Optimization of power and hydrogen production from glycerol by supercritical water reforming

Chemical Engineering Journal 218 (2013) 309–318

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Optimization of power and hydrogen production from glycerol by supercritical water reforming Francisco Javier Gutiérrez Ortiz ⇑, Pedro Ollero, Ana Serrera, Sebastián Galera Departamento de Ingeniería Química y Ambiental, Universidad de Sevilla, Camino de los Descubrimientos s/n, 41092 Sevilla, Spain

h i g h l i g h t s " A process design is proposed and simulated for reforming glycerol using supercritical water. " The product gas is conditioned to obtain a hydrogen-rich gas stream, which is sent to a fuel cell. " The process includes energy integration and it is energy self-sufficient by burning the off-gas. " The process is assessed by an energy and exergy analysis. " Maximum power generation is obtained by an expander and a fuel cell.

a r t i c l e

i n f o

Article history: Received 16 July 2012 Received in revised form 6 December 2012 Accepted 10 December 2012 Available online 20 December 2012 Keywords: Reforming Supercritical water Glycerol Hydrogen Simulation

a b s t r a c t A process design is proposed and simulated for reforming glycerol using supercritical water aimed to produce maximum power and hydrogen in an energy self-sufficient system. The selected route takes advantage of the huge pressure energy of product gas just at the outlet of the reformer converting that into power by a turbine. The expanded product gas is conditioned by two water gas shift reactors and a pressure swing adsorption unit, so a hydrogen-rich gas stream is sent to a proton exchange membrane fuel cell to be converted into electrical energy and the pressure swing adsorption off-gas stream is used as fuel gas to provide the thermal energy required by the reforming process. The evaluation of the global efficiency of the process is carried out by energy and exergy analysis. Required glycerol feed concentration in aqueous solution was obtained for a self-sufficient process, both for pure and pretreated crude glycerol, at reforming temperatures from 600 to 1000 °C and 240 atm. Thus, reforming and preheating at 800 °C and 240 atm, it was obtained a power of 1592 kW per ton/h of glycerol, with exergy and energy efficiencies of 33.8% and 35.8%, respectively. Ó 2012 Elsevier B.V. All rights reserved.

1. Introduction The rising surplus of biodiesel-derived crude glycerol from the transesterification process requires a further processing. The crude glycerol stream leaving the homogeneous base-catalyzed transesterification reactor normally contains glycerol, methanol, alkalies (catalyst and soap), methyl esters and water. Normal refining of glycerol involves an acid treatment to split the soaps into free fatty acid and salts. Fatty acids are not soluble in glycerol and are separated from the top and recycled to the process; then, if methanol is recovered by vacuum distillation, some salts remain with the glycerol, which is in aqueous solution and its purity is about 85 wt.%; if methanol is present, glycerol purity may be reduced to about 75 wt.%.

⇑ Corresponding author. Tel.: + 34 95 448 72 68/60. E-mail addresses: [email protected], [email protected] (F.J. Gutiérrez Ortiz). 1385-8947/$ - see front matter Ó 2012 Elsevier B.V. All rights reserved. http://dx.doi.org/10.1016/j.cej.2012.12.035

A great deal of research has been carried out following a thermochemical route of glycerol valorization to obtain H2 or synthesis gas, by steam reforming [1–4], autothermal reforming [5,6] and aqueous phase reforming [7–10]. Many of these studies are centered on the development and characterization of catalysts. Supercritical water (SCW) is an emerging and promising medium to obtain hydrogen by reforming of glycerol, due to its relevant thermophysical properties such as a high capability to solubilize gaseous organic molecules and high diffusivity, among others [11–14] with only a few papers published regarding with the glycerol [15–17]. SCW is extremely reactive and, it may be possible to perform the process in the absence of a catalyst, although this premise requires to be experimentally verified. Hydrogen is very attractive as a clean fuel for proton exchange membrane fuel cells (PEMFCs), which are more efficient than combustion engines and have zero carbon emissions. In order to lessen the CO content and to increase the hydrogen content in the reformate stream, two water gas shift reactors (WGS), at two

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Nomenclature B_ H_ _ LW _ m Q_ S_ T _ W

exergy flow (kW) enthalpy flow (kW) exergy flow loss (kW) mass flow rate (kg/s) heat flow (kW) entropy flow (kJ/(s K)) temperature (K) power (kW)

Greek

g

efficiency (%)

temperature levels (HWGS and LWGS), are normally used. As a next stage, a pressure swing adsorption (PSA) unit or a preferential oxidation reactor (PrOx) may be used to reduce the carbon monoxide concentration to less than 10 ppm to avoid the poisoning of the PEMFC catalyst. SCW reforming is a new technology, which needs a better understanding of the fundamental phenomena for a suitable reactor and process design. Since it is in its beginnings, research on the key parameters influencing the overall process efficiency is essential. Some previous works have been carried out focused on fundamental aspects of the process of glycerol reforming using supercritical water such as a thermodynamic study [15,16,18,19], but there is only one published paper that deals with a methodology based on the exergy concept to assess the process [20]. An exergy analysis quantifies the lost efficiency in a process due to the loss in energy quality by evaluating the irreversibilities of the process, once the suitable control volume is selected, providing thus where the process can be improved. This paper specifically proposes a route of achieving maximum electrical power and efficiency in a process of glycerol reforming using supercritical water from the huge pressure energy of the reformate product by means of an expander and from pure hydrogen by a PEM fuel cell. In the process, a trend of different stages is included for the conditioning of the product gas (two WGS reactors and a PSA unit), once expanded to produce electrical energy, so the final purified hydrogen-rich gas can be used in a fuel cell. Moreover, the thermal energy required by the process to achieve an energy self-sufficient system is obtained by burning the PSA off-gas, so an external fuel is not needed as a heat source for performing the process. Thus, the proposed process is assessed by an energy and exergy methodology. Therefore, in the present study, the methodology described in the previous paper [20] is applied to an energy integrated process, and an optimization procedure is developed to achieve maximum power in the process as well as to reach a maximum exergy efficiency. These are the main aspects to be claimed within the paper novelty.

2. Process design and simulation The process design and simulation have been performed by coupling the reforming process with a turbine and a PEM fuel cell, using energy integration to improve the energy use. Fig. 1 shows the process flow-sheet. Appendix A gives details of the exergy analysis used to assess the process. Regarding with the energy integration, a number of exchangers were suitably located following a strategy based on minimizing the lost work or exergy losses, so small temperature-driving forces

Subscripts Gly glycerol in referred to some entering the control volume considered (unit, subsystem or system) MeOH methanol net referred to the net power of the turbine out referred to some leaving the control volume considered (unit, subsystem or system) PEMFC proton exchange membrane fuel cell PSA pressure Swing Adsorption WGS water Gas Shift

must be achieved, by using countercurrent flow and small temperature approaches at the ends of the exchangers. A ‘from inside to outside’ approach was used to design the heat exchanger network, in such a way that the internal streams close to each other at high thermal levels are the first to contact each other (the high-temperature hot streams heat the high-temperature cold streams) and the external more separated streams at low thermal levels are the last to contact each other (the low-temperature hot streams warm up the low-temperature cold streams), by taking the reformer as a central point, where the maximum temperature is required. Then, several possible arrangements of heat exchangers and streams pairing were performed in alternative flow sheets, and the redistribution of exchanger duties was assessed. The flow sheet that provides the best heat integration performance, i.e., minimize the use of external utilities and maximizing internal heat flows into the system, is finally selected. Besides, a series of gas-conditioning stages are added to the process for achieving a hydrogen-rich gas stream, which is fed into a PEMFC. A glycerol–water mixture is first pumped and then heated up to a desired temperature before entering the supercritical water reforming reactor. The heating is performed by means of five heat exchangers taking thus heat from hot streams leaving first the reformer-furnace, next the reforming reactor and then the two WGS reactors, which need to be cooled. In the reformer, glycerol is converted at temperature and pressure previously established. The reforming reactor is simulated as a Gibbs reactor, where the product composition and the heat of overall reaction are calculated under conditions that minimize the Gibbs free energy. The furnace–combustor where the PSA off-gas stream is burnt is simulated as a stoichiometric reactor, which transfer the heat released from the combustion reaction to the reforming reactor to operate at the specified temperature. The flow-rate of the gas to be burnt is computed as that needed to match the heat flow required in the reforming reactor. A fan is used to enter the air required in the furnace–combustor, so there is an O2 content of 3%vol in the flue-gas. The product gas leaving the reformer is expanded in a turbine converting thus the pressure energy into electrical energy, and then conditioned. By two adiabatic shift reactors in series with a cooler between them to remove the heat of reaction from the exothermic water gas shift reaction, the concentration of carbon monoxide produced in the reformer is reduced and the hydrogen content is increased. The first-stage reactor (HWGS) operates at 350 °C, and the second (LWGS) operates at 200 °C, which proceed on exothermal reaction at equilibrium, simulated by two adiabatic equilibrium reactors. The co-existence of CO, CO2 or CH4 has been reported to have a negative influence on the PEMFC [21], so a high purity hydrogen stream is beneficial to the process efficiency. The carbon monoxide

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11

09

P5 CW1PEM CW2PEM

G2

W PEM

W TUR

REFORMING REACTOR MIX1 HE01

HE02

HE03

HE04

WORK

TURB

HE05

QPEM

WORK

P1 GLY

COOLER

HEAT 02

03

04

05

06

07

08

01 WATER

R

SPL-PEM

Tail gas

CW3PEM

HWGS

LWGS ΔHPEM

10

GS

PEMFC

12

FAN1

HE09

HE08 AIR

to surroundings

G4

HEAT

G1

AIRC

H2FC1

AIRIN

P4

PSA

TGIN

Cooling water

W1

FURNACE AIR2IN

H2 CW01 P2

V1

W2

HE07 SG CW04

CW05

Hot water

SPLIT2

H2IN AIR2

CW02

TG1

MIX FAN

H2W AIR2C

H2FC2

V2 SEP

HE06

TG3 TG2

15

MIX2 MIX3

Separated water

P3

SW

W4 W3

G3 CW03

14

Fig. 1. Heat-integrated flow-sheet for producing hydrogen and electrical power via supercritical water reforming of glycerol.

concentration can be reduced to less than 10 ppm by means of a PSA unit [22]. Against the selective oxidation of CO (PrOx), which can be also used for this purpose, the PSA is probably the most mature technology for CO removal from the reformate gas aside from it is able to remove all the other gases present to obtain a very pure hydrogen at the PSA outlet. In addition, while PrOx is mainly used for a small scale, the PSA may be used for small-to-medium scale plants in stationary applications, such as the studied process. Since water vapor is strongly adsorbed in the PSA unit, the amount of water present in the gas must be reduced as much as possible before entering this unit to minimize the loss of its adsorption capacity. For this reason, the gas stream coming from the LWGS reactor is cooled down to 35 °C and the condensed water is separated from the gas stream by means of a gas–liquid separator. Furthermore, operating at low temperature allows to reduce the bed volume, and thus less of the H2 is lost during the bed regeneration steps [23]. The PSA process is based on sorbents that preferentially adsorb impurities at high partial pressures and then desorb them at a lower partial pressure (during the regeneration steps of the PSA cycle). Within the PSA unit, the adsorbent regeneration and desorption of components retained take place in the depressurization. Therefore, the removal of species adsorbed is done by total pressure reduction. This is simulated by disposing a valve at the PSA off-gas stream. On the other hand, the purified hydrogen stream exits the PSA bed at a pressure close to the PSA inlet pressure. Likewise, another valve has been disposed at the hydrogen stream to enter the PEMFC at atmospheric pressure. Increasing the pressure shifts the curves towards higher recovery although at pressures greater than 18 atm the change will be very small [23]. In addition, it is not convenient to operate the PSA unit at too high pressure in order to obtain more power in the expander. Therefore, the PSA unit has been simulated in a simplified way, by an ideal separator, although working at a realistic pressure and temperature (15 atm and 35 °C, respectively), and the pressure swing adsorber (PSA) is assumed to give a high enough purity (99.999%) with a hydrogen recovery of 80%, which are feasible val-

ues [24,25]. Thus, to yield a high purity hydrogen stream, some of the hydrogen is lost in the off-gas stream from the PSA, which is used as fuel in the furnace, and its energy is still used for heating the reforming reaction, reducing thus the amount of glycerol sent to the reformer to supply the required heat. The hydrogen-rich gas stream leaving the PSA unit is then heated up to 80 °C before it enters the PEM fuel cell to improve its performance. A simple PEM fuel cell process has been included in the scheme as a subsystem. Water is added to the air stream and the water mass flow rate is calculated to reach a relative humidity of 35% at 80 °C, inside the fuel cell. Similarly, water is also added to the hydrogen stream to reach a 100% relative humidity at 80 °C. Humidification of the entering streams is necessary to maintain membrane saturation and to prolong the life of the membrane. Water diffuses through the electrodes and membrane as vapor. Hydrogen coming from the reforming process diffuses through the negatively charged electrode (anode) to the membrane and reacts with oxygen coming from the positively charged electrode (cathode) and diffusing through the cathode electrode. Thus, the electrochemical reaction occurs in the membrane electrode assembly, which is simulated as an isothermal reactor (PEMFC in the scheme). The pure hydrogen obtained in the PSA enables to operate the fuel cell without splitting a part of the hydrogen in the anode off-gas, which only consists of water vapor because practically no net hydrogen exits the cell at steady state. For the simulation of the PEM fuel cell, anode and cathode channels are represented as the streams entering the reactor separately, and leaving it in a single stream. A fraction of the chemical energy of the electrochemical reaction is converted into electrical power, another fraction is internally used to complete air and hydrogen humidification process (some integrated fuel cell and membrane humidifier systems are being developed [26]) and to heat all the reactants to 80 °C, and the rest of the energy is released as heat leaving the fuel cell body toward, respectively, the cooling water and the surroundings. The air enters the fuel cell to have a flow rate of oxygen twice the flow rate required for stoichiometric oxygen–hydrogen reaction. A fan

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makes it possible to feed the air into the cathode channel of the fuel cell. Likewise, the waste heat produced in the stack module is removed through the cooling loop, which usually consists of a radiator, cooling pump, radiator fan. In this study, only the pump has been considered, and the cooling water flow rate is calculated to obtain a maximum increase of 15 °C through the heat exchanger. A description of the energy and exergy analysis of the PEMFC is given in the Appendix B. Specifications of the elements used in the simulation are shown in Table 1. The computation has been made with the aid of AspenPlus™ version 2006.5 (Aspen Technology, Inc., USA). The used thermodynamic method has been the predictive Soave–Redlich–Kwong (PSRK), as the most suitable [19]. Two feeds have been studied: pure and pretreated crude glycerol. The main components of crude glycerol considered were glycerol and methanol, as other researchers did [27], i.e., the starting point is taken just after separating the methyl ester fraction and performing the acid treatment, but before removing the methanol fraction. Thus, the pretreated crude glycerol feed

consists of glycerol with variable methanol content (10, 20 and 30 wt.%). Hydrogen, carbon monoxide, carbon dioxide, methane, ethane, propane, water, methanol, ethanol, glycerol and oxygen as well as pure carbon (in solid phase) were added manually, as they were considered the possible species from the SCW reforming of glycerol. The simulation did not predict coke formation for any of the experimental conditions in this paper, and any other compound has a mole fraction lower than 1012. The contents of ethane, propane and ethanol were also so low that they were not included in the analyses. Likewise, for all operating conditions simulated glycerol and methanol conversion were always 100%, at equilibrium condition. Based on reforming reactions (1) and (2), two hydrogen yields are computed using Eqs. (3) and (4): Glycerol reforming:

C3 H8 O3 þ 3H2 O ¼ 3CO2 þ 7H2

ð1Þ

Methanol reforming:

CH3 OH þ H2 O ¼ CO2 þ 3H2

ð2Þ

Table 1 Specifications of the individual process units for the simulation of the conceptual design of the SCW reforming of glycerol. Code

Equipment

Specifications

MIX1–MIX4 P1

Mixers Pump

HE01

Heat exchanger

HE02

Heat exchanger

HE03

Heat exchanger

HE04

Heat exchanger

HE05

Heat exchanger

R

Reforming Reactor

TURB

Turbine

HWGS (High-temperature stage) LWGS (Low-temperature stage) HE06

Water gas shift reactor Water gas shift reactor

SEP

L–G Separator

PSA (plus valves) HE07

Pressure swing adsorption unit

FURNACE

Furnace–combustor

FAN1–FAN2

Fans

HE08

Heat exchanger

P2–P5

Pumps

HE09

Heat exchanger

SPL

Splitter

PEMFC Reactor

Stoichiometric

COOLER

Heat exchanger

SPL-PEM

Splitter

Pressure drop: 0.0 atm Efficiency: 0.8 Outlet pressure: 240 atm Pressure drop: 0.0 atm (Tin)hot fluid  (Tout)cold fluid = 2 °C Pressure drop: 0.0 atm (Tin)hot fluid – (Tout)cold fluid = 3 °C Pressure drop: 0.0 atm Hot stream outlet temperature: 200 °C Pressure drop: 0.0 atm Hot stream outlet temperature: 350 °C Pressure drop: 0.0 atm (Tin)hot fluid – (Tout)cold fluid = variable on reforming temperature Operating temperature: variable (600–1000 °C) Pressure drop: 0 atm Type: Isentropic; Isentropic Efficiency: 0.72 Outlet pressure: 15 atm Adiabatic Reactor (heat duty = 0) CO + H2O = CO2 + H2 Adiabatic Reactor (heat duty = 0) CO + H2O = CO2 + H2 Pressure drop: 0.0 atm Hot stream outlet temperature: 35 °C Pressure drop: 0.0 atm Adiabatic It removes most the H2 (80%) from the other gases Outlet pressure: 1.1 atm Pressure drop: 0.0 atm (Tout)hot fluid – (Tout)cold fluid = 2 °C Combustion of everything able to be oxidized Operating temperature: 1000 °C Operating pressure: 1 atm Type: Isentropic; Isentropic Efficiency: 0.72 Outlet pressure: 1.1 atm Pressure drop: 0.0 atm Hot stream outlet temperature: 120 °C Efficiency: 0.8 Outlet pressure: 1.1 atm Pressure drop: 0.0 atm Hot stream outlet temperature: 80 °C Split fraction stream H2: variable (objective: zero) Pressure drop: 0.0 atm Heat + Work to SPL-PEM H2 + 0.5O2 ? H2O Isothermal Reactor: 80 °C Pressure drop: 0.0 atm Cooling water outlet temperature: 40 °C Auxiliary for simulation Fraction of work and heat

Heat exchanger

Heat exchanger

F.J. Gutiérrez Ortiz et al. / Chemical Engineering Journal 218 (2013) 309–318

Pure glycerol:

gH2 ¼

mol=h H2 1  mol=h C3 H5 ðOHÞ3 7

ð3Þ

Pretreated crude glycerol:

gH2 ¼

mol=h H2 mol=h C3 H5 ðOHÞ3  7 þ mol=h CH3 ðOHÞ  3

ð4Þ

313

Although the effect of pressure was proved to be little significant in a previous study [19] sensitivity analysis was also carried out by ranging the pressure from 200 to 300 atm. The results of such simulations verified the small impact of pressure on the process performance (changes in the total net power less than 1%), and all simulations described in this paper were carried out at a reforming pressure of 240 atm. 3.1. Optimization of power and hydrogen production

3. Results and discussion In this study, the process was assessed using a mass flow-rate of glycerol (pure or pretreated crude) fed into the system of 1000 kg/ h. This glycerol stream may be mixed with more or less water changing thus the feed concentration to the plant. The aim was to achieve not only maximum power in the expander but also maximum hydrogen production to send to the PEMFC, and keeping the energy self-sufficiency of the system by burning the PSA off-gas. With this target, the analyses were carried out over the following variable ranges: reforming temperature of 600– 1000 °C, glycerol feed concentration of 5–55 wt.%, for both pure and pretreated crude glycerol. MeOH content in the crude glycerol was changed from 10 to 30 wt.%. As relevant outputs for assessing the process, the gross and net power of the turbine, the molar flowrates of hydrogen routed to PEMFC along with the power obtained with this device and the exergy losses were used. The net power of the turbine is calculated as the power obtained in the expander minus the power required in the fans and the pumps. The conversion from electrical into mechanical energy and vice versa was considered as 100%. Likewise, exergy, combined heat-power exergy, energy efficiencies and hydrogen yields were also used for the assessment.

For the reforming temperature range above mentioned, the conditions to reach a self-sufficient process with no use of external fuel and achieve maximum power and hydrogen production were inspected. The conditions match the cases in which hydrogen was completely separated from the PSA unit, and the required glycerol concentration was computed for achieving the self-sufficiency, by burning the PSA off-gases (that contains a small fraction of the produced hydrogen) in the furnace where the heat needed to sustain the process is released. To verify this, a sensitivity analysis was carried out for each reforming temperature, by changing the glycerol concentration between 5 and 50%wt. By this way, conditions from a heat deficit in the reactor to a heat excess are obtained corresponding to cases with lower and higher glycerol concentration, respectively. Thus, e.g., for a reforming temperature of 800 °C, the heat deficit is obtained if the glycerol concentration is lower than 16 wt.%, and the heat surplus is got if a glycerol concentration higher than 27 wt.% is used. Between these limits, an energy selfsufficient reforming is possible, and by increasing the glycerol concentration the fraction of pure hydrogen separated from the PSA unit to be burnt in the furnace (apart from the PSA off-gas) is lower. There is a specific value in which the hydrogen-rich gas splitted is null, so all the hydrogen obtained in the PSA unit is routed to the PEM fuel cell. For this condition, from an energy and exergy point

(a)

(b)

(c)

(d)

Fig. 2. Mass flow rate of hydrogen sent to PEMFC and hydrogen fraction sent to the Furnace at: (a) 700 °C, (b) 800 °C, (c) 900 °C, and (d) 1000 °C (pure glycerol, 240 atm).

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of view, the optimum is achieved as shown in Fig. 2, for four reforming temperatures (700, 800, 900 and 1000 °C) and, equally, the maximum total net power is reached as illustrated in Fig. 3, where the gross and the net power of the turbine (once the consumption of other machines is discounted), as well as the power in PEM fuel cell and the total net power in the entire process are represented. In Fig. 3, it can be observed that at low glycerol concentrations (less than about 15 wt.%), all the product gas must be sent to the furnace to achieve the energy self-sufficiency. Indeed, there is a value below which the energy demand for heating the big amount of water is higher than that provided by the system itself and the process cannot be performed even though all the expander power is converted into heat. By increasing the glycerol concentration, it becomes possible to send a fraction of H2 to the PEM fuel cell. Although the hydrogen yield decreases as the glycerol feed concentration increases, more power is obtained in this unit because a higher fraction of hydrogen is sent to the fuel cell. Beyond the glycerol concentration in which all the hydrogen is routed to the fuel cell, the process provides a heat flow surplus and the power obtained in the fuel cell is lower because the hydrogen yield lessens. On the other hand, the power generated by the expander and consumed by pumps and fans decreases as the flow rate decreases, i.e., as the glycerol concentration increases, since the study was performed using a constant mass flow rate of glycerol. Therefore, maximum total power matches the glycerol concentration at which maximum power is obtained in the PEM fuel cell.

In addition, other effect obtained (but not shown) was the increase in hydrogen molar flow-rate when rising the reforming temperature, as well as the turbine power, due to the higher thermal level of the product gas and the higher product-gas molar flow-rate (dry basis), which contains more chemical energy available in the furnace–combustor. Likewise, the total net power, which takes into account the PEM fuel cell power, increases once hydrogen-rich gas is routed to the PEMFC, achieving the maximum, once again, just when all the hydrogen separated in the PSA is sent to the fuel cell. With the aim of giving more details of the analyses carried out, the case of 800 °C reforming temperature for pure glycerol has been selected. Table 2 shows the overall energy balance, illustrating the gross power obtained in both the expander and the PEM fuel cell, which are of similar magnitude, as well as the power consumed by pumps and fans. The net power may be also obtained by an overall energy balance of the process. Table 3 depicts the heat flows, the temperatures for all the ten heat exchangers used and the mass flow rates of fluids, for achieving maximum power and hydrogen production. It can be observed that the two highest heat duties are obtained for those heat exchangers located at the supercritical reformer inlet and for cooling down the product gas to condensate water before entering the PSA unit (HE05 and HE06, respectively). Likewise, the cooling water, with a mass flow rate five times higher than that of the feed, may be used as hot water since it leaves the heat exchanger HE07 at 84 °C, making it possible to increase the overall efficiency of the process (see Appendix B). Table 4 collects the changes in exergy flows, work and heat flows,

0

Fig. 3. Turbine gross power, turbine net power of the turbine (once consumption in other machines is discounted), PEMFC power and total power obtained in the overall process at reforming temperatures of: (a) 700 °C, (b) 800 °C, (c) 900 °C, and (d) 1000 °C (pure glycerol, 240 atm).

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F.J. Gutiérrez Ortiz et al. / Chemical Engineering Journal 218 (2013) 309–318 Table 2 Overall energy balance for reforming of glycerol using SCW (for pure glycerol at 800 °C and 240 atm). Work entering the system (kW)

Enthalpy of the inlet-streams (kW)

P1 P2 P3 P4 P5 FAN1 FAN2

Glycerol Water Air Air2 W1 W3 CW1PEM CW01

44.8 0.1 4.7E03 5.8E03 0.2 12.4 12.8

Work leaving the system (kW) TURB PEMFC

2053.6 12233.9 0.2 0.2 1912.7 1722.3 112797.5 88218.9

Enthalpy of the outletstreams (kW) 788.7 873.4

Tail gas (G4) Separated Water (SW) CW5 GS CW3PEM

5220.6 11200.0 86832.8 4935.9 112344.6

Note: A heat flow of 3.3 kW is released from the PSA unit due to the condensation of a part of the small water fraction still present in the gas entering the PSA unit.

and exergy losses in the individual process units, and it shows that the heat released in the combustion of the PSA off-gas is just the heat flow required by the reforming reactor (1370.6 kW). Likewise, it can be seen that 28.5% of the exergy losses are due to the PEM fuel cell. Similarly, the contribution of the reaction section of the plant (coupled reformer-furnace) and all the heat exchangers are 42.2% and 18.9% of the total exergy losses, respectively. The overall exergy efficiency of the process is 33.8%, under these conditions. The combined heat-power exergy efficiency is 35.4%. Likewise, the energy efficiency is 35.8%, and the hydrogen yield is 47.8%. Similar results are obtained for the other reforming temperatures simulated, and they are partially depicted in Table 5, which summarizes the exergy analysis of the pure glycerol simulations, showing the overall thermal efficiency, the overall exergy efficiency, the combined heat-power exergy efficiency and the hydrogen yield. It can be observed that the energy, exergy and combined efficiencies increase as the reforming temperature rises as well as the required glycerol feed concentration does. Finally, a scheme of an energy self-sufficient process using a CO-PrOx reactor instead of a PSA unit was simulated, so 99% of CO is converted into CO2 and 1% of the H2 is burnt, operating at 150 °C, and the gas leaving this unit was cooled down to 80 °C. Hence, the heat exchanger HE07 was not longer necessary. The reformate gas was expanded to 1.1 atm, and 20% of the gas entering the PEMFC was anode off-gas to be burnt in the furnace. The other specifications were the same. By reforming at 800 °C and using the same glycerol feed concentration (26.5 wt.%), the exergy, energy and combined heat-power exergy efficiencies were 29.8%, 31.8% and 31.0%, respectively. The H2 yield was 47.8%. A total net

power of 1379.4 kW was obtained, which is 13.4% lower than that obtained using the scheme shown in Fig. 1. Thus, the process simulated including a CO-PrOx instead of the PSA stage resulted in a worse performance. 3.2. Assessment of the process performance when feeding pretreated crude glycerol When pretreated crude glycerol is fed into the system, the molar flow rate of the feed increases as compared with pure glycerol feed as the methanol concentration increases, since the total mass flow rate is the same but the molar weight of methanol is lower than that of glycerol. Besides, the higher hydrogen yield of methanol reforming than that of glycerol reforming makes it possible to obtain higher hydrogen yields and hydrogen molar flow-rate. Similar graphical representations to Figs. 2 and 3 are obtained but not shown. Instead of it, Table 5 shows the results of simulations run for pretreated crude glycerol. It can be observed that the crude glycerol feed concentration is lower for pretreated crude glycerol at the same reforming temperature as the methanol concentration increases. Thus, the flow rates of water fed into the reformer will rise as compared to those required in the pure glycerol case, increasing thus the total flow rate entering and leaving the reformer, and hence, the gross power of the turbine and the power of the machines that consume energy. Also, exergy and combined efficiencies give somewhat higher values than those for pure glycerol, although the exergy losses are slightly higher. In addition, since the total molar flow rate is higher, the molar flow rate of hydrogen increases as the methanol concentration increases. Therefore, the electrical power produced in the PEM fuel cell is higher. As it can be observed in Table 5, the above differences are accentuated when the methanol concentration in the feed is increased. In case of reforming at 800 °C, when the MeOH content in the feed is 30 wt.%, the net power produced by the expander and by the PEM fuel cell was 767.8 kW and 940.4 kW, respectively, obtaining an H2 yield of 48.2% (0.4% points higher than in the case of using pure glycerol). The glycerol concentration is 19.6 wt.%, but the methanol concentration is 8.4 wt.%, so less water is used than that for pure glycerol. Likewise, the exergy losses are 6% higher, but the increase in exergy and combined heat-power exergy efficiencies are 0.3% and 0.4% higher, respectively, because the total net power is also higher. Similarly, the increase in the energy efficiency is 0.6% higher. Since these differences are less accentuated as the methanol concentration in the feed decreases, it can be concluded that although the presence of methanol in the pretreated crude glycerol is beneficial to the process, it must be accounted for a material loss in the biodiesel production. Due to the small improvement obtained, it is recommendable to recover the methanol in the pretreatment of crude glycerol to return it to the biodiesel production process.

Table 3 Heat flows, temperatures and mass flow rates of fluids for all the ten heat exchangers (pure glycerol at 800 °C and 240 atm). Heat exchanger

Q_ (kW)

(Tin)cold

(Tout)cold

(Tin)hot

(Tout)hot

_ cold (kg/h) m

_ hot (kg/h) m

HE01 HE02 HE03 HE04 HE05 HE06 HE07 HE08 HE09 COOLER

229.5 379.6 400.1 283.5 1204.9 1639.1 23.5 302.1 10.4 452.7

30.1 83.2 172.7 263.2 318.6 25.0 24.1 36.4 35.5 25.0

83.2 172.7 263.2 318.6 441.1 95.0 82.1 309.6 80.0 40.0

95.0 202.1 381.0 473.7 1000.0 175.7 85.2 320.6 129.6

85.2 175.7 200.0 350.0 320.6 35.0 84.1 129.6 122.8

3774.9 3774.9 3774.9 3774.9 3774.9 20009.9 1163.6 3865.6 58.6 25584.8

20009.9 3774.9 3774.9 3774.9 5029.2 3774.9 20009.9 5029.2 5029.2

Temperatures in °C.

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Table 4 Changes in exergy flows, work and heat flows, and exergy losses in the individual process units (pure glycerol at 800 °C and 240 atm). Equipment

DB_ (kW)

MIX1 P1 HE01 HE02 HE03 HE04 HE05 R TURB HGWS LWGS HE06 SEP PSA + valves HE07 FURNACE FAN1 HE08 P2 HE09 FAN2 P3 P4 P5 MIX2 MIX3 COOLER PEMFC Total exergy losses

24.2 31.5 19.7 33.3 29.9 26.1 147.1 541.5 923.6 11.2 0.2 272.4 0.01 103.9 3.9E03 1857.6 9.0 21.7 0.07 1.6 9.3 1.1E03 1.9E03 0.1 1.6 2.5 11.0 1811.3

_ in (kW) W

_ out (kW) W

Q_ in (kW)

Q_ out (kW)

DS_ cold (kW/K)

DS_ hot (kW/K)

_ (kW) LW 24.2 13.3 19.7 33.3 29.9 26.1 147.1 508.3 134.8 11.2 0.2 272.4 0.01 103.8 1.9 807.8 3.4 25.4 0.02 1.6 3.5 3.5E03 3.9E03 0.1 1.6 2.5 59.5 890.3 3122.0

44.8 229.5 379.6 400.1 283.5 1204.9 1370.6

229.5 379.6 400.1 283.5 1204.9

0.7 1.0 0.8 0.5 1.8

0.6 0.8 0.7 0.4 1.3

1639.1

1639.1

4.9

4.0

23.5

3.3 23.5 1370.6

0.1

0.1

302.1

302.1

0.7

0.6

10.4

10.4

3.2E02

2.6E02

788.7

12.4 0.1 12.8 4.7E03 5.8E03 0.2

452.7 873.4

452.7

Table 5 Comparison on the thermodynamic performance of the overall process for conditions to maximize work and hydrogen production versus feed type and reforming temperature. Temperature (°C)

Feed concentration* (wt.%)

_ LW (kW)

_ Tgross W (kW)

_ Tnet W (kW)

_ PEMFC W (kW)

gEnergy

gExergy

gcombined

gH2

(%)

(%)

(%)

(%)

600 700 800 900 1000

17.9 21.5 26.5 32.9 42.8

3489.0 3285.1 3122.0 3017.8 2921.8

839.6 835.6 788.7 690.7 603.0

748.8 755.1 718.6 629.5 549.5

411.6 645.4 873.4 1087.4 1274.4

26.1 31.5 35.8 38.6 41.0

25.0 29.9 33.8 36.3 38.4

28.4 32.4 35.4 37.3 39.0

22.6 35.4 47.8 59.6 69.8

600 700 800 900 1000

15.8 19.0 23.5 29.1 38.0

(1.8) (2.1) (2.6) (3.2) (4.2)

3557.5 3348.8 3181.8 3074.8 2976.9

858.6 854.4 807.0 708.0 618.1

765.7 773.1 734.6 644.5 562.6

423.2 662.7 895.5 1114.5 1305.1

26.3 31.7 36.0 38.9 41.3

25.1 30.0 33.9 36.4 38.6

28.5 32.5 35.6 37.5 39.1

22.7 35.5 47.9 59.7 69.9

(20 wt.% MeOH)

600 700 800 900 1000

13.8 16.6 20.5 25.5 33.3

(3.5) (4.2) (5.1) (6.4) (8.3)

3629.0 3414.8 3244.3 3135.0 3034.7

877.4 873.5 825.3 724.5 623.4

781.6 789.4 751.2 658.5 575.5

435.0 680.5 917.8 1141.1 1336.0

26.4 31.9 36.2 39.0 41.5

25.1 30.1 34.0 36.5 38.7

28.6 32.6 35.7 37.6 39.3

22.8 35.6 48.0 59.7 69.9

(30 wt.% MeOH)

600 700 800 900 1000

11.8 14.3 17.7 22.0 28.7

(5.1) (6.1) (7.6) (9.4) (12.3)

3700.4 3481.6 3308.1 3196.3 3093.6

896.7 892.6 843.6 741.1 648.3

798.3 805.7 767.8 673.2 589.0

447.3 698.3 940.4 1168.0 1367.0

26.5 32.0 36.4 39.2 41.7

25.2 30.2 34.1 36.6 38.7

28.7 32.6 35.8 37.7 39.4

22.9 35.8 48.2 59.8 70.0

Feed type

Pure glycerol

Crude glycerol (10 wt.% MeOH)

Note 1: At a reforming temperature of 600 °C, HE04 is out of service, since the product gas temperature is 288.3 °C. Note 2: For pretreated crude glycerol, in the column of feed concentration there are two numbers: pure glycerol wt.% (methanol wt.%); the sum is the crude glycerol feed concentration. * The first number is the glycerol feed concentration (wt.%) and the number between brackets represents the methanol concentration (wt.%) in the feed.

3.3. Simulation results compared with some experimental results Finally, a comparison of the results obtained in this study with other published results was carried out in the main process unit, i.e., the supercritical reformer reactor, which was simulated as a Gibbs reactor where chemical equilibrium was assumed. Although

the number of publications on the reforming of glycerol using supercritical water is fairly limited, the few experimental results found in the literature match reasonably well with the results obtained in this study based on the assumption of reaching chemical equilibrium. Thus, at feed concentration of 30 wt.% glycerol, Byrd et al. [15] obtained a product gas composition H2/CO/CO2/CH4 of

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47.2/3.2/34.0/15.6, at a temperature of 800 °C and a pressure of 241 bars, over Ru/Al2O3 catalysts in a tubular fixed-bed flow reactor. In the present study, the product gas composition H2/CO/CO2/ CH4 was 47.5/11.1/25.4/15.9, when using a glycerol feed concentration of 26.5 wt.%, at 800 °C and 240 atm. At 15 wt.% glycerol, Byrd et al. obtained a product gas composition H2/CO/CO2/CH4 of 57.9/0.6/30.9/10.7, and in this study, at 14 wt.% glycerol, the product gas composition H2/CO/CO2/CH4 was 59.4/9.8/24.4/6.4, using the same pressure and temperature. Experimental yields of CO and CO2 are somewhat different (smaller and higher, respectively) from the predicted values at equilibrium, probably because the water–gas shift reaction takes place in presence of the used catalyst.

B_ ¼ H_  T 0 S_

4. Conclusions



where H_ and S_ are the total stream enthalpy and entropy flows, respectively, evaluated at the corresponding state, i.e., the liquid, vapor or supercritical state. For all the streams entering (J) and leaving (K) the system, it results

DB_ outin ¼

Acknowledgments This research is supported by the Science and Technology Ministry of Spain under the research Project ENE2009-13755, as a Project of Fundamental Research inside the framework of the National Plan of Scientific Research, Development and Technological Innovation 2008–2011. Appendix A. Exergy analysis The studied process is continuous and steady-state; it is open to interactions with the surroundings, and potential and kinetic energies are not considered. Likewise, the reactors are at equilibrium condition. The dead-state conditions (of surroundings) are 25 °C and 1 atm. All flows (mass, energy, exergy) entering the system are considered as positive, while flows exiting are indicated by a negative sign. The terms of energy and exergy flows refer to energy and exergy per unit of time (power). Likewise, the terms of heat flow and work flow will be used, although power will be also used for the latter. By combining energy and entropy balances, the lost work flow _ for the overall process can be (exergy flow) or lost power (LW) computed as follows:

_ ¼ LW

X

_  DB_ outin  W

outin

 X T0 _ 1 Qj Tj outin

ðA1Þ

_ ¼ T 0 DS_ irr is the lost work or exergy losses due to the inwhere LW _ is the power relative to crease in entropy of the Universe. W mechanical shaft work and electrical work as well as work resulting from the expansion or contraction of the system against the surroundings, computed as follows:



X

outin

q h X X _ kÞ þ _ kÞ _ ¼  ðW ðW W out in k¼1

ðA2Þ

k¼1

Electricity and mechanical work are perfectly convertible and, for these forms, exergy contents equals to the energy content. B_ is the exergy flow of a stream, which is calculated via:

J X

ðB_ s Þin 

K X ðB_ s Þout

ðA4Þ

s¼1

s¼1

Q_ j is the heat flow transferred to the system (positive) or by the system (negative) at Tj, which is the temperature of the jth heat source or sink outside the system. Only a part (given by the Carnot factor) of the heat transferred is a useful work, so:

X

outin

A SCW reforming of glycerol process was designed and simulated, by including energy integration and conditioning stages (HWGS, LWGS and PSA), to obtain a hydrogen-rich gas for a PEMFC. Maximum net power production by an expander (718.6 kW @ 800 °C, 240 atm), once consumption in other machines was discounted, and by a PEMFC (873.4 kW @ 80 °C, 1 atm), once pure hydrogen production was maximized from the PSA unit (58.6 kg/ h), was achieved by feeding 1000 kg/h of glycerol in an aqueous solution (26.5 wt.%). The combustion of the PSA off-gas provided the heat required by the reforming process.

ðA3Þ

   p  m  X X T0 _ T0 T0 1 1 ðQ_ j Þoutþ 1 ðQ_ j Þin Q j¼ Tj T j out T j in j¼1 j¼1

ðA5Þ

In the reforming process, the lost power is the net sum of exergy _ minus the net power produced flows of inlet and outlet streams (B) _ net ) since the overall reforming process is adiain the system (W batic. However, since the fuel cell system is not adiabatic the heat flow transferred to the surroundings by this subsystem must be taken into account (see Appendix B). Therefore:

_ LW¼

X

  X X _ out þ W _ in  1 T 0 jQ_ cooling j W T FC

ðB_ in B_ out Þ

system

ðA6Þ

As the final goal is to get maximum power through the huge pressure energy of the reformate gas, by means of an expander, plus the conversion of hydrogen energy, by means of a fuel cell, the exergy efficiency of the process to be inspected is the following:

gexergy ¼

_ net W _ _ W net  LW

ðA7Þ

_ net includes both the net power of the turbine, The term W once the power by other machines is discounted, and the fuel cell. The term of Lost Power includes internal and external exergy losses, both in the overall reforming process and the PEM fuel cell system. In the term of Lost Work (flow), both internal and external exergy losses are considered. The former are due to the system itself and their units and the latter is a wasted exergy no longer usable by subsequent processes due to dissipations to the environment as heat losses in individual process units (not included in this study, where a perfect thermal isolation has been assumed), cooling water or smokestack effluents. Anyway, the cooling water heated in the process could be used as a local source of heating, i.e., in the plant and other surrounding plants, next to the concept of district heating. As a result, a combined heat-power exergy efficiency (gh–p) may be defined as follows:

ghp ¼

_ net  P DB_ cooling water W 8EX _ net  LW _  P DB_ cooling water W 8EX

ðA8Þ

where

DB_ cooling water ¼

X ðB_ out  B_ in Þ

cooling water

ðA9Þ

Finally, for pure glycerol, the energy efficiency is defined as

_ W

net genergy ¼ _ mGly LHV Gly

ðA10Þ

_ Gly is the mass flow rate of glycerol fed into the system and where m LHV Gly is the Lower Heating Value of glycerol (16 MJ/kg). In case of using pretreated crude glycerol, the energy efficiency is calculated as

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_ W

net genergy ¼ _ _ MeOH LHV MeOH mGly LHV Gly þ m

ðA11Þ

_ MeOH is the mass flow rate of methanol fed into the system where m and LHV MeOH is the Lower Heating Value of methanol (19 MJ/kg). More information can be found elsewhere [20]. Appendix B. Energy and exergy analysis of the PEMFC The generic energy balance in the stack is the following:

X

H_ in 

X

_ PEMFC þ Q_ cooling H_ out ¼ W

ðB1Þ

where H_ in is the enthalpy flow of the reactants gases in, H_ out is the enthalpy flow of the unused reactants gases and the product gases _ PEMFC is the electricleaving the fuel cell (all the water is as vapor), W _ ity generated, Q cooling is the heat flow taken away from the stack by active cooling. The stack is assumed to be well thermally isolated. Assuming negligible potential and kinetic energy effects on the fuel cell electrochemical process, the total exergy transfer per unit mass of each reactant and product consists of the combination of both physical and chemical exergies. Thus, using Eq. (A1), the exergy loss of the PEM fuel cell unit has been computed as follows:

_ PEMFC ¼ B_ out þ LW

X

 B_ in  1 

T0 T PEMFC

 _ PEMFC jQ_ cooling j  W

ðB2Þ

where TPEMFC is the temperature inside the PEM fuel cell. To obtain the exergy losses of this subsystem, the lost work of the fan, pump and cooler must be taken into account, and so it has been computed in the analyses, in a similar way to other similar units. Specifically, to calculate the unusable waste exergy of the PEM fuel cell through the cooler, the following equation has been used, using Eq. (A1):

 _ cooler ¼ B_ out þ B_ in  1  LW

T0 T PEMFC

 jQ_ cooling j

ðB3Þ

where the heat source is the PEM fuel cell at TPEMFC and the heat sink is the cooling water through the cooler. The electric power generated by the fuel cell is calculated as:

_ PEMFC ¼ g m _ W FC H2 LHV H2

ðB4Þ

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