ecient mass and heat transfer, ideally leading to equilibrium between the exiting streams. A novel cycle has been devised to exploit countercurrent contacting, and this article describes the cycle which can be viewed as a hybrid of traditional vapour compression and absorption refrigeration. In theory absorption refrigeration is very attractive for integration with process plant requiring a refrigeration service because expensive compressor power input can be replaced by low level heat recovered from the process. The low level heat is often viewed as free, i.e. no cost, because if not used as heat input to the absorption refrigeration regenerator, then it would be rejected to cooling water. However, the absorption refrigeration cycle is a more complex facility than vapour compression and generally requires a bigger capital investment, hence the additional capital investment must be justi®ed by the saving on compressor power. Most process plant rejects heat at a variety of temperature levels and the main concern is to match the heat released to possible users. Generally, less heat will be available at higher temperatures and this is viewed as more valuable because it can be directly recovered for process heating. This leaves the more plentiful lower temperature heat to be rejected to drive an absorption refrigeration system. It is important for the designer to note that the coecient of performance, COP (de®ned
Fig. 3. Illustration of how counter-current liquid/vapour contacting could be used to achieve the desired composition/temperature pro®le. Fig. 3. ScheÂma de l'utilisation de la mise en contact aÁ contrecourant du liquide et de la vapeur a®n d'obtenir le pro®l composition/tempeÂrature deÂsireÂ.
as the refrigeration duty, QE, divided by the heat input to the regenerator, QR), decreases when using lower level heat to drive the refrigeration unit. This means quite large heat loads can be involved, but this is not a problem because the heat can be viewed as free. The new cycle, being a hybrid, contains features of both vapour compression and absorption refrigeration, and hence appears slightly more complex. Both absorption refrigeration and the new cycle contain a regenerator/ distillation column and the recovered heat is used to drive these units. 2. Analysis In order to understand how the counter-current contacting system would operate, it is necessary to consider the behaviour of a typical single stage n, as shown in Fig. 4(a). The analysis considers a binary mixture, and it is assumed that equilibrium is achieved between liquid and vapour streams leaving the stage. Many standard textbooks deal with equilibrium stage calculations; a comprehensive treatment is given by Smith [4]. The analysis is performed on an equilibrium diagram as shown in Fig. 4(b). The equilibrium line relates molar compositions of the MVC in the liquid, x, and vapour,
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Fig. 4. Representation of a single contacting stage and associated compositions of entering and exiting streams: (a) a typical contacting stage; (b) the four stage compositions shown on an equilibrium diagram. Fig. 4. ScheÂma d'un eÂtage aÁ contact unique, avec les compositions des ¯ux entrants et sortants : (a) un eÂtage aÁ contact typique ; (b) les quatre compositions d'eÂtage sous forme de diagramme.
y, that are in equilibrium. Hence the compositions xn and yn of the two exiting streams from stage n must lie on this line, because they are assumed to be in equilibrium. Note also that the vapour is richer in MVC than the liquid, hence yn>xn. The shape of the equilibrium line in Fig. 4(b) is typical for a near ideal binary mixture (e.g. light hydrocarbons). Relative volatility, a, is a parameter commonly used to characterise the shape of the equilibrium line. Equilibrium liquid and vapour compositions, such as xn and yn, are then related by: yn
xn 1 1xn
where
Psat;MVC Psat;LVC
1 2
Psat,MVC and Psat,LVC are the vapour pressures of the pure components measured at the same temperature. For a near ideal mixture, a does not vary very much with temperature. In order to establish the desired composition pro®le, it is necessary to transfer MVC from the vapour stream into the liquid stream, and LVC from the liquid stream into the vapour stream. To achieve this, the entering vapour stream Vn1 must be signi®cantly richer in MVC than the entering liquid stream Ln 1 (i.e. yn1 > xn 1 ). Finally, to achieve the desired composition pro®le, whereby the content of MVC increases as temperature decreases (i.e. as n increases), it is required that xn > xn 1 and yn1 > yn . Hence, it can be said that yn1 > yn > xn > xn 1 . The relative positions of the four stage compositions are shown on Fig. 4(b). It will be noted from this that
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pairs of compositions for passing streams ([xn 1 ; yn ] and [xn ; yn1 ]) lie above the equilibrium line, as opposed to the situation in distillation where compositions of passing streams are below the equilibrium line. Moving on to consider a series of contacting stages as shown in Fig. 5(a), it becomes apparent that the composition change through the stages can be represented diagrammatically in the same way that a McCabe± Thiele diagram can be used to represent composition change in a distillation column [4]. An absorption heat pump using the above concept and graphical construction has been described elsewhere by Le Go [5,6] and co-workers. A further important assumption here is that of constant molar over¯ow [4]; this is approximately true for many mixtures such as light hydrocarbons, and it means that vapour and liquid ¯ows will be constant through a series of equilibrium stages, i.e. Vn=Vn+1=V and Ln 1=Ln=L. Fig. 5(b) shows a graphical construction for ®ve equilibrium contacting stages. The compositions of passing streams, e.g. [xn 1 ; yn ] and [xn ; yn1 ], are related by the operating line. The mathematical form of the operating line is easily determined by performing a molar balance on the MVC over part of the cascade of stages. For example, considering stages 1 to n the following relationship is obtained: yn1
L xn y1 V
L xF V
3
This shows the points representing pairs of passing streams lie on a straight line with gradient L/V. A physical interpretation can now be given to Fig. 5(b). The equilibrium line is drawn for a relative volatility of 3; this corresponds approximately to a mixture of propane (MVC) and isobutane (LVC) over a temperature range from 20 to 20 C. Both feeds to the series of contacting stages have been taken as 98% pure; the top (saturated liquid) feed stream contains 2 mol% MVC and the bottom (saturated vapour) feed stream contains 98 mol% MVC. The gradient of the operating line is 1, i.e. equal liquid and vapour ¯ows are to be used through the ®ve stages. Table 1 shows the vapour and liquid compositions for each stage, taken from Fig. 5(b). This illustrates more clearly how the aim of establishing a composition pro®le has been achieved by transferring the MVC from the vapour to the liquid and the LVC from the liquid to the vapour. The MVC rich vapour feed has thereby been converted to a vapour product that is poorer in MVC and richer in LVC, and is thus more easily condensed, i.e. for a given condensing temperature a lower pressure will be required, which was the initial aim. The greater the composition change that can be achieved, the greater this bene®t will be.
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Fig. 5. Representation of a series of contacting stages and associated compositions. (a) A series of N equilibrium stages. Stage n is a typical equilibrium stage. The liquid stream becomes richer in MVC as it ¯ows down through the stages, and the vapour stream becomes depleted in MVC as it ¯ows up. (b) A graphical construction for a series of ®ve equilibrium stages, drawn for a binary mixture with a relative volatility of 3. Fig. 5. ScheÂma d'une seÂrie d'eÂtages aÁ contact avec les compositions associeÂes. (a) Une seÂrie d'eÂtages d'eÂquilibre N. L'eÂtage n est un eÂtage d'eÂquilibre typique. Le ¯ux liquide en eÂcoulement descendant devient plus riche en MVC au fur et aÁ mesure qu'il coule aÁ travers les eÂtages, et le ¯ux de vapeur a une concentration de moins en moins eÂleveÂe en MVR au fur et aÁ mesure qu'il poursuit son ascension. (b) ScheÂma graphique d'une seÂrie de cinq eÂtages en eÂquilibre pour un meÂlange binaire avec une volatilite relative de 3.
Table 1 Comparison between composition pro®les from graphical construction and computer simulation, and summary of stage temperatures from the simulation Tableau 1 Comparaison entre les pro®ls de composition graphiques et la simulation informatique. Les tempeÂratures des eÂtages sont reÂsumeÂes aÁ partir de la simulation Stage
Top feed 1 2 3 4 5 Bottom feed a
Compositions (mol% MVC) Graphical construction
Computer simulation
x
y
x
y
2 17 27 37 47 62 ±
± 38 53 63 73 83 98
2 17 28 38 48 63 ±
± 35 51 63 73 83 98
Assumed saturated.
Temperature ( C) (from computer simulation)
16.1a 8.7 4.0 0.3 3.3 7.8 15.4a
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Also shown in Table 1 are the results (stage compositions and temperatures) of a computer simulation of the propane/isobutane system, produced using Hyprotech's HYSYS package. This shows two important results. Firstly, it con®rms that the compositions are very close to those predicted by the graphical construction in Fig. 5(b). Secondly, at an operating pressure of 2.75 bar, a temperature lift of 16.5 C was obtained between top and bottom stages (with a bottom stage temperature of 7.8 C and a top stage temperature of 8.7 C). Whilst this temperature lift is insucient to be of practical use, it demonstrates the principle that with a mixed refrigerant, composition variation alone (no pressure change is considered at this stage) can provide a temperature lift using a series of counter-current contacting stages. 3. Practical considerations So far, the use of countercurrent vapour/liquid contacting through a series of equilibrium stages in order to establish a composition/temperature pro®le has been demonstrated. Practical implementation of the scheme has not been considered. In common with other stagewise equilibrium processes, such as distillation, this process could most easily be carried out in a packed or plate column. The individual contacting stages would therefore be implemented as either packing or plates within the column designed to promote ecient mass/heat transfer. For the column to be of practical application in a refrigeration cycle, a number of other factors must also be taken into account. Firstly, the provision of evaporation and condensation units has not yet been considered. Secondly a greater temperature lift is needed. For example, a typical refrigeration duty might require an evaporator temperature of 25 C, and the cooling water available would determine the minimum condenser temperature, typically 40 C. This implies a temperature lift of 65 C. Finally, in order to generate the column composition pro®le, it is necessary to provide relatively pure feed streams to the top and bottom of the refrigeration column. Therefore, a refrigerating column would need to form part of a larger system, which would include a separation operation to provide the top and bottom feed streams. 3.1. Provision of evaporator and condenser In order to maximise the temperature lift, the evaporator and condenser must be positioned as shown in Fig. 6. This allows us to claim an increased temperature lift of 24.1 C, from 15.4 C (saturated vapour leaving the evaporator and entering the bottom stage) to 8.7 C (saturated vapour leaving the top stage). However, the bubble point temperature of the exiting vapour stream is only 1.2 C at the column pressure, which is the tempera-
Fig. 6. Illustration of the arrangement for the column and auxiliaries. Fig. 6. ScheÂma de la con®guration de la colonne et des composants annexes.
ture to which it would have to be cooled to complete condensation. This is obviously too low for condensation against cooling water or air, so a greater temperature lift is still needed. 3.2. Increasing the temperature lift The two most obvious ways to achieve this are either to introduce a pressure increase between evaporator and condenser, as in a normal vapour compression refrigeration cycle, or to extend the temperature lift resulting from the composition pro®le. 3.2.1. Pressure Raising the pressure between evaporator and condenser might be performed in a number of ways, depending on where the pressure change was introduced, as shown in Fig. 7. Considering a system with two pressure levels, three variations are possible: ®rstly to split the column into two parts operating at dierent pressures; secondly to introduce the pressure variation between the column and the condenser; or thirdly to put it between the column and the evaporator. For practical application, the con®guration of Fig. 7(b) Ð compressing the overhead vapour from the refrigerating column before feeding it to the condenser Ð is considered best. The other two con®gurations shown in Figs. 7(a) and (c) are less suitable because they both
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Fig. 7. Possible arrangements for introducing a pressure change between the evaporator and condenser of the refrigeration column: (a) pressure change within column; (b) pressure change between column and condenser; (c) pressure change between column and evaporator. Fig. 7. Con®gurations possibles a®n d'obtenir un changement de pression entre l'eÂvaporateur et le condenseur de la colonne frigori®que : (a) changement de pression aÁ l'inteÂrieur de la colonne ; (b) changement de pression entre la colonne et l'eÂvaporateur.
place the compression heat within the column, increasing the vapour ¯ow and thus requiring a larger column. 3.2.2. Composition To increase the temperature lift resulting from composition change would require maximising the composition change between evaporator and condenser, and/or the use of a mixture with a wider dierence in volatility between the MVC and the LVC. Considering the propane/isobutane system at 2.75 bar studied previously, the boiling points of the pure components are 16.6 and
17.2 C respectively, so even if the change in composition could be fully exploited, the temperature lift would be insucient (33.8 C). Using a mixture with a greater dierence in volatility is not a solution either. If the equilibrium line is constructed for a=5, then the available working region above the equilibrium line is much smaller and this in turn restricts the usable composition range, thus defeating the object of using a pair of components with a higher relative volatility. Hence the pressure change alternative is pursued as the best means to increase the temperature lift.
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Fig. 8. One possible con®guration for incorporating the refrigeration column into system with a distillation column. Fig. 8. Con®guration possible permettant d'inteÂgrer la colonne frigori®que dans un systeÁme aÁ colonne de distillation.
3.3. Completing the cycle In order to supply the relatively high purity feed streams needed to operate the refrigeration column, it would be necessary to run it in conjunction with a separation operation. A distillation column is an eective way to perform the separation needed and hence becomes an integral part of the cycle. The complete cycle is illustrated in Fig. 8. The distillation column top product (MVC rich) would form the feed to the evaporator, where it is completely vaporised before passing to the bottom of the refrigeration column. The distillation column bottom product (LVC rich) would feed the top of the refrigeration column. To develop a design the assumption of 98 mol% propane in the distillation column top product and 98 mol% isobutane in the bottoms will be maintained. The cycle is completed by returning the refrigeration column overhead vapour and bottoms liquid to appropriate locations in the distillation column. An important parameter in designing a distillation column is the re¯ux ratio. This de®nes the ratio in which the liquid from the condenser is divided for return to the column as re¯ux and taken as top product. In theory there is a minimum re¯ux ratio which leads to the requirement for a very large number of contacting stages in the distillation column. For a practical design, a commonly used heuristic suggests a re¯ux ratio of around 1.25 times the minimum re¯ux ratio [7]. It is also
important to introduce the feed(s) onto the correct stage(s) of the distillation column. In the design outlined below, and the Sensitivity Study in the next section, the 1.25 factor has been used and the feed location optimised to put all the designs on a consistent basis. The highest pressure needed in the cycle will be that of the distillation unit, which will be determined by the minimum temperature (say 40 C) obtainable with cooling water. The pressure is ®xed by the requirement to completely condense the overhead vapour, i.e. a propane/isobutane mixture containing 98 mol% propane has a bubble-point pressure of 13.5 bar at 40 C. Similarly, the pressure of the refrigeration evaporator and column is ®xed by the requirement for the evaporator to produce a saturated vapour at 25 C, at which temperature the dew-point pressure for the above mixture is 1.9 bar. The distillation column product streams therefore have to be let down from 13.5 to 1.9 bar. If these streams are not subcooled before being let down, approximately 40% of both streams ¯ashes to vapour, which is a problem since it reduces the refrigerating capacity of the evaporator whilst increasing the load on the compressor and condenser. The bottom product stream from the distillation column will be hot enough to be cooled against cooling water (to 40 C), in the heat exchanger labelled ``pre-cooler (water)'' in Fig. 8, before being let down, thus reducing the ¯ash to around 25%. However, since the top product stream is already at 40 C, it can only be further cooled using refrigeration
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recovery, in the heat exchanger labelled ``pre-cooler (refrigerant)''. The ¯ash for this stream is thus reduced to approximately 15%. The return of the top vapour and bottom liquid streams from the refrigeration column to the distillation column requires power input. The bottom liquid is easily pumped from 1.9 up to 13.5 bar, but the top vapour need not be compressed beyond its bubble-point pressure at 40 C, in this case 8.2 bar. This represents a signi®cant saving on compressor power. Once condensed, the liquid stream is easily pumped to 13.5 bar. The performance of the cycle discussed above and illustrated in Fig. 8 is summarised in Table 2. It is worth noting that the highest temperature in the cycle, to be found at the bottom of the distillation column, is 79 C. This is approximately 100 C cooler than the distillation temperature for a simple ammonia-water absorption refrigeration cycle operating under similar conditions. Included in Table 2 are two new terms, derived from coecient of performance. Power ratio is calculated as the refrigeration heat load divided by the compressor power, and heat ratio as the refrigeration heat load divided by the distillation column reboiler duty. These two terms are useful for later comparison purposes.
4. Sensitivity study The idea of introducing the refrigeration column with two volatile working ¯uids is new so it is worth examining the sensitivity of the design to changes in the following important parameters:
Table 2 Performance of the refrigeration cycle as illustrated in Fig. 8 Tableau 2 Performance du cycle frigori®que illustre en Figure 8 Refrigerant evaporator
Temperature ( C) Pressure (bar) Heat load (kW)
25 1.9 1000
Refrigerant condenser
Bubble-point temperature ( C) Pressure (bar) Heat load (kW)
40 8.2 1373
Work input
Compressor (kW) Pump 1 (kW) Pump 2 (kW) Total (kW)
328 7 6 341
Refrigeration Number of stages column Pressure (bar) Bottom stage temperature ( C) Top stage temperature ( C)
5 1.9 19 2
Distillation column
32 13.5 79 40 2848 2410
Number of stages Pressure (bar) Reboiler temperature ( C) Condenser bubble-point temperature ( C) Reboiler heat load (kW) Condenser heat load (kW)
Power ratio
2.93
Heat ratio
0.35
Fig. 9. Eect of varying the number of stages in the refrigeration column, keeping the composition of the feed streams ®xed (98% pure), on: (a) heat and mechanical work input; (b) composition of streams exiting refrigeration column. Fig. 9. Eet de la variation du nombre d'eÂtages dans la colonne frigori®que, en maintenant une concentration des ¯ux entrants stable (98 %) sur : l'eÂnergie thermique et le travail meÂcanique ; (b) la composition des ¯ux sortant de la colonne frigori®que.
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. Number of contact stages in the refrigeration column, whilst maintaining both feed streams at 98 mol% purity. . Purity of the feed streams to the refrigeration column whilst maintaining ®ve contacting stages in the column.
Fig. 9(a) shows that as the number of stages in the refrigeration column is increased, both heat input to the distillation reboiler and work input to the compressor decrease. This is to be expected, because with a greater number of stages in the refrigeration column, the irreversibilities are reduced, enabling a greater change in composition to be obtained through the column, as shown in Fig. 9(b). The overhead vapour from the column will become richer in LVC, so its bubble-point pressure will fall, which reduces the work input to the compressor. Similarly, the bottom liquid stream will become richer in MVC. With higher purity feed streams being returned to the distillation column, the separation will become easier, reducing the reboiler heat load. The number of stages in the distillation column does not change signi®cantly, but the re¯ux ratio falls slightly as the number of stages in the refrigeration column is increased. With reduced vapour trac through the distillation column, a smaller diameter column will be needed. Therefore, increased capital expenditure on the refrigeration column is accompanied by a reduction in utility costs and also a reduction in capital expenditure on the distillation column. Fig. 10 shows the results of changing the purity of the feed streams to the refrigeration column, the number of stages in the column being ®xed at 5. For simplicity, the composition of both feed streams has been changed in parallel. For example, 90% purity means that the bottom feed stream contains 90 mol% MVC and 10 mol% LVC, and the top feed stream contains 90 mol% LVC and 10 mol% MVC. Fig. 10(a) shows that as the feed purity is decreased from its initial value of 98%, the compressor work input rises and the reboiler heat load falls. With less pure feed streams, the available composition change in the refrigeration column is reduced, so the overhead vapour stream becomes poorer in LVC, increasing its condensing pressure and hence the compressor work input. Lower purity feed streams means a less rigorous separation is required, so less heat input is needed to the distillation column reboiler. This illustrates a trade-o between heat and work input. Fig. 10(b) shows the eect on the distillation column of changing the purity of its product streams. As the purity is reduced, the number of contacting stages falls, as does the reboiler temperature. With a ®xed condenser temperature, the column pressure also decreases as the overhead product stream becomes poorer in MVC. Therefore, reducing the purity of the feed streams to the refrigeration column also implies a reduction in capital expenditure on the distillation column.
Fig. 10. Eect of varying the purity of the feed streams to the refrigeration column, with ®ve stages, on: (a) heat and mechanical work input; (b) number of stages in the distillation column, and distillation column reboiler temperature. Fig. 10. Eet des variations de concentration des ¯ux entrant dans la colonne frigori®que, avec cinq eÂtages, sur : l'eÂnergie thermique et le travail meÂcanique ; (b) le nombre d'eÂtages dans la colonne de distillation, et la tempeÂrature du rebouilleur de la colonne de distillation.
5. Comparison to standard refrigeration cycles The proposed cycle is a hybrid so it is interesting to make a comparison with absorption and vapour compression refrigeration. The new cycle is closest to absorption refrigeration, so this is considered ®rst. The comparison
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needs to be made under two headings, namely complexity (as an indication of capital investment) and utilities consumption (as an indication of operating costs). The ammonia±water cycle has a refrigerant operating range similar to the new cycle and has been extensively investigated, hence it forms a good basis for comparison. Ammonia absorption refrigeration cycles can vary widely in complexity [8,9]. To start the comparison, the simplest absorption refrigeration unit is considered, i.e. single stages for evaporation, absorption and regeneration. Table 3 makes a comparison on the number of items in the cycles and their direct equivalence. As may be noted each item in an absorption refrigeration cycle has a direct equivalent in the new cycle with the exception of the compressor, which is an additional item. In order to understand why the new cycle is of interest, despite the addition of a compressor, it is necessary to compare the operating capabilities of absorption refrigeration to those of the new cycle.
Table 3 Equipment comparison between ammonia absorption refrigeration cycle and new cycle Tableau 3 Comparaison du cycle frigori®que aÁ absorption aÁ ammoniac et le nouveau cycle Simple ammonia absorption New cycle refrigeration cycle Regenerator (includes reboiler and condenser) Absorber±condenser Evaporator ± One pump Two heat exchangers (Two interchangers)
Distillation column (includes reboiler and condenser) Refrigeration column condenser (refrigerant) Evaporator Compressor Two pumps Two heat exchangers (interchanger, cooler)
Table 4 Comparison of absorption refrigeration with new cycle Tableau 4 Comparaison du cycle frigori®que aÁ absorption et le nouveau cycle Boiling refrigerant at ( C)
Description
Heat recoverable to ( C)
Single stage New cycle
150 85
25 25
Dual pressure regeneration A Dual pressure regeneration B
106
25
85
+2
COP or power/ heat ratio COP 0.4 Power ratio 2.93 Heat ratio 0.35 COP 0.26 COP 0.34
Briley [10] provides a useful table detailing the main operating characteristics for some alternative ammonia absorption refrigeration designs. Table 4 is derived from this work and is intended to summarise alternative designs which can be compared with the new cycle. As may be seen from Table 4, to provide boiling refrigerant at 25 C, a simple ammonia absorption refrigeration unit would need heat at 150 C or greater, and operate with a COP of approximately 0.4. These numbers are broadly con®rmed by Holldorf [9] and Borsig [11], and can be compared with the requirements of the new cycle, namely heat available at or above 85 C and a heat ratio of 0.35. However, a compressor and its power input is also needed; this might be viewed as the penalty paid to permit heat recovery to such low temperatures. The ability to recover low level heat is an important feature of the new cycle, so consideration should be given to the changes which are necessary to permit an ammonia absorption unit to operate on colder heat supplies. Bogart [8] and Holldorf [9] describe two approaches: . multiple reboilers on the regenerator, each reboiler located higher up the column and operating at successively lower temperatures; . dividing the regenerator into two stages, with the water rich section operating at a lower pressure and hence capable of accepting heat at a lower temperature.
Both alternatives add to the complexity and hence cost of the unit and in particular the dierence between the two stage regeneration and the new cycle is not clear cut. Two stage regeneration is the most frequently quoted route for reducing the required temperature of the heat source and useful data is given by Briley [10]. This has been used to generate two alternative designs given in Table 4. These alternatives have been chosen to highlight performance dierences compared to the new cycle. In Table 4, dual pressure A shows that the minimum heat source temperature feasible to produce 25 C refrigerant is 106 C and the COP has dropped signi®cantly to 0.26. To highlight the importance of heat source temperature, dual pressure B shows that taking heat down to 85 C would permit a refrigeration temperature of 2 C and a COP of 0.34. Clearly the new cycle is able to operate at more extreme conditions but the penalty paid is the addition of the compressor. It is probably worth observing that a refrigeration temperature of 25 C is at the cold extreme for absorption refrigeration. Absorption refrigeration is much more eective for refrigeration temperatures around 0 C and one would anticipate similar performance improvement for the new cycle. Obviously, the refrigeration column pressure would be higher leading to less ¯ash on liquid pressure reduction and reduced compressor power consumption. Finally, to complete the comparison vapour compression refrigeration needs to be considered. Here pure
J. Swinney et al. / International Journal of Refrigeration 24 (2001) 208±219
propane is considered and the COP for a cycle operating with an evaporator at 2.03 bar and condenser at 13.7 bar is 1.88. This value should be compared with the power ratio of 2.93 for the new cycle. 6. Comparison with previous work As noted previously Le Go and co-workers [5,6,12,13] have studied absorption heat pumps using ``reverse recti®cation''. However, the binary mixtures considered (e.g. water/ethylene glycol, ethylamine/water and pentane/octane) are much further apart in terms of volatility than our example. In such cases a straight operating line, as used in this work, would permit only a very restricted composition change. Le Go and coworkers create a change in operating line gradient by a combination of heat removal and withdrawing liquid from the middle section of the column. The heat removed is the up-graded energy from the heat pump but this is obviously not feasible for the refrigeration column discussed above because the entire refrigeration column operates below ambient temperature. Heat pumps combining absorption and compression are investigated by a number of authors. The systems considered all seem to be based on standard absorption cycles using either ammonia/water or water/lithiumbromide refrigerant/absorbent pairs, with vapour compression added at some point in the cycle to improve either eciency or ¯exibility. The development of these systems is reviewed by Morawetz [14] and Ahlby and Hodgett [15]. One variant of these combined cycles is analysed by Ziegler and Spindler [16], who look at optimising the performance of a single stage vapour compression refrigeration cycle by adding an economiser ¯ash. The use of a second compression stage to deal with the intermediate pressure ¯ash vapour is compared with the use of an absorption circuit. The compression plus absorption con®guration is shown to be advantageous because the heat input to the absorption circuit can be supplied solely by the superheat from the compressor discharge. Experimental results are presented for the compression plus absorption cycle. Aly [17] looks at a water/ lithiumbromide absorption cycle where vapour recompression is used in place of the conventional condenser. This system is aimed at high temperature use, not refrigeration. 7. Conclusions The use of composition change with a mixed refrigerant to achieve a temperature lift has been examined. It has been demonstrated theoretically how this may be achieved by using a series of equilibrium stages with countercurrent vapour/liquid contacting. These stages could be arranged in column format. A simple graphical
219
analysis of the system along the lines of a McCabe±Thiele construction for distillation has been presented for the case of a binary mixture. The possibility of integrating the column into a closed cycle has been examined, along with the practical implications of this in terms of energy requirements. Performance is shown to be comparable to conventional absorption refrigeration units bearing in mind the possibility of recovering much lower level heat. The designer has considerable ¯exibility over the use of shaft power and recovered low level heat. Like all absorption refrigeration based systems, the new cycle is best suited to situations where waste heat at low temperatures is readily available.
References [1] Didion DA, Bivens DB. Role of refrigerant mixtures as alternatives to CFCs. Int J Refrig 1990;13:163±75. [2] Bensa® A, Haselden GG. Wide-boiling refrigerant mixtures for energy saving. Int J Refrig 1994;17(7):469±74. [3] Swinney J, Jones WE, Wilson JA. The impact of mixed non-azeotropic working ¯uids on refrigeration system performance. Int J Refrig 1998;21(8):607±16. [4] Smith BD. Design of equilibrium stage processes. McGraw-Hill, 1963 (chapter 5). [5] Ranger PM, Matsuda H, Le Go P. Modelling of a new type of absorption heat pump combining recti®cation and ``reverse-recti®cation''. J Chem Eng Japan 1990;23(5):530±6. [6] Ranger PM, Matsuda H, Le Go P. Modelling and experimental operation of a ``reverse-recti®cation column'' for an absorption heat pump. J Chem Eng Japan 1990;23(5):537±42. [7] Yaws CL, Li KY, Fang CS. How to ®nd the minimum re¯ux for binary systems in multiple-feed columns. Chem Eng 1981;88(10):153±6. [8] Bogart M. Ammonia absorption refrigeration in industrial processes. Houston (TX): Gulf Publishing Company, 1981. (chapter 17) [9] Holldorf G. Revisions up absorption refrigeration eciency. Hydrocarbon Processing. July 1979;58:149±155. [10] Briley G.C. Conserve energy. . . refrigerate with waste heat. Hydrocarbon Processing. May 1976;55:173±174. [11] Anon. Borsig pocket book. 5th ed. p. 336 [12] Le Go P., Labidi J. The concept of `reverse-fractionaldistillation' exergy analysis and applications. IChemE Symposium Series 1992; 128 B55-B67. [13] Labidi J, Schwarzer BP, Le Go P. Absorption heat pumps composed of multiple stages either independent or belonging to a unique column. Int Absorption Heat Pump Conf, ASME 1993;31:251±6. [14] Morawetz E. Sorption-compression heat pumps. Int J Energy Res 1989;13(1):83±102. [15] Ahlby L, Hodgett DL. Compression±absorption heat pumps. In: 3rd Int. Energy Agency Heat Pump Conf. Oxford: Pergamon Press, 1990. [16] Ziegler F, Spindler U. An ammonia refrigerator with an absorption circuit as economizer. Int J Refrig 1993;16(4):230±9. [17] Aly SE. Energy-ecient absorption/mechanical vapourrecompression system. J Inst Energy 1991;64(458):36±40.
www.elsevier.com/locate/ijrefrig
A novel hybrid absorption±compression refrigeration cycle J. Swinney, W.E. Jones *, J.A. Wilson School of Chemical, Environmental and Mining Engineering, University of Nottingham, University Park, Nottingham NG7 2RD, UK Received 22 November 1999; received in revised form 28 February 2000; accepted 8 March 2000
Abstract When used in traditional pool-boiling type refrigeration cycles, non-azeotropic mixed refrigerants tend to result in a reduced eciency compared to pure refrigerants. This results from the composition shift eect, which distributes the mixture components: concentrating the more volatile component in the high pressure part of the cycle, and the less volatile component in the low pressure part. The obvious eect of this is to increase the compression ratio relative to a single component. This article investigates a way of manipulating the composition change of a refrigerant mixture, using two components of similar volatility, in order to reduce the compression ratio. Counter-current vapour±liquid contact is used in a ``refrigeration column'', which is combined with a distillation column. The cycle is able to exploit heat sources below 100 C as input to the distillation column and the designer is able to optimise the consumption of compressor power and distillation heat input. # 2001 Elsevier Science Ltd and IIR. All rights reserved. Keywords: Refrigerating system; Refrigerant; Zeotropic mixture; Compression system; Absorption system; Hybrid system; Design
Nouveau cycle frigori®que aÁ absorption aÁ compression hybride ReÂsume La performance obtenue lorsqu'on utilise les meÂlanges de frigorigeÁnes azeÂotropes dans les cycles frigori®ques classiques aÁ eÂbullition libre est en geÂneÂrale diminueÂe si on la compare avec celle obtenue avec les frigorigeÁnes purs. Ce pheÂnomeÁne est duà aux changements de composition avec une distribution des composants variable : le composant le plus volatil est concentre dans la partie haute pression du cycle alors que le composant le moins volatil se trouve concentre dans la partie basse pression du cycle. L'eet eÂvident est d'augmenter le taux de compression vis-aÁ-vis d'un seul composant. Cette communication eÂtudie une facËon d'ajuster les changements de composition d'un meÂlange de frigorigeÁnes aÁ l'aide de deux composants d'une volatilite comparable, permettant ainsi de diminuer le taux de compression. On assure le contact vapeur-liquide aÁ contre-courant dans une « colonne frigori®que» combineÂe avec une colonne de distillation. Ce cycle utilise des sources de chaleur en dessous de 100 C pour la colonne de distillation. Le concepteur peut donc optimiser la consommation eÂnergeÂtique du compresseur et la chaleur requise pour le processus de distillation. # 2001 Elsevier Science Ltd and IIR. All rights reserved. Mots cleÂs : SysteÁme frigori®que ; FrigorigeÁne ; MeÂlange zeÂotrope ; SysteÁme aÁ compression ; SysteÁme aÁ absorption ; SysteÁme mixte ; Conception
1. Introduction * Corresponding author. Tel.: +44-115-951-4172; fax: +44115-951-4181. E-mail address: [email protected] (W.E. Jones).
Non-azeotropic mixed refrigerants [1] may be used to advantage in counter-current ¯ow-boiling refrigeration systems, where they can be used to provide a temperature-
0140-7007/01/$20.00 # 2001 Elsevier Science Ltd and IIR. All rights reserved. PII: S0140-7007(00)00025-6
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Nomenclature a COP COP L LVC MVC n N Pc Pe Psat QE QR V x y
Relative volatility Coecient of performance for vapour compression refrigeration Coecient of performance for absorption refrigeration Liquid molar ¯ow rate Less volatile component More volatile component Stage number Total number of stages Condenser pressure Evaporator pressure Saturated vapour pressure Refrigeration duty Heat input to the regenerator Vapour molar ¯ow rate Liquid phase composition (mol fraction MVC) Vapour phase composition (mol fraction MVC).
glide during evaporation and condensation. By matching the refrigerant temperature glide to that of the process ¯uid or cooling water in the heat exchangers, irreversibilities in heat transfer can be reduced, leading to energy
209
saving [2]. However, in pool-boiling systems, the use of mixtures imposes a distinct penalty due to the composition shift, which tends to concentrate the more volatile components (MVC) at the high temperature (condenser) end and the less volatile components (LVC) at the low temperature (evaporator) end (see Fig. 1). This has the eect of tending to lower the evaporator pressure and raise the condenser pressure, thus increasing the compression ratio and reducing the coecient of performance (COP), relative to that given by a comparable pure ¯uid [3]. This can be easily demonstrated for a binary mixed refrigerant, where it can be shown that the calculated COP passes through a minimum at some point in the composition range, as shown in Fig. 2 for a mixture of propane (MVC) and isobutane (LVC). The situation described above, where the LVC tends to become concentrated in the cold end of the cycle and the MVC at the hot end, is the opposite of the ideal situation, which would be to concentrate the LVC at the hot (condenser) end of the cycle, and the MVC at the cold (evaporator) end. This would tend to lower the condenser pressure and raise the evaporator pressure for given temperatures, thus reducing the compression ratio and increasing the COP. One way to achieve this would be to introduce a counter-current ¯ow between liquid and vapour in a series of contacting stages, as shown in Fig. 3. If the vapour feed stream is signi®cantly richer in MVC than the
Fig. 1. Illustration of the mechanism giving rise to a composition shift in a pool-boiling refrigeration system. (a) A single pool-boiling evaporator. Liquid from the condenser is pressure reduced to feed the evaporator, which produces a vapour stream that feeds the compressor. The vapour stream and the liquid pool in the evaporator are in equilibrium, hence the liquid pool is richer in the less volatile component and poorer in the more volatile component then the vapour. The liquid stream from the condenser and the vapour feed to the compressor have the same composition. (b) A single pool-boiling evaporator with the addition of an economiser ¯ash. The process described above is now repeated twice, so that the vapour removed from the intermediate pressure ¯ash is richer, so that the vapour removed from the intermediate pressure ¯ash is richer in MVC (and poorer in LVC) than the vapour from the evaporator. Similarly the liquid pool in the evaporator is poorer in MVC (and richer in LVC) than the liquid pool in the intermediate pressure ¯ash drum. Fig. 1. ScheÂma du meÂcanisme de changement de composition dans un systeÁme frigori®que aÁ eÂbullition libre. (a) Evaporateur aÁ eÂbullition libre simple. La pression du liquide du condenseur est reÂduite avant son passage par l'eÂvaporateur qui produit un ¯ux de vapeur alimentant le compresseur. Le ¯ux de vapeur et le liquide contenu dans l'eÂvaporateur sont en eÂquilibre ; le liquide a donc une concentration plus eÂleveÂe en composant moins volatil et de la meÃme manieÁre a une concentration moins eÂleveÂe en composant plus volatil que la vapeur. Le ¯ux liquide sortant du compresseur et le ¯ux de vapeur alimentant le compresseur ont la meÃme composition. (b) Un eÂvaporateur aÁ eÂbullition libre simple muni d'un eÂconomiseur. Le proceÂde deÂcrit ci-dessus est reÂpeÂte deux fois, de telle sorte que la vapeur enleve du ballon de deÂtente aÁ pression intermeÂdiaire est plus riche en MVC et moins riche en LVC que la vapeur en provenance de l'eÂvaporateur. De la meÃme manieÁre, le lquide dans l'eÂvaporateur a une concentration en MVC moins eÂleveÂe et une concentration en LVC plus eÂleveÂe que le liquide dans le ballon de deÂtente aÁ pression intermeÂdiaire.
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Fig. 2. Variation of calculated COP with refrigerant charge composition for a single stage pool-boiling refrigeration system, using a binary propane/isobutane refrigerant. Fig. 2. Variation du COP calcule selon la composition de la charge pour un systeÁme frigori®que monoeÂtage aÁ eÂbullition libre, pour un frigorigeÁne binaire propane/isobutane.
liquid feed stream, then this will tend to impose the desired composition and temperature pro®le on the stages. This assumes good mixing between liquid and vapour in each stage to promote ecient mass and heat transfer, ideally leading to equilibrium between the exiting streams. A novel cycle has been devised to exploit countercurrent contacting, and this article describes the cycle which can be viewed as a hybrid of traditional vapour compression and absorption refrigeration. In theory absorption refrigeration is very attractive for integration with process plant requiring a refrigeration service because expensive compressor power input can be replaced by low level heat recovered from the process. The low level heat is often viewed as free, i.e. no cost, because if not used as heat input to the absorption refrigeration regenerator, then it would be rejected to cooling water. However, the absorption refrigeration cycle is a more complex facility than vapour compression and generally requires a bigger capital investment, hence the additional capital investment must be justi®ed by the saving on compressor power. Most process plant rejects heat at a variety of temperature levels and the main concern is to match the heat released to possible users. Generally, less heat will be available at higher temperatures and this is viewed as more valuable because it can be directly recovered for process heating. This leaves the more plentiful lower temperature heat to be rejected to drive an absorption refrigeration system. It is important for the designer to note that the coecient of performance, COP (de®ned
Fig. 3. Illustration of how counter-current liquid/vapour contacting could be used to achieve the desired composition/temperature pro®le. Fig. 3. ScheÂma de l'utilisation de la mise en contact aÁ contrecourant du liquide et de la vapeur a®n d'obtenir le pro®l composition/tempeÂrature deÂsireÂ.
as the refrigeration duty, QE, divided by the heat input to the regenerator, QR), decreases when using lower level heat to drive the refrigeration unit. This means quite large heat loads can be involved, but this is not a problem because the heat can be viewed as free. The new cycle, being a hybrid, contains features of both vapour compression and absorption refrigeration, and hence appears slightly more complex. Both absorption refrigeration and the new cycle contain a regenerator/ distillation column and the recovered heat is used to drive these units. 2. Analysis In order to understand how the counter-current contacting system would operate, it is necessary to consider the behaviour of a typical single stage n, as shown in Fig. 4(a). The analysis considers a binary mixture, and it is assumed that equilibrium is achieved between liquid and vapour streams leaving the stage. Many standard textbooks deal with equilibrium stage calculations; a comprehensive treatment is given by Smith [4]. The analysis is performed on an equilibrium diagram as shown in Fig. 4(b). The equilibrium line relates molar compositions of the MVC in the liquid, x, and vapour,
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Fig. 4. Representation of a single contacting stage and associated compositions of entering and exiting streams: (a) a typical contacting stage; (b) the four stage compositions shown on an equilibrium diagram. Fig. 4. ScheÂma d'un eÂtage aÁ contact unique, avec les compositions des ¯ux entrants et sortants : (a) un eÂtage aÁ contact typique ; (b) les quatre compositions d'eÂtage sous forme de diagramme.
y, that are in equilibrium. Hence the compositions xn and yn of the two exiting streams from stage n must lie on this line, because they are assumed to be in equilibrium. Note also that the vapour is richer in MVC than the liquid, hence yn>xn. The shape of the equilibrium line in Fig. 4(b) is typical for a near ideal binary mixture (e.g. light hydrocarbons). Relative volatility, a, is a parameter commonly used to characterise the shape of the equilibrium line. Equilibrium liquid and vapour compositions, such as xn and yn, are then related by: yn
xn 1 1xn
where
Psat;MVC Psat;LVC
1 2
Psat,MVC and Psat,LVC are the vapour pressures of the pure components measured at the same temperature. For a near ideal mixture, a does not vary very much with temperature. In order to establish the desired composition pro®le, it is necessary to transfer MVC from the vapour stream into the liquid stream, and LVC from the liquid stream into the vapour stream. To achieve this, the entering vapour stream Vn1 must be signi®cantly richer in MVC than the entering liquid stream Ln 1 (i.e. yn1 > xn 1 ). Finally, to achieve the desired composition pro®le, whereby the content of MVC increases as temperature decreases (i.e. as n increases), it is required that xn > xn 1 and yn1 > yn . Hence, it can be said that yn1 > yn > xn > xn 1 . The relative positions of the four stage compositions are shown on Fig. 4(b). It will be noted from this that
211
pairs of compositions for passing streams ([xn 1 ; yn ] and [xn ; yn1 ]) lie above the equilibrium line, as opposed to the situation in distillation where compositions of passing streams are below the equilibrium line. Moving on to consider a series of contacting stages as shown in Fig. 5(a), it becomes apparent that the composition change through the stages can be represented diagrammatically in the same way that a McCabe± Thiele diagram can be used to represent composition change in a distillation column [4]. An absorption heat pump using the above concept and graphical construction has been described elsewhere by Le Go [5,6] and co-workers. A further important assumption here is that of constant molar over¯ow [4]; this is approximately true for many mixtures such as light hydrocarbons, and it means that vapour and liquid ¯ows will be constant through a series of equilibrium stages, i.e. Vn=Vn+1=V and Ln 1=Ln=L. Fig. 5(b) shows a graphical construction for ®ve equilibrium contacting stages. The compositions of passing streams, e.g. [xn 1 ; yn ] and [xn ; yn1 ], are related by the operating line. The mathematical form of the operating line is easily determined by performing a molar balance on the MVC over part of the cascade of stages. For example, considering stages 1 to n the following relationship is obtained: yn1
L xn y1 V
L xF V
3
This shows the points representing pairs of passing streams lie on a straight line with gradient L/V. A physical interpretation can now be given to Fig. 5(b). The equilibrium line is drawn for a relative volatility of 3; this corresponds approximately to a mixture of propane (MVC) and isobutane (LVC) over a temperature range from 20 to 20 C. Both feeds to the series of contacting stages have been taken as 98% pure; the top (saturated liquid) feed stream contains 2 mol% MVC and the bottom (saturated vapour) feed stream contains 98 mol% MVC. The gradient of the operating line is 1, i.e. equal liquid and vapour ¯ows are to be used through the ®ve stages. Table 1 shows the vapour and liquid compositions for each stage, taken from Fig. 5(b). This illustrates more clearly how the aim of establishing a composition pro®le has been achieved by transferring the MVC from the vapour to the liquid and the LVC from the liquid to the vapour. The MVC rich vapour feed has thereby been converted to a vapour product that is poorer in MVC and richer in LVC, and is thus more easily condensed, i.e. for a given condensing temperature a lower pressure will be required, which was the initial aim. The greater the composition change that can be achieved, the greater this bene®t will be.
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Fig. 5. Representation of a series of contacting stages and associated compositions. (a) A series of N equilibrium stages. Stage n is a typical equilibrium stage. The liquid stream becomes richer in MVC as it ¯ows down through the stages, and the vapour stream becomes depleted in MVC as it ¯ows up. (b) A graphical construction for a series of ®ve equilibrium stages, drawn for a binary mixture with a relative volatility of 3. Fig. 5. ScheÂma d'une seÂrie d'eÂtages aÁ contact avec les compositions associeÂes. (a) Une seÂrie d'eÂtages d'eÂquilibre N. L'eÂtage n est un eÂtage d'eÂquilibre typique. Le ¯ux liquide en eÂcoulement descendant devient plus riche en MVC au fur et aÁ mesure qu'il coule aÁ travers les eÂtages, et le ¯ux de vapeur a une concentration de moins en moins eÂleveÂe en MVR au fur et aÁ mesure qu'il poursuit son ascension. (b) ScheÂma graphique d'une seÂrie de cinq eÂtages en eÂquilibre pour un meÂlange binaire avec une volatilite relative de 3.
Table 1 Comparison between composition pro®les from graphical construction and computer simulation, and summary of stage temperatures from the simulation Tableau 1 Comparaison entre les pro®ls de composition graphiques et la simulation informatique. Les tempeÂratures des eÂtages sont reÂsumeÂes aÁ partir de la simulation Stage
Top feed 1 2 3 4 5 Bottom feed a
Compositions (mol% MVC) Graphical construction
Computer simulation
x
y
x
y
2 17 27 37 47 62 ±
± 38 53 63 73 83 98
2 17 28 38 48 63 ±
± 35 51 63 73 83 98
Assumed saturated.
Temperature ( C) (from computer simulation)
16.1a 8.7 4.0 0.3 3.3 7.8 15.4a
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Also shown in Table 1 are the results (stage compositions and temperatures) of a computer simulation of the propane/isobutane system, produced using Hyprotech's HYSYS package. This shows two important results. Firstly, it con®rms that the compositions are very close to those predicted by the graphical construction in Fig. 5(b). Secondly, at an operating pressure of 2.75 bar, a temperature lift of 16.5 C was obtained between top and bottom stages (with a bottom stage temperature of 7.8 C and a top stage temperature of 8.7 C). Whilst this temperature lift is insucient to be of practical use, it demonstrates the principle that with a mixed refrigerant, composition variation alone (no pressure change is considered at this stage) can provide a temperature lift using a series of counter-current contacting stages. 3. Practical considerations So far, the use of countercurrent vapour/liquid contacting through a series of equilibrium stages in order to establish a composition/temperature pro®le has been demonstrated. Practical implementation of the scheme has not been considered. In common with other stagewise equilibrium processes, such as distillation, this process could most easily be carried out in a packed or plate column. The individual contacting stages would therefore be implemented as either packing or plates within the column designed to promote ecient mass/heat transfer. For the column to be of practical application in a refrigeration cycle, a number of other factors must also be taken into account. Firstly, the provision of evaporation and condensation units has not yet been considered. Secondly a greater temperature lift is needed. For example, a typical refrigeration duty might require an evaporator temperature of 25 C, and the cooling water available would determine the minimum condenser temperature, typically 40 C. This implies a temperature lift of 65 C. Finally, in order to generate the column composition pro®le, it is necessary to provide relatively pure feed streams to the top and bottom of the refrigeration column. Therefore, a refrigerating column would need to form part of a larger system, which would include a separation operation to provide the top and bottom feed streams. 3.1. Provision of evaporator and condenser In order to maximise the temperature lift, the evaporator and condenser must be positioned as shown in Fig. 6. This allows us to claim an increased temperature lift of 24.1 C, from 15.4 C (saturated vapour leaving the evaporator and entering the bottom stage) to 8.7 C (saturated vapour leaving the top stage). However, the bubble point temperature of the exiting vapour stream is only 1.2 C at the column pressure, which is the tempera-
Fig. 6. Illustration of the arrangement for the column and auxiliaries. Fig. 6. ScheÂma de la con®guration de la colonne et des composants annexes.
ture to which it would have to be cooled to complete condensation. This is obviously too low for condensation against cooling water or air, so a greater temperature lift is still needed. 3.2. Increasing the temperature lift The two most obvious ways to achieve this are either to introduce a pressure increase between evaporator and condenser, as in a normal vapour compression refrigeration cycle, or to extend the temperature lift resulting from the composition pro®le. 3.2.1. Pressure Raising the pressure between evaporator and condenser might be performed in a number of ways, depending on where the pressure change was introduced, as shown in Fig. 7. Considering a system with two pressure levels, three variations are possible: ®rstly to split the column into two parts operating at dierent pressures; secondly to introduce the pressure variation between the column and the condenser; or thirdly to put it between the column and the evaporator. For practical application, the con®guration of Fig. 7(b) Ð compressing the overhead vapour from the refrigerating column before feeding it to the condenser Ð is considered best. The other two con®gurations shown in Figs. 7(a) and (c) are less suitable because they both
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Fig. 7. Possible arrangements for introducing a pressure change between the evaporator and condenser of the refrigeration column: (a) pressure change within column; (b) pressure change between column and condenser; (c) pressure change between column and evaporator. Fig. 7. Con®gurations possibles a®n d'obtenir un changement de pression entre l'eÂvaporateur et le condenseur de la colonne frigori®que : (a) changement de pression aÁ l'inteÂrieur de la colonne ; (b) changement de pression entre la colonne et l'eÂvaporateur.
place the compression heat within the column, increasing the vapour ¯ow and thus requiring a larger column. 3.2.2. Composition To increase the temperature lift resulting from composition change would require maximising the composition change between evaporator and condenser, and/or the use of a mixture with a wider dierence in volatility between the MVC and the LVC. Considering the propane/isobutane system at 2.75 bar studied previously, the boiling points of the pure components are 16.6 and
17.2 C respectively, so even if the change in composition could be fully exploited, the temperature lift would be insucient (33.8 C). Using a mixture with a greater dierence in volatility is not a solution either. If the equilibrium line is constructed for a=5, then the available working region above the equilibrium line is much smaller and this in turn restricts the usable composition range, thus defeating the object of using a pair of components with a higher relative volatility. Hence the pressure change alternative is pursued as the best means to increase the temperature lift.
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Fig. 8. One possible con®guration for incorporating the refrigeration column into system with a distillation column. Fig. 8. Con®guration possible permettant d'inteÂgrer la colonne frigori®que dans un systeÁme aÁ colonne de distillation.
3.3. Completing the cycle In order to supply the relatively high purity feed streams needed to operate the refrigeration column, it would be necessary to run it in conjunction with a separation operation. A distillation column is an eective way to perform the separation needed and hence becomes an integral part of the cycle. The complete cycle is illustrated in Fig. 8. The distillation column top product (MVC rich) would form the feed to the evaporator, where it is completely vaporised before passing to the bottom of the refrigeration column. The distillation column bottom product (LVC rich) would feed the top of the refrigeration column. To develop a design the assumption of 98 mol% propane in the distillation column top product and 98 mol% isobutane in the bottoms will be maintained. The cycle is completed by returning the refrigeration column overhead vapour and bottoms liquid to appropriate locations in the distillation column. An important parameter in designing a distillation column is the re¯ux ratio. This de®nes the ratio in which the liquid from the condenser is divided for return to the column as re¯ux and taken as top product. In theory there is a minimum re¯ux ratio which leads to the requirement for a very large number of contacting stages in the distillation column. For a practical design, a commonly used heuristic suggests a re¯ux ratio of around 1.25 times the minimum re¯ux ratio [7]. It is also
important to introduce the feed(s) onto the correct stage(s) of the distillation column. In the design outlined below, and the Sensitivity Study in the next section, the 1.25 factor has been used and the feed location optimised to put all the designs on a consistent basis. The highest pressure needed in the cycle will be that of the distillation unit, which will be determined by the minimum temperature (say 40 C) obtainable with cooling water. The pressure is ®xed by the requirement to completely condense the overhead vapour, i.e. a propane/isobutane mixture containing 98 mol% propane has a bubble-point pressure of 13.5 bar at 40 C. Similarly, the pressure of the refrigeration evaporator and column is ®xed by the requirement for the evaporator to produce a saturated vapour at 25 C, at which temperature the dew-point pressure for the above mixture is 1.9 bar. The distillation column product streams therefore have to be let down from 13.5 to 1.9 bar. If these streams are not subcooled before being let down, approximately 40% of both streams ¯ashes to vapour, which is a problem since it reduces the refrigerating capacity of the evaporator whilst increasing the load on the compressor and condenser. The bottom product stream from the distillation column will be hot enough to be cooled against cooling water (to 40 C), in the heat exchanger labelled ``pre-cooler (water)'' in Fig. 8, before being let down, thus reducing the ¯ash to around 25%. However, since the top product stream is already at 40 C, it can only be further cooled using refrigeration
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recovery, in the heat exchanger labelled ``pre-cooler (refrigerant)''. The ¯ash for this stream is thus reduced to approximately 15%. The return of the top vapour and bottom liquid streams from the refrigeration column to the distillation column requires power input. The bottom liquid is easily pumped from 1.9 up to 13.5 bar, but the top vapour need not be compressed beyond its bubble-point pressure at 40 C, in this case 8.2 bar. This represents a signi®cant saving on compressor power. Once condensed, the liquid stream is easily pumped to 13.5 bar. The performance of the cycle discussed above and illustrated in Fig. 8 is summarised in Table 2. It is worth noting that the highest temperature in the cycle, to be found at the bottom of the distillation column, is 79 C. This is approximately 100 C cooler than the distillation temperature for a simple ammonia-water absorption refrigeration cycle operating under similar conditions. Included in Table 2 are two new terms, derived from coecient of performance. Power ratio is calculated as the refrigeration heat load divided by the compressor power, and heat ratio as the refrigeration heat load divided by the distillation column reboiler duty. These two terms are useful for later comparison purposes.
4. Sensitivity study The idea of introducing the refrigeration column with two volatile working ¯uids is new so it is worth examining the sensitivity of the design to changes in the following important parameters:
Table 2 Performance of the refrigeration cycle as illustrated in Fig. 8 Tableau 2 Performance du cycle frigori®que illustre en Figure 8 Refrigerant evaporator
Temperature ( C) Pressure (bar) Heat load (kW)
25 1.9 1000
Refrigerant condenser
Bubble-point temperature ( C) Pressure (bar) Heat load (kW)
40 8.2 1373
Work input
Compressor (kW) Pump 1 (kW) Pump 2 (kW) Total (kW)
328 7 6 341
Refrigeration Number of stages column Pressure (bar) Bottom stage temperature ( C) Top stage temperature ( C)
5 1.9 19 2
Distillation column
32 13.5 79 40 2848 2410
Number of stages Pressure (bar) Reboiler temperature ( C) Condenser bubble-point temperature ( C) Reboiler heat load (kW) Condenser heat load (kW)
Power ratio
2.93
Heat ratio
0.35
Fig. 9. Eect of varying the number of stages in the refrigeration column, keeping the composition of the feed streams ®xed (98% pure), on: (a) heat and mechanical work input; (b) composition of streams exiting refrigeration column. Fig. 9. Eet de la variation du nombre d'eÂtages dans la colonne frigori®que, en maintenant une concentration des ¯ux entrants stable (98 %) sur : l'eÂnergie thermique et le travail meÂcanique ; (b) la composition des ¯ux sortant de la colonne frigori®que.
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. Number of contact stages in the refrigeration column, whilst maintaining both feed streams at 98 mol% purity. . Purity of the feed streams to the refrigeration column whilst maintaining ®ve contacting stages in the column.
Fig. 9(a) shows that as the number of stages in the refrigeration column is increased, both heat input to the distillation reboiler and work input to the compressor decrease. This is to be expected, because with a greater number of stages in the refrigeration column, the irreversibilities are reduced, enabling a greater change in composition to be obtained through the column, as shown in Fig. 9(b). The overhead vapour from the column will become richer in LVC, so its bubble-point pressure will fall, which reduces the work input to the compressor. Similarly, the bottom liquid stream will become richer in MVC. With higher purity feed streams being returned to the distillation column, the separation will become easier, reducing the reboiler heat load. The number of stages in the distillation column does not change signi®cantly, but the re¯ux ratio falls slightly as the number of stages in the refrigeration column is increased. With reduced vapour trac through the distillation column, a smaller diameter column will be needed. Therefore, increased capital expenditure on the refrigeration column is accompanied by a reduction in utility costs and also a reduction in capital expenditure on the distillation column. Fig. 10 shows the results of changing the purity of the feed streams to the refrigeration column, the number of stages in the column being ®xed at 5. For simplicity, the composition of both feed streams has been changed in parallel. For example, 90% purity means that the bottom feed stream contains 90 mol% MVC and 10 mol% LVC, and the top feed stream contains 90 mol% LVC and 10 mol% MVC. Fig. 10(a) shows that as the feed purity is decreased from its initial value of 98%, the compressor work input rises and the reboiler heat load falls. With less pure feed streams, the available composition change in the refrigeration column is reduced, so the overhead vapour stream becomes poorer in LVC, increasing its condensing pressure and hence the compressor work input. Lower purity feed streams means a less rigorous separation is required, so less heat input is needed to the distillation column reboiler. This illustrates a trade-o between heat and work input. Fig. 10(b) shows the eect on the distillation column of changing the purity of its product streams. As the purity is reduced, the number of contacting stages falls, as does the reboiler temperature. With a ®xed condenser temperature, the column pressure also decreases as the overhead product stream becomes poorer in MVC. Therefore, reducing the purity of the feed streams to the refrigeration column also implies a reduction in capital expenditure on the distillation column.
Fig. 10. Eect of varying the purity of the feed streams to the refrigeration column, with ®ve stages, on: (a) heat and mechanical work input; (b) number of stages in the distillation column, and distillation column reboiler temperature. Fig. 10. Eet des variations de concentration des ¯ux entrant dans la colonne frigori®que, avec cinq eÂtages, sur : l'eÂnergie thermique et le travail meÂcanique ; (b) le nombre d'eÂtages dans la colonne de distillation, et la tempeÂrature du rebouilleur de la colonne de distillation.
5. Comparison to standard refrigeration cycles The proposed cycle is a hybrid so it is interesting to make a comparison with absorption and vapour compression refrigeration. The new cycle is closest to absorption refrigeration, so this is considered ®rst. The comparison
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needs to be made under two headings, namely complexity (as an indication of capital investment) and utilities consumption (as an indication of operating costs). The ammonia±water cycle has a refrigerant operating range similar to the new cycle and has been extensively investigated, hence it forms a good basis for comparison. Ammonia absorption refrigeration cycles can vary widely in complexity [8,9]. To start the comparison, the simplest absorption refrigeration unit is considered, i.e. single stages for evaporation, absorption and regeneration. Table 3 makes a comparison on the number of items in the cycles and their direct equivalence. As may be noted each item in an absorption refrigeration cycle has a direct equivalent in the new cycle with the exception of the compressor, which is an additional item. In order to understand why the new cycle is of interest, despite the addition of a compressor, it is necessary to compare the operating capabilities of absorption refrigeration to those of the new cycle.
Table 3 Equipment comparison between ammonia absorption refrigeration cycle and new cycle Tableau 3 Comparaison du cycle frigori®que aÁ absorption aÁ ammoniac et le nouveau cycle Simple ammonia absorption New cycle refrigeration cycle Regenerator (includes reboiler and condenser) Absorber±condenser Evaporator ± One pump Two heat exchangers (Two interchangers)
Distillation column (includes reboiler and condenser) Refrigeration column condenser (refrigerant) Evaporator Compressor Two pumps Two heat exchangers (interchanger, cooler)
Table 4 Comparison of absorption refrigeration with new cycle Tableau 4 Comparaison du cycle frigori®que aÁ absorption et le nouveau cycle Boiling refrigerant at ( C)
Description
Heat recoverable to ( C)
Single stage New cycle
150 85
25 25
Dual pressure regeneration A Dual pressure regeneration B
106
25
85
+2
COP or power/ heat ratio COP 0.4 Power ratio 2.93 Heat ratio 0.35 COP 0.26 COP 0.34
Briley [10] provides a useful table detailing the main operating characteristics for some alternative ammonia absorption refrigeration designs. Table 4 is derived from this work and is intended to summarise alternative designs which can be compared with the new cycle. As may be seen from Table 4, to provide boiling refrigerant at 25 C, a simple ammonia absorption refrigeration unit would need heat at 150 C or greater, and operate with a COP of approximately 0.4. These numbers are broadly con®rmed by Holldorf [9] and Borsig [11], and can be compared with the requirements of the new cycle, namely heat available at or above 85 C and a heat ratio of 0.35. However, a compressor and its power input is also needed; this might be viewed as the penalty paid to permit heat recovery to such low temperatures. The ability to recover low level heat is an important feature of the new cycle, so consideration should be given to the changes which are necessary to permit an ammonia absorption unit to operate on colder heat supplies. Bogart [8] and Holldorf [9] describe two approaches: . multiple reboilers on the regenerator, each reboiler located higher up the column and operating at successively lower temperatures; . dividing the regenerator into two stages, with the water rich section operating at a lower pressure and hence capable of accepting heat at a lower temperature.
Both alternatives add to the complexity and hence cost of the unit and in particular the dierence between the two stage regeneration and the new cycle is not clear cut. Two stage regeneration is the most frequently quoted route for reducing the required temperature of the heat source and useful data is given by Briley [10]. This has been used to generate two alternative designs given in Table 4. These alternatives have been chosen to highlight performance dierences compared to the new cycle. In Table 4, dual pressure A shows that the minimum heat source temperature feasible to produce 25 C refrigerant is 106 C and the COP has dropped signi®cantly to 0.26. To highlight the importance of heat source temperature, dual pressure B shows that taking heat down to 85 C would permit a refrigeration temperature of 2 C and a COP of 0.34. Clearly the new cycle is able to operate at more extreme conditions but the penalty paid is the addition of the compressor. It is probably worth observing that a refrigeration temperature of 25 C is at the cold extreme for absorption refrigeration. Absorption refrigeration is much more eective for refrigeration temperatures around 0 C and one would anticipate similar performance improvement for the new cycle. Obviously, the refrigeration column pressure would be higher leading to less ¯ash on liquid pressure reduction and reduced compressor power consumption. Finally, to complete the comparison vapour compression refrigeration needs to be considered. Here pure
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propane is considered and the COP for a cycle operating with an evaporator at 2.03 bar and condenser at 13.7 bar is 1.88. This value should be compared with the power ratio of 2.93 for the new cycle. 6. Comparison with previous work As noted previously Le Go and co-workers [5,6,12,13] have studied absorption heat pumps using ``reverse recti®cation''. However, the binary mixtures considered (e.g. water/ethylene glycol, ethylamine/water and pentane/octane) are much further apart in terms of volatility than our example. In such cases a straight operating line, as used in this work, would permit only a very restricted composition change. Le Go and coworkers create a change in operating line gradient by a combination of heat removal and withdrawing liquid from the middle section of the column. The heat removed is the up-graded energy from the heat pump but this is obviously not feasible for the refrigeration column discussed above because the entire refrigeration column operates below ambient temperature. Heat pumps combining absorption and compression are investigated by a number of authors. The systems considered all seem to be based on standard absorption cycles using either ammonia/water or water/lithiumbromide refrigerant/absorbent pairs, with vapour compression added at some point in the cycle to improve either eciency or ¯exibility. The development of these systems is reviewed by Morawetz [14] and Ahlby and Hodgett [15]. One variant of these combined cycles is analysed by Ziegler and Spindler [16], who look at optimising the performance of a single stage vapour compression refrigeration cycle by adding an economiser ¯ash. The use of a second compression stage to deal with the intermediate pressure ¯ash vapour is compared with the use of an absorption circuit. The compression plus absorption con®guration is shown to be advantageous because the heat input to the absorption circuit can be supplied solely by the superheat from the compressor discharge. Experimental results are presented for the compression plus absorption cycle. Aly [17] looks at a water/ lithiumbromide absorption cycle where vapour recompression is used in place of the conventional condenser. This system is aimed at high temperature use, not refrigeration. 7. Conclusions The use of composition change with a mixed refrigerant to achieve a temperature lift has been examined. It has been demonstrated theoretically how this may be achieved by using a series of equilibrium stages with countercurrent vapour/liquid contacting. These stages could be arranged in column format. A simple graphical
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analysis of the system along the lines of a McCabe±Thiele construction for distillation has been presented for the case of a binary mixture. The possibility of integrating the column into a closed cycle has been examined, along with the practical implications of this in terms of energy requirements. Performance is shown to be comparable to conventional absorption refrigeration units bearing in mind the possibility of recovering much lower level heat. The designer has considerable ¯exibility over the use of shaft power and recovered low level heat. Like all absorption refrigeration based systems, the new cycle is best suited to situations where waste heat at low temperatures is readily available.
References [1] Didion DA, Bivens DB. Role of refrigerant mixtures as alternatives to CFCs. Int J Refrig 1990;13:163±75. [2] Bensa® A, Haselden GG. Wide-boiling refrigerant mixtures for energy saving. Int J Refrig 1994;17(7):469±74. [3] Swinney J, Jones WE, Wilson JA. The impact of mixed non-azeotropic working ¯uids on refrigeration system performance. Int J Refrig 1998;21(8):607±16. [4] Smith BD. Design of equilibrium stage processes. McGraw-Hill, 1963 (chapter 5). [5] Ranger PM, Matsuda H, Le Go P. Modelling of a new type of absorption heat pump combining recti®cation and ``reverse-recti®cation''. J Chem Eng Japan 1990;23(5):530±6. [6] Ranger PM, Matsuda H, Le Go P. Modelling and experimental operation of a ``reverse-recti®cation column'' for an absorption heat pump. J Chem Eng Japan 1990;23(5):537±42. [7] Yaws CL, Li KY, Fang CS. How to ®nd the minimum re¯ux for binary systems in multiple-feed columns. Chem Eng 1981;88(10):153±6. [8] Bogart M. Ammonia absorption refrigeration in industrial processes. Houston (TX): Gulf Publishing Company, 1981. (chapter 17) [9] Holldorf G. Revisions up absorption refrigeration eciency. Hydrocarbon Processing. July 1979;58:149±155. [10] Briley G.C. Conserve energy. . . refrigerate with waste heat. Hydrocarbon Processing. May 1976;55:173±174. [11] Anon. Borsig pocket book. 5th ed. p. 336 [12] Le Go P., Labidi J. The concept of `reverse-fractionaldistillation' exergy analysis and applications. IChemE Symposium Series 1992; 128 B55-B67. [13] Labidi J, Schwarzer BP, Le Go P. Absorption heat pumps composed of multiple stages either independent or belonging to a unique column. Int Absorption Heat Pump Conf, ASME 1993;31:251±6. [14] Morawetz E. Sorption-compression heat pumps. Int J Energy Res 1989;13(1):83±102. [15] Ahlby L, Hodgett DL. Compression±absorption heat pumps. In: 3rd Int. Energy Agency Heat Pump Conf. Oxford: Pergamon Press, 1990. [16] Ziegler F, Spindler U. An ammonia refrigerator with an absorption circuit as economizer. Int J Refrig 1993;16(4):230±9. [17] Aly SE. Energy-ecient absorption/mechanical vapourrecompression system. J Inst Energy 1991;64(458):36±40.