The formula related to the calculation of the system efficiency is described in eq. (1). All the system’s outputs are divided by the same magnitude which is the energy content of the initial fuel input on a Lower Heating Value (LHV) basis. In this way it is possible to identify how much energy contained in the MSW is transformed into electricity, heating and cooling and how much are the system losses. It is important to mention, that the formula below is introduced in a simplified manner and the reason is to estimate the overall performance of the plant. An exergy based approach is required in order to identify the location and the real magnitude of inefficiencies occurring in the plant, apart from the energy losses to the environment. However, an exergy analysis is out of the scope of the current work:

ṖNet is the net electrical power of the system and is calculated in eq. (2) by taking into account the gross electrical power ṖFuelcell and the power consumption of the four compressors. The power consumption of the pump is very small (< 2 kW), consequently it is neglected. The terms Q̇Absorption and Q̇Heatex imply the cooling effect of the absorption cycle and the heat duty of heat exchanger (HEATEX6) assigned for the hot water production respectively:

In order to determine the energy value of the waste delivered to a waste to energy plant, the energy content has to be calculated. There are two methods for determining the energy content:

• The Higher Heating Value (HHV);

• The Lower Heating Value (LHV).

The HHV relates to the complete combustion of a fuel from an initial state of 25 °C and the return of the combustion products to the initial state. In this way the latent heat produced from the vaporization of the water contained in the fuel, is considered in the overall heating value. On the contrary, LHV refers to the complete combustion of a fuel from an initial state of 25 °C and the return of the combustion products in the temperature of 150 °C by cooling. The heat of water vaporization in this case is not taken into account. The calculation of these two values can be performed either by knowing the chemical analysis found by conducting experiments, either in the laboratory by using a bomb calorimeter. In the case of known chemical analysis, well known equations are introduced in order to calculate HHV and LHV respectively [16], [17]. The elements in eq. (3) represent percentage by weight (ww%) in the fuel on a dry basis:

The building of the complete CHCP model was developed in Aspen Plus simulator. Aspen Plus has been widely used for modelling complex steady-state processes in the fields of chemical and petrochemical engineering. It includes a large library of blocks such as compressors, heat exchangers and reactors that represent most operations in industrial processes. Furthermore, it constitutes of a large property databank containing information about thermo physical properties of different chemical substances and mixtures. The calculations related to the chemical reactors are performed based on the assumption of minimization of Gibbs free energy. It is assumed that the gas streams in the gasification and fuel cell sections have an ideal gas behavior (IDEAL property method). SOLID method was implemented regarding the modelling of sorbents in gas cleaning reactors. The fluid property package developed by REFPROP was used to model the absorption cycle system, as the working fluids are highly non-ideal mixtures for which reference equations of state are preferred over conventional ones.

Determination of the given input parameters to the system is based on the following assumptions. The gasifier is of the downdraft type, in isothermal, steady-state conditions and chemical reactions reach their equilibrium [18]-[21]. MSW at the ambient conditions (T = 15 °C and p = 1.013 bar), with a rate of ṁ = 130 kg/h after drying and moisture content 10% [22] enters a downdraft gasifier. The gasifier, where the decomposition of the fuel takes place followed by the combustion and gasification processes, operates at 800 °C and in atmospheric pressure. Ambient air is first compressed and then preheated before it is supplied into the gasifier. The produced syngas is then processed through the cleaning system, where Hydrogen chloride (HCl) and Hydrogen sulfide (H₂S) are reduced to the desired ppm levels.

The physical and chemical composition of MSW is derived from [23] and consists of 52% plastics, 10% textiles, 22% food waste and 16% paper. The volatile matter content was taken as an average value of the data found in [24] due to the lack of information. The author of [23] studied the chemical as well the physical composition of the waste produced in the National Hospital of Taiwan, thus it is considered representative for modelling hospital solid wastes. The chemical composition of waste is displayed in Table 1.

Syngas produced through gasification contains different kinds of impurities such as particles, tars, sulfur, chlorine compounds and alkali metals, which can poison the catalysts used in the fuel cells. Pretreating the syngas is crucial to ensure high system performance and life expectancy [25]. The cleaning methods can be divided into two types, depending on their operating range: cold (below 200 °C) and hot (possibly above 500 °C) gas cleaning [25]. Cold gas cleaning requires additional heating and cooling compared to hot gas cleaning, which results in an efficiency penalty and can generate waste streams that need further treatment. Hot gas cleaning may be more energy-efficient but is not a mature technology. Gas cleaning systems related to syngas produced from waste gasification are in general terms much simpler when compared to coal gasification. The feedstock supplied to the gasifier does not contain large particles, since they have been usually removed through recycling process, before the feedstock is fed to the gasifier. Alkali metals, such as potassium and sodium, are not present in the given MSW composition and consequently only the removal of HCl and H₂S compounds is taken into consideration. HCl can be removed in an absorption process by using calcium or sodium based sorbents, whereas metal oxides such as Zinc (Zn) and Molybdenum (Mo) can be used for the H₂S removal. Other particles can be removed in a high-temperature filter prior to HCl and H₂S catalysts. In this study the use of the sorbent Sodium bicarbonate (NaHCO₃) for HCl removal to few ppm level is determined, since it has proved to be very effective (99% of HCl removal) within a temperature range of 350-500 °C [26]. Similarly, H₂S, which is the most common sulfur-derived compound in the synthesis gas [26], is removed by absorption with commercially available sorbent Zinc oxide (ZnO) in temperatures between 350-450 °C [27]-[30].

Fuel cells produce electricity directly from the fuel oxidation process, which can result in higher system performances. The electrical efficiencies range between 40 and 55%, while the fuel utilization factor can reach up to 80% in CHP mode [17]. Meanwhile, the efficiency of internal combustion engines is approximately 30% and therefore, the introduction of fuel cell technologies in the power generation system can be highly beneficial. Fuel cells are characterized by low noise due to the lack of moving parts and very low emissions compared to other conventional technologies [5].

The calculation of the SOFC electrical output is based on the work of Zhang et al. [31]. The challenge of creating a SOFC model in Aspen Plus software lies on the fact that there is no model available in the software’s component library. Thus, the author introduced an alternative method for the design of a tubular SOFC, by utilizing the existed building blocks of Aspen Plus. Afterwards, it was validated by performing a comparison to a Siemens-Westinghouse 100 kW class tubular SOFC stack.

The main input data is presented in Table 2. The electrical output results from an energy balance conducted in the anode reactor, taking into consideration the heat produced from the electrochemical reactions, the amount of air needed to keep the operating temperature of the cell stable, as well as an assumed percentage of 2% of heat losses. The SOFC operates in isothermal (780 °C) steady-state atmospheric conditions. Syngas is compressed and preheated to 650 °C before entering the fuel cell on the anode side. The inlet gas composition (mol%) is: H₂ 34%, CO 20%, CH₄ 0.04 %, Nitrogen (N₂) 27.14%, H₂O 11.3%, CO₂ 7.47%. Similarly, the ambient air entering the SOFC is first compressed and preheated to 600 °C. The flow leaving the SOFC on the anode side (off-fuel) still contains some unburnt fuel (depending on the fuel utilization factor U_F), which is burned in a catalytic burner where the depleted air (off-air from cathode), is also processed. The value of U_F, set to 0.85, is taken from [31], [32]. The assumptions of the SOFC temperature inlets have been reported in [17]. Waste heat from the exhaust gases is used to drive the absorption plant as well as the hot water production. The off-gases after the burner are cooled to 150 °C in order to provide heat for the absorption cycle, while they are cooled to 100 °C for hot water production, thus avoiding acid gases condensation. The desired domestic hot water temperature is 55 °C and the isentropic efficiency of the compressors is 70%.

Refrigeration systems based on vapour absorption cycles are a well-known technology, but their market share is still limited compared to the vapour compression systems, because of their low efficiency and high capital costs. Typical values of the Coefficient of Performance (COP) for absorption cycles range between 0.5 and 1.5, while they exceed 3 in the case of vapour compression cycles [33]. Albeit their disadvantages, the utilization of absorption cycles is significantly favoured when waste heat is available, especially if hot exhaust gases resulting from industrial processes are directly discharged in the surroundings. The integration of absorption chillers which utilize this heat that otherwise would be wasted, can lead to an increase in the overall plant efficiency.

An ammonia-water rich solution consisting of 60% H₂O and 40% Ammonia (NH₃) on a molar basis, is pumped to the desorber after internal heat recovery with the hot weak solution (poor concentration of ammonia). Heat from the burner exhausts is introduced to the desorber resulting in partial evaporation of ammonia and water. In the outlet of the desorber the mixture generally contains about 5-10% water on a mass basis, thus further decrease in the water concentration is needed. This is achieved through cooling and condensation in the rectifier. As a result, almost pure ammonia mixture is led to the condenser where it is transformed into a liquid solution. A Vapor-Liquid Heat Exchanger (VLHEATEX) is implemented afterwards in order to sub cool the mixture, which then enters in an expansion valve where its pressure is decreased to the low pressure level of the cycle. The weak ammonia-water solution exiting the desorber is mixed with the water removed in the rectifier. It is then expanded and mixed with ammonia exiting the vapor-liquid heat exchanger in the absorber, thus closing the cycle. The rectifier temperature, as shown in Table 3, is set to 95 °C in order to achieve an ammonia purity of 98% in the gas phase at the outlet. The mass flow of the solution is adjusted to achieve the desired temperature of 105 °C in the desorber, based on the waste heat available of 266 kW. The evaporator and condenser temperatures are deduced from the constraint related to the assumed minimum temperature difference of 5 K in the heat exchangers.

The assumptions needed for a realistic modelling of the absorption cycle are as follows:

• Saturated liquid conditions in the outlet of the absorber and the condenser;

• The cooling medium in the evaporator is air which is cooled from 25 °C to 22.5 °C, based on the environmental conditions for thermal comfort [34];

• Water in initial temperature of 15 °C is used as a cooling medium in the absorber, rectifier and condenser.

A sensitivity analysis needs to be performed in order to identify the impact of varying parameters on the system efficiency. The parameters to be varied are the Air Equivalent Ratio (AER) and the operating temperatures of the gasifier and desorber respectively.

Apart from the thermodynamic analysis of the studied system, a techno-economic analysis is also performed to investigate the cost effectiveness of the plant. The data related to the Purchased Equipment Cost (PEC) stems from [35]-[38]. The unit cost is in (USD_2015). The cost estimation of SOFC, which is not yet in the market, builds on the following considerations. The cost of SOFC is calculated from eq. (14) based on the study of [40] which was derived for serial cell production in a future scenario. It should be noted that such hypothesis may change in the future and the authors use this relation since no other relation can be found in the open literature. At present, the cost is four times larger due to non-commercialization. Regarding the purchased cost of the gasifier, a number of fixed-bed applications have been recorded around the world and cost estimates are available in the literature. Absorption technology is a mature technology and the related cost data is readily available as well.

In eqs. (13-16) the cost correlations related to gasifier, SOFC, absorption chiller and counter flow heat exchanger respectively, are presented. In eq. (13) ṁMSW refers to the mass flow (kg/sec) of dried MSW entering the gasifier reactor. The terms N_Cells and D_Cells in eq. (14) relate to the number of cells within a stack as well the diameter of each cell in (m). The cost factor F_AC in eq. (15) depends on the heat source as well as from the chiller’s single or double effect. For a single effect absorption chiller and water as the heat source which is the case of the specific study, the cost factor is determined to be 1. The term A_Heatex refers to the heat transfer area, which is calculated in eq. (17). Q is the heat exchanger duty in (W), LMTD is the mean logarithmic temperature difference in every heat exchanger and U (W/m²K) is the overall heat transfer coefficient. The values of U depend on the working fluids and the considered values are taken from literature. Finally the correlations related to other components of the plant such as compressors, pumps are well known, thus there is no need of mentioning them and can be found in the references stated above:

where:

The costing method, for the calculation of the Total Capital Investment (TCI) is described thoroughly in [39] and presented in Figure 2. The TCI is the sum of Fixed Capital Investment (FCI) and other outlays which are referring to the start-up costs and the licensing of a project as well as to the working capital. Working capital relates not only to the payment of salaries but also to the expenses associated with materials, fuels and different kinds of prerequisites regarding the operation. FCI is the sum of Total Direct Costs (TDC) and Total Indirect Costs (TIC), respectively. TDC are further divided in onsite and offsite costs. The former costs are associated with the Purchased cost of Each Component (PEC), the installation costs and the equipment needed for system controlling. The latter costs refer to service facilities such as water and electricity utilities and the civil and architectural work. Finally TIC relates to the oversight of the system, the arrangement of any uncertainties or risks that will arise during the project and the costs that are destined for the contractor’s payment.

Figure 2.

Overview of TCI calculation

The calculation of the Annual Investment (ACI) is essential to understand the different capital costs involved in the installation and maintenance of the plant. First, the operating time of the plant is considered to be 20 years. The maintenance factor (f_Maintenance) refers to the deterioration of the equipment along the operating years and the consequent adding costs that arise for service and maintenance. The chosen value is 1.1% and has been retrieved from [37]. The factor of discount rate (f_Discount) has to do with the costs that stem from the fact that the money for the construction of the plant shall be borrowed. Even if the money is supplied from internal sources, there is still an opportunity cost, since the money could be utilized elsewhere or even deposited in to a bank [40]. The value of 5% is selected from [41] and refers to a socio-economic perspective. The inflation factor, which is the price level changing rate over the years, is taken to a typical value of 2%. Finally, the last two factors, f_Insurance and f_Taxation are related to the impacts of taxation and insurance elements on the annual investment. The corresponding values of 0.2% and 0.54% have been reported in [40].

Eqs. (18-22) describe the formulas for the calculation of different costs based on the factors explained above:

The annual capital investment (C_ACI) is simply the sum of the different costs as presented in eqs. (8-16):

The operating temperature inside the gasifier is a very crucial parameter regarding the gasification process since it greatly affects the syngas composition. In Figure 3 the trends of the different chemical elements are illustrated while varying the gasifier temperature in the range of 650-950 °C. It can be observed that the H₂ content increases slightly from 650 °C to 700 °C, where it starts to decline. CO shows a significant increase from 22.5% to 32.5% (40% increase), while CO₂ follows the opposite trend with a 45% decrease (from 22% to 12%). The CH₄ content decreases with an increasing temperature, while the water fraction rises from 700 °C and forward (24% increase). The diverse trends in the syngas composition can be explained by the thermochemical reactions occurring inside the gasifier. The increase in the gasification temperature favors the products of the endothermic reactions, meaning that the hydrogen content increases due to water-gas and steam reforming reactions. The small reduction of the produced hydrogen, from the temperature of 700 °C and so forth, results from the lack of methane. The same reactions that favor the production of H₂, are responsible for the CO increase, together with the Boudouard reaction. The constant decrease of the methane content results from the predominance of the steam reforming reaction at higher temperatures. Furthermore, the decrease of CO₂ can be attributed to the facilitation of Boudouard reaction in the expense of the combustion reaction of carbon. Finally, steam is consumed in the water-gas, water-gas-shift and steam reforming reactions. After the temperature of 650 °C, the flow of CH₄ decreases dramatically, limiting the steam reforming reaction. Furthermore, the increase in gasifier temperature impedes the water-gas-shift reaction, as it is mildly exothermic. These two aspects explain the upward trend of H₂O. In general, higher operating temperatures of the gasifier are more preferable since syngas with high heating values is produced. On the other hand, too high temperatures can cause ash agglomeration inside the gasifier.

Figure 3.

Effect of gasification temperature on syngas composition

The term AER is coined for the ratio between the air which is introduced in the gasifier and the necessary amount of air to achieve stoichiometric combustion. In Figure 4a the effect of AER on the syngas composition is illustrated. CO shows a large decrease of 79% (from 29.4% to 6.2%), while CO₂ rises from 15.6% to 20% (28% increase) as the AER varies from 0.2 to 0.6. The trends of CO and CO₂ can be attributed to combustion reactions. In practice, larger values of AER mean higher amounts of oxygen inside the gasifier, something that favors the combustion reactions. Complete carbon combustion is favored over partial combustion because of the higher oxygen concentration, which explains the increase of CO₂ and the decrease of CO. H₂ also decreases from 3.6% to 0.8% (78% decrease), whereas H₂O increases by 32% (from 10% to 13.3%). The downward trend of H₂ can be ascribed again to the dominance of combustion reaction of hydrogen, which causes also the rise of H₂O fraction as the AER increases. The percentage of methane is also lowered but its percentage is very low and thus it is not possible to show it in Figure 4a. In Figure 4b the impact on the system’s outputs when varying the AER is depicted: an increase in AER has a negative effect on the system efficiency. Particularly, net power production decreases in higher AERs. The same applies to the cooling effect, since less heat is transferred to the absorption system from cooling the exhaust gases exiting the burner. Heat duty in HEATEX6 depicts a constant trend since the same mass flow of exhaust gases enter the heat exchanger where the ambient water is heated up to 55 °C. System efficiency drops from 73% to 38% with an increase of the AER from 0.2 to 0.6. In general terms, the AER should be as low as possible to avoid the complete combustion of fuel, while ensuring in the same time complete conversion of carbon.

Figure 4.

Syngas composition versus AER (a); electrical, heating and cooling outputs versus AER (b)

In Figure 5 the effect of varying the desorber temperature on the cooling effect created in the evaporator, as well as on system efficiency, is presented. Since the changes made will only affect the absorption cycle, it is needless to show net electrical output and heating duty which shall remain stable. The optimal cooling output and thus the highest system efficiency is pointed at the temperature of 105 °C, 154 kW and 73%, respectively.

Figure 5.

Cooling output and total system efficiency versus desorber temperature

In general terms, higher desorber temperatures lead to lower ammonia concentrations because of water evaporation, although more ammonia goes to the gas phase, leading to a lower cooling effect. This trade-off between the higher ammonia flow rate and lower concentration in the solution results into the peak of the absorption cooling effect observed at 105 °C.

The TCI of the complete system amounts to about USD 2,326,000. The PEC is USD 540,200 while the ACI, based on eq. (23), is 81,400 USD/year. The fuel cost, which is, in this case, municipal solid waste, is taken to be zero. In Table 5 the purchased costs of the most expensive components are illustrated.

Regarding the energy savings if the system is to be implemented, it is observed the electricity and cooling demand of the building are fully covered. Excess electricity can be sold to the grid, but, since the case study focuses primarily on the energy savings, no further research will be conducted regarding the electricity market. The heating demand is covered up to 55% and the rest is supplied by a natural gas boiler. The average heat demand is 1,927,200 kWh, while the heat produced from the system is 1,061,856 kWh. The annual cost of natural gas needed in the boiler amounts 43,267 EUR/year.

The Payback Period (PBP) and Net Present Value (NPV) in Table 6 are calculated in order to investigate the economic feasibility of this system. The former indicator is a simple form of financial analysis, which takes into account the capital cost of the investment and compares it with the annual revenues that the system would create. The latter indicator allows the future value of cash flows to be adjusted to a reference year by applying an interest rate. The TCI reaches about 1,358,770 EUR, while the PBP is 4.5 years. The earnings from the energy savings are approximately 5 M EUR, based on the NPV for the system lifetime.