Hydrocarbon Processing 03 2012

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MARCH 2013 | Volume 92 Number 3 HydrocarbonProcessing.com

8

34 SPECIAL REPORT: CORROSION CONTROL

35 Monitor and minimize corrosion in high-TAN crude processing S. Ghoshal and V. Sainik

39 Mitigate reactor failures due to graphitization A. Al-Meshari, A. Abdelgalil and S. Al-Enazi

43 Consider dewpoint corrosion in reactor design K. Ramesh

DEPARTMENTS

4 8 11 15 106 109

Industry Perspectives Brief Impact Innovations Marketplace Advertiser index

51 Corrosion under insulation is a hidden problem T. Hanratty

BONUS REPORT: SAFETY DEVELOPMENTS

COLUMNS

21

Reliability How the best lubricant provides added value

23

Integration Strategies Wireless sensing benefits from new technology

25

Boxscore Construction Analysis Saudi Arabia’s plan for near-zero-sulfur fuels

55 Process safety management: Going beyond functional safety M. A. Turk and A. Mishra

65 Operator response to alarms is an important protection layer R. Limaye

71 Key aspects of design and operational safety in offsite storage terminals V. Ramnath, Aker Solutions, Pune, India

CATALYST—SUPPLEMENT

77 Optimize feed treatment for polypropylene process

31

H. Poorkar, H. Shbrain, S. Al-Harbi and A. Al-Saeed

78 Catalyst news REFINING DEVELOPMENTS

95 Improve benzene control L. McDermott and A. Malik

CLEAN FUELS

99 FCC catalyst design evolves to maximize propylene C. Pouwels and K. Bruno

Cover Image: Catalytic reformer No. 3 at BP’s Kwinana refinery, Western Australia.

110

Viewpoint Refiners have a new learning curve with shale oil Engineering Case Histories Case 71: Statistical visual data can be useful in troubleshooting—Part 2

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Industry Perspectives EU leans forward on clean-fuel strategies In late January 2013, the European Commission (EC) announced ambitious measures to ensure the buildup and standardization of alternative fueling stations across Europe. Much of the earlier EC policy initiatives addressed fuels and vehicles, but offered little guidance on distribution. Vicious cycle. In the press release, the EC noted that the implementation of “clean-fuel” programs are hindered by three primary barriers: • High cost of alternative-fuel vehicles • Low consumer acceptance of alternative-fuel vehicles • Lack of recharging and refueling stations. Refueling stations are not constructed because there are not enough vehicles. Vehicles are not sold at competitive prices because demand is low. Consumers do not buy the alternative vehicles because they are expensive and there are insufficient fuel stations. The EC is proposing a package of binding targets for Member States requiring a minimum level of infrastructure to support clean fuels, along side with common European Union (EU)-wide standards for distribution equipment. Standardization efforts. The new measures seek to simplify

distribution issues for vehicles powered by electricity, hydrogen and natural gas. Electricity. Across the EU, electric charging points vary greatly. Germany, France, The Netherlands, Spain and the UK are making progress with electric-refuel stations. The EC announced the “Type 2” plug as the common standard for recharging electric vehicles. EU automobile makers have sought a common plug to build alternative-fuel stations. Hydrogen. A significant number of hydrogen fuel stations are operating in Germany, Italy and Denmark, although not all are publically accessible. Under the proposal, existing hydrogenfuel stations would form a network with common standards to ensure fueling capabilities. Natural gas—compressed and liquefied. In the EU, over one million vehicles, mostly passenger cars, are powered by compressed natural gas (CNG) and represent only 0.5% of the vehicle fleet. By 2020, the industry aims to increase the CNG vehicle fleet to 10 million units. Such an aggressive target requires publically accessible refueling stations. These stations must have standardize equipment and be available EU-wide. Further development of alternative-fuel vehicles will require cooperation, standardization and investment. The EC estimates that approximately €10 billion will be necessary to build the minimum infrastructure for alternative fuels by 2020. The proposal is based on the premise that the market will grow and finance itself without the need to involve public funds. An expanded version of Industry Perspectives can be found online at HydrocarbonProcessing.com. 4MARCH 2013 | HydrocarbonProcessing.com

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)/(;,7$//,&6$)(,6025(7+$1$352*5$0,7·6$:$ 455°C). As discussed in this case history, the shell of a methyl tertiary butyl ether (MTBE) reactor became susceptible to elevated-temperature oxidation that initiated corrosion conditions. In addition, high-temperature fluctuations fostered stresses that promoted fatigue of the MTBE reactor shell and, ultimately, caused an emergency shutdown of the unit.

structure representing a plane of weakness. Chain graphitization is usually favored along planes of localized yielding or in regions that experience plastic deformation due to cold working or bending. This process leads to considerable reduction in stressrupture strength and fatigue resistance of the steel. Furthermore, chain graphitization can cause a significant reduction in the fracture toughness and, consequently, facilitate fast crack growth.2,3

Background. Cracking in the reactor shell of an MTBE unit

Case history. An emergency shutdown of an MTBE unit oc-

was detected by hot-air leaks through the reactor. This failure resulted in the emergency shutdown. Several investigation techniques were used to identify the root cause for the failure, including X-ray fluorescence (XRF), C/S analyzer, optical microscopy, scanning electron microscope/energy dispersive Xray (SEM/EDX), tensile testing and hardness testing. Results from the thorough investigation identified that the cracking of the reactor shell was attributed to corrosion fatigue (oxidation fatigue), due to thermal cycling and graphitization. Graphitization was the root cause for the degradation of the steel mechanical properties and for significant reductions in the fracture toughness, thus facilitating fast crack growth. The recommendation from this failure and study was the practice to inspect all reactor shells for graphitization. The rise in the shell-metal temperature may have been attributed to inefficient thermal insulation provided by the refractory lining. For this MTBE process, the metal temperature of the reactor should be kept below 427°C to prevent graphitization. The condition of the reactor shells must be evaluated using metallographic techniques. Ideally, the extent of graphitization can be investigated through representative sampling for metallographic examination. Another option is upgrading the construction material for the reactor shell to a chromium (Cr)containing, low-alloy steel.

curred due to cracking of the main reactor shell. This reactor had been in service for 15 years, and it is constructed of ASTM A515 Grade 60 CS and lined with refractory. In the reactor, isobutene is converted into isobutylene in the presence of catalyst. In this process, hot isobutene (T ≈ 600°C) is fed to the reactor at 170 tph. Then, the reactor is purged with steam at 250°C–260°C. Finally, hot air (T ≈ 675°C) flowing at 630 tph is introduced to regenerate the catalyst. Each cycle takes about eight minutes.

Definition of graphitization. Graphitization occurs in CS as an end result from prolonged exposure (over 40,000 hr) to high-temperature (> 455°C) process conditions. It is a process in which the pearlite decomposes into ferrite and randomly dispersed graphite. It has been reported that graphitization is accelerated by high stresses and/or temperature fluctuations.1 Two types of graphitization were observed in the CS microstructure—random and “chain” graphitization. In the random graphitization, the graphite nodules are scattered randomly across the microstructure, leading to degradation in steel mechanical properties, e.g., strength and hardness. In the chain graphitization, the graphite nodules form an aligned

Investigation. Three plate samples from the reactor shell were submitted for analysis: A, B and C. Cracks were observed on the three samples, as shown in FIGS. 1–3. The cracks appeared to have initiated at the bolt holes (attached to the shell internal surfaces) and propagated through the shell plate. Visual examination of areas near the cracks revealed the presence of multiple parallel cracks. The internal surfaces of the three samples had been covered with grayish and reddish deposits that were col-

FIG. 1. General view of reactor sample A. Note: The cracks initiated at the bolt holes attached to the shell-plate internal surface. Hydrocarbon Processing | MARCH 2013

39

Corrosion Control TABLE 1. Chemical composition of the reactor shell plates, wt% C 4

ASTM A515 Grade 60 Reactor shell plate

0.24 (max) 0.13

Mn

P

S

Si

0.98 (max) 0.035 (max) 0.035 (max) 0.88

0.032

Al

Cr

Ni

Cu

Fe

0.03

0.03

0.03

0.03

bal

0.13–0.45

0.006

0.22

bal

TABLE 2. Chemical composition of the deposits collected from the sample internal surfaces, wt% O

Al

Si

P

Ca

Mn

Fe

Zn

Sample A

30.4

0.2

0.3

0.1

0.2

0.6

bal

0.3

Sample B

30.2

0.2

0.4

0.1

0.3

0.6

bal

0.2

Sample C

30.1

0.1

0.3

0.1

0.2

0.6

bal

0.3

FIG. 2. Cracks at the bolt holes were also observed on Sample B.

FIG. 4. Optical photomicrograph showing the presence of nodular graphite across the microstructure, as polished.

FIG. 3. Two through-wall cracks were observed on Sample C. The sample internal surface was covered with reddish layer.

lected for chemical analysis. No appreciable wall thinning was noticed on the three samples. Material deposit identification. The chemical composition

of the shell material was determined using XRF spectrometry and C/S analyzer; the results are summarized in TABLE 1. These materials conform to the chemical requirement for ASTM A515 Grade 60 CS. The chemical compositions of the deposits collected from the sample internal surfaces are listed in TABLE 2. The deposits were composed mainly of iron oxides. Traces of Al, Si, P, Ca, Mn and Zn were also detected. Metallographic examination. Cross sections from different areas of the collected samples were prepared for metallographic examination. Examination of the steel microstructure revealed the presence of a high concentration of graphite nodules, as 40MARCH 2013 | HydrocarbonProcessing.com

shown in FIG. 4. These graphite nodules formed aligned structures, i.e., chain graphitization, in localized areas, as shown in FIG. 5. The SEM/EDX analysis was carried out to confirm the composition of the graphite nodules (FIG. 6). The steel was then nital etched to reveal the microstructure. No cementite was noticed in the microstructure, thus suggesting that a complete transformation of cementite in the steel to graphite and ferrite has occurred. Severe, localized chain graphitization was also observed in several cross sections, as shown in FIG. 7. No creep voids and/ or micro fissures were noticed throughout the microstructure. Parallel, unbranched wedge cracks, oriented perpendicular to the surface, were observed on the cross sections removed from areas near the ruptures (FIG. 8). The cracks nucleated on the internal surfaces of the shell plate and propagated toward the external surface. Also, the crack root walls were covered with dense oxide layers. Some of the cracks have grown through the graphite particles, suggesting that fatigue resistance was adversely influenced by graphitization, as depicted in FIG. 9. Mechanical testing. Three round tensile test samples were machined from each sample. The test was conducted at room temperature on a universal testing machine. The tensile test results are summarized in TABLE 3 and FIG. 10. The three specimens from each sample performed fairly consistently. The results indicated small but definite differences in strength levels among the three different sections of the shell. The steel strength did not meet the minimum tensile requirements for ASTM A515 Grade 60 CS. Also, the hardness test results showed that the hardness of the samples is lower than the minimum hardness requirement (TABLE 4).

Corrosion Control TABLE 3. Tensile test results Yield strength, Tensile strength, Elongation, MPa MPa %

FIG. 5. Optical micrograph showing local concentrations of graphite nodules, i.e., chain graphitization, as polished.

ASTM A515 Grade 60

220

415–550

25

Sample A-1

263

361

41.7

Sample A-2

267

359

40.0

Sample A-3

263

355

40.0

Sample B-1

271

375

37.3

Sample B-2

263

370

37.3

Sample B-3

273

376

38.8

Sample C-1

242

342

41.1

Sample C-2

215

347

39.6

Sample C-3

249

344

40.3

TABLE 4. Rockwell hardness test results Specimen

Rockwell, HRB

ASTM A515 Grade 60

68–84

Sample A

58.5

Sample B

58.6

Sample C

58.5

FIG. 6. Backscattered electron image of the graphite nodules and EDX of the graphite nodules.

Discussion. The original microstructure of the reactor shell’s

steel is composed of pearlite (i.e., a mixture of ferrite and cementite) and ferrite. However, the metallographic examination (as well as the mechanical testing) showed that the steel underwent graphitization, and cementite appeared to have completely decomposed into ferrite and graphite. Random and “chain” graphitization were observed throughout the microstructure. It is obvious that the reactor shell was exposed to temperatures greater than 455°C for extended periods. To control graphitization, the metal temperature must be kept below 427°C, where the graphitization rate is extremely low. Overheating of the reactor shell may be sourced to inefficient thermal insulation provided by the refractory lining. The visual and metallographic examinations indicated that the cracking of the reactor shell was caused by corrosion fatigue (oxidation fatigue) nucleated on the shell internal surface. Oxidation fatigue is characterized by unbranched, wedge cracks oriented perpendicular to the surface, and it often appears as multiple parallel cracks. Indeed, systems that operate cyclically (e.g., an MTBE reactor) and/or subject to rapid startup and shutdown procedures are the most vulnerable to oxidation fa-

FIG. 7. Optical photomicrograph showing a microstructure composed of ferrite and graphite nodules. Chain graphitization was observed throughout the steel microstructure as nital etched.

tigue. Also, the formation of chain graphitization indicated significant reduction in fatigue resistance of the reactor shell material. The oxidation fatigue process involves these steps: • Thermal cycling causes expansion of the base metal, generating high internal stresses in the oxide layer. Accordingly, the oxide layer could not withstand high stresses, and, thus, it began to crack, rendering the metal surface exposed to the environment. • Oxides form on the exposed metal surface at the root of the crack, creating a notch effect. The next thermal cycle generates cracking in the oxide along the notch, thus causing the original crack to deepen. • Repeating the listed steps results in the development of a wedge-shaped crack propagating through the shell, eventually leading to a rupture. Hydrocarbon Processing | MARCH 2013

41

Corrosion Control

FIG. 10. Tensile testing results. FIG. 8. Unbranched wedge cracks oriented perpendicular to the surface and often appearing as multiple parallel cracks, as polished.

shutdowns. If graphitization was only found in localized areas of the reactor shell, then patch repair can be conducted. When replacing the reactor shell, upgrading the shell material with Cr-containing low-alloy steel is recommended. Graphitization can be prevented by using steels containing more than 0.7% Cr. • Inspect the refractory lining of the reactor. Shell temperature rise can be linked to inefficient thermal insulation provided by the refractory lining. A refractory specialist should investigate the performance of the refractory lining inside the reactor. Al Ca Fe Mn O

FIG. 9. The crack root and walls were covered with oxides. The cracking of the graphite nodule is near the crack initiation point, and the crack was transgranular, as nital etched.

Lessons learned. With the final root cause identified in the failure of the MTBR reactor, several recommendations were made: • Monitoring and estimating the shell-metal temperature is critical when predicting graphitization. The metal temperature should be kept within the design requirements. It has been reported that exposure at temperatures below 427°C support an extremely low graphitization rate. • Metallographic techniques should be used to evaluate the condition of the reactor shell. Ideally, the extent of graphitization can be investigated through representative sampling of the reactor for metallographic examination. • Graphitization may be evaluated using metallographic field replication procedures. Remember: Damage may occur in the internal part of the shell wall, making it pointless to conduct the field replica. • Replacement. If the condition of the reactor shell is the same as the submitted samples, then the shell should be replaced. Cracking and ruptures facilitated by graphitization may occur suddenly and unpredictably, leading to emergency 42MARCH 2013 | HydrocarbonProcessing.com

Aluminum Calcium Iron Manganese Oxygen

NOMENCLATURE P Si S Zn

Phosphorus Silicon Sulfur Zinc

LITERATURE CITED ASM Handbook, Vol. 1, Properties and Selection: Irons, Steels, and High Performance Alloys, ASM International, 1993. 2 Foulds J. R., and R. Viswanathan, “Graphitisation of Steels in Elevated-Temperature Service,” Journal of Materials Engineering and Performance, October 2001, pp. 484–492. 3 Azevedo C. R. F. and G. S. Alves, “Failure Analysis of a Heat-Exchanger Serpentine,” Engineering Failure Analysis, December 2005, pp. 193–200. 4 ASTM A515/A515M-03 (Reapproved 2007), “Standard Specification for Pressure Vessel Plate, Carbon Steel, for Intermediate- and Higher-Temperature Service.” 1

ABDULAZIZ AL-MESHARI is a failure analyst at SABIC Technology Centre in Jubail, Saudi Arabia. He holds a PhD in material science and metallurgy from the University of Cambridge (UK) and an MSc degree in corrosion science and engineering from UMIST (UK). He is member of a NACE and ASM member since 2000. SAAD AL-ENAZI is a failure analyst at SABIC Technology Centre-Jubail, Saudi Arabia. He holds an MS degree in manufacturing management and a BSc degree in mechanical engineering from the University of Toledo.

ABDELGADER ABDELGALIL is a computer simulation specialist at SABIC Technology Centre in Jubail, Saudi Arabia. He holds a PhD in solid mechanics from the University of British Columbia, Canada, and MSc degree in solid mechanics from Benghazi University, Libya.

Special Report

Corrosion Control K. RAMESH, Reliance Industries Ltd., Hazira, India

Consider dewpoint corrosion in reactor design In a unique series of events at an Indian petrochemicals plant, an 18-mm-thick shell on a 14-month-old oxyhydrochlorination (OHCL) reactor in a vinyl chloride monomer production unit suddenly leaked, resulting in a total plant breakdown over the course of several days (FIG. 1). The cause of the problem was found to be incorrectly designed external cleats on a hydrocarbon reactor. The leak was due to a puncture in the reactor shell at the bracket support plate (meant for a 5T davit) welded to the shell (FIG. 2). A similar reactor had been in service for 18 years without any problem. Technical data and process description. The technical

specifications for the OHCL reactor are shown in TABLE 1. The function of the OHCL reactor is to make ethylene dichloride (EDC) by exothermic reaction of ethylene, hydrogen chloride (HCl) and oxygen (O2 ) in the presence of a fluidized catalyst [copper (II) chloride and alumina]. Fluidization is induced by feeding air to the bottom of the reactor after filtration and compression. All streams are fed to the reactor after sufficient preheating to avoid HCl dewpoint corrosion in the reactor. The complete section of the OHCL reactor is shown in FIG. 3. Observations. A number of observations were recorded during the leak response: • The plant was running at a normal, 1,053-metric-ton load (design capacity is 1,100 metric tons) when the leak was detected, resulting in an emergency shutdown. • A visual inspection revealed localized corrosion of the shell at the inner diameter. Severe thinning and metal loss were observed in an area 300 mm wide by 75 mm high, with an original thickness of 18 mm (FIG. 4).

FIG. 1. Leak site at the OHCL reactor.

• The rest of the shell sections were found to be normal, without any significant erosion or corrosion (see FIG. 5 and TABLE 2). Metallurgical investigations. A chemical analysis of an alloy, using the optical emissions spectroscopy method, confirmed the alloy composition to be the specified standard of SA 516 grade 70, with respect to the elements analyzed. The carbon equivalent of steels used in pressure vessel construction governs hardenability—i.e., the extent of microstructural damage to the material due to welding. The carbon equivalent, based on alloy composition, was determined to be 0.39, which was within the acceptable limit of TABLE 1. Technical specifications for the reactor Specification

Measurement

Shell material

SA 516 grade 70

Design pressure (max.)

6 kg/cm2g (top)

Design temperature (max.)

288°C

Operating pressure

3.1 kg/cm2g (top)

Operating temperature

226°C (top)/245°C (bottom)

Service

EDC + HCl + C2H2 + O2 + CO2

FIG. 2. A puncture in the reactor shell at the bracket support plate was found to be the source of the leak. Hydrocarbon Processing | MARCH 2013

43

Corrosion Control 0.43 for the specified standard. Also, a tensile test showed satisfactory physical property conditions meeting the requirements of SA 516 grade 70 material. The hardness of bulk materials was found to be within acceptable limits. Micro-hardness measurements revealed the highest values (210 Vickers hardness number) at the fillet weld, which was determined to be normal. An impact test performed at −46°C showed acceptable values. Nondestructive testing by wet fluorescent magnetic particle inspection on both the inner and outer surfaces did not detect any linear indication, and no cracks were observed around the leakage. Ultrasonic flaw detection suggested that the plate was free from significant internal defects. Macrostructural observation discovered a bright band of heat-affected zone (HAZ) beneath the fillet weld, along with dark streaks at the center of the cross-section. The welding was determined to be internally sound and absent of any significant defects. Observation of the microstructure indicated fine-grained banded ferrite and a pearlite structure to the plate material. Welding on the external surface showed a microstructure comprising of a dendritic structure of ferrite and carbides, whereas the HAZ microstructure showed a fine and coarsegrained ferrite and pearlite structure. No significant abnormality in the welding was observed. The microstructure of the inner diameter of the reactor clearly showed the effects of erosion and corrosion. In some places, preferential corrosion of the pearlite phase was observed. The scale formation was observed at a microlevel all over the internal surface. The microstructure of the outer diameter showed no significant degradation or damage. Scanning electron microscopy (SEM) of the inner surface revealed a defined corrosion/ erosion pattern, although the surface was found to be free from cracking (FIG. 6). Conclusion from metallurgical investigations. The steel

alloy shell plate showed acceptable conditions for chemical composition, strength, toughness and microstructure properties. The puncture was located at the fillet weld toe of a bracket—a localized region.

FIG. 3. The complete section of the OHCL reactor.

44MARCH 2013 | HydrocarbonProcessing.com

FIG. 4. A visual inspection revealed severe thinning and metal loss.

Corrosion Control Damage was not observed on side-support fillet welds on the same bracket; nor was it observed on other shell plates where pressure, temperature and fluid conditions remained constant at the same elevation inside the reactor. None of the other fillet welds located at the same elevation outside of the shell showed evidence of material degradation underneath. This suggested that a localized phenomenon was affecting the shell plate, leading to erosion and corrosion of the surface of the reactor. The logical next step was to find the reason for the concentrated metal loss. Detailed calculations estimated the heat transfer and resulting temperature profile in the davit support welded to the OHCL reactor. The basis of the study used a process temperature in the reactor of 220°C, with normal process fluid composition, and an ambient temperature of 30°C. Basis for calculations. In the steady state, some of the heat contained in the reactor effluent gases upstream of the cyclones is lost through the reactor wall to the atmosphere. With a heat balance, the heat loss from the process fluid to the wall equals the heat transmitted through a fouling layer (if present); the heat transmitted through the wall equals the heat transmitted through the insulation (if present), which equals the heat loss to the air, as shown in Eq. 1. Q /A = h (tip – tf ) = h (tff – t) = kiw (t – tio )/x = kw I (to – tI )/xI = ha (tI – ta )

TABLE 2. Thickness measurements for shell sections, mm

(1)

where: Q /A= heat loss per unit area. FIG. 7 shows a sketch of the davit support and reactor wall showing the legend for the different temperatures and distances used in the calculations. The thermal conductivities for carbon steel at 225°C and glass-wool insulation at ambient temperature are shown in Eqs. 2 and 3: kw = 47 W/mK

(2)

and kI = 0.035 W/mK

FIG. 5. Plot showing corroded shell sections.

In this case, when estimating the highest shell temperature for the steady-state condition, no fouling was assumed. Any buildup of fouling acts as insulation and further lowers the shell temperature. Deposits were discovered, but the thickness was unknown. The area around the davit support was insulated to the same thickness as the shell of the reactor (insulation thickness of 100 mm), but the davit support itself was an uninsulated metal finger attached to the shell wall. The davit support acts like an aircooler fin, increasing the surface area available for heat transfer. The total heat loss through the davit support is the sum of the heat transferred directly from the process to the shell wall behind the support, along with the heat transferred to insulated sections of the shell in the surrounding areas, and then transmitted through the wall to the support—the coldest spot. Heat-transfer coefficients. In an air-based reactor, the reactor gas effluent has a significant concentration of nitrogen. For simplification purposes, the formula for the heat-transfer coefficient for nitrogen is also used for the process gas. For convection to and from a vertical plate (e.g., a reactor wall), the heat-transfer coefficient is hc = 0.19 (tp – ti )0.33, where

(3)

Region

Thickness

Region

Thickness

Region

Thickness

A

19.2

M

19

Y

18.2

B

19.2

N

17.2

Z

18.2

C

13.7

O

17.2

A1

18.7

D

13.7

P

17.2

B1

18.7

E

13.7

Q

18.2

C1

18.4

F

17.9

R

18.2

D1

18.4

G

19.7

S

17.9

E1

18.3

H

5.3

T

17.1

F1

18.3

I

5.3

U

17.1

G1

18.8

J

5.3

V

17.1

H1

18.8

K

5.3

W

18.2

I1

18.9

L

17.9

X

18.2

J1

18.9

FIG. 6. This SEM image shows pit holes similar to those found in corrosion patterns. Hydrocarbon Processing | MARCH 2013

45

Corrosion Control

Heat-transfer calculations used in case study Goal: Estimate temperature profile in davit support Sketch of davit support tp ti

Temperatures: tp = Process temperature, = 211°C = 412°F ti = Inside metal temp., °C = outside metal temp., °C tI = Insulation surface temp.,°C (when present) ta = Ambient temperature, = 30°C = 86°F

a

to b

e

tI

to

ta

f c

d

g

Dimensions: a = wall thickness, mm = 18 b = support height, mm = 30 c = support length, mm = 504 d = support width, mm = 454 e = distance to hole, mm = 162 f = hole diameter, mm = 277 g = insulation thickness = 100

FIG. 7. Calculations used for heat-transfer measurements.

Calculations for top davit support as installed • tp = 211°C = 412°F • ta = 30°C = 86°F • Wall thickness = 18 mm • Width of support = 454 mm • Height of support = 30 mm • Insulation thickness on reactor = 100 mm • Length of support = 504 mm • Distance to hole = 162 mm • Diameter of hole = 277 mm • Kwall = 47 W/mK at 225°C • Kins. = 0.035 W/m°C at ambient • Cross-section of shell behind support = 0.01362 m2 • Surface area of support exposed to air = 0.4332 m2 • hc = B3 × dt0.33 Btu/ft2h°F, t = °F • B3h = 0.22 • B3v = 0.19 Estimate temperature profile without hot metal heat input • hr = ε × σ × (Tε4 − Tu4) ÷ (Tε − Tu ) • ε = 0.95, T = °R • Estimate to = 34.5°C = 94°F • tref = 273.16 K • σ = 5.67 × 10–8 W/m2K4 • hr = 6.1382 W/m2K • 1 W/m2 °C = 0.176 Btu/ft2h °F • 1.73 × 10–9 Btu/hft2R4 • hc = 2.49 W/m2K • ha = 8.63 W/m2K • Q o = 16.8 W • ti (°C) = 38.7447°C • (Q /A)a = 38.8 W/m2, due to larger surface • For inside, hc = 7.17 W/m2K • Q i = 16.8 W • Iterate estimate to until Q o = Q i 46MARCH 2013 | HydrocarbonProcessing.com

• (Q/A)p = 1234.78 W/m2 through small cross-section • Neglecting any heat input from nearby hot metal, the shell behind the support is 38.7°C Estimate hot shell temperature for insulated sections of reactor • Estimate tl , 36.2°C = 97°F • hr = 6.1892 W/m2K • Radiant heat transfer is much higher than conductive to atmosphere • hc = 2.77 W/m2K • ha = 8.96 W/m2K • (Q /A)a = 55.3 W/m2 • to = 194°C • ti = 194°C • For inside, hc = 3.31665 W/m2K • (Q /A)p = 55.3 W/m2 • Iterate estimate tl until (Q /A)a = (Q /A)p Try to estimate heat transfer from hot shell wall to cold davit support • Areas for heat transfer from various sources, m2: 0 From process gas = 0.01362 m2 0 From side metal = 0.00054 m2 0 From top/bottom = 0.008172 m2 • Add heat duties from all sources and check resulting metal temperature to atmosphere • Estimate to = 102.545°C • Q p = 10.5887 W • Q s = 16.7241 W • Distance to hole = 278.588 mm • Q t = 218.291 W • Distance to hole = 162 mm • Q b = 218.291 W • Total duty = 463.90 W • Iterate to until Q total = Q a • hr = 8.5223 W/m2K • hc = 6.24 W/m2K • ha = 14.8 W/m2K • Q o = 463.891 W • ti = 110.5°C • In still air (no rain), the inside shell wall temperature is around dewpoint • In rain, much more heat can be removed, lowering the wall temperature to below dewpoint Calculations for thinner (20-mm) top davit support • tp = 211°C = 412°F • ta = 30°C = 86°F • Wall thickness = 18 mm • Width of support = 454 mm • Height of support = 20 mm • Insulation thickness on reactor = 100 mm

Corrosion Control

• • • • • • • • • •

Length of support = 504 mm Distance to hole = 162 mm Diameter of hole = 277 mm Kwall = 47 W/mK at 225°C Kins. = 0.035 W/m°C at ambient Cross-section of shell behind support, 0.00908 m2 hc = B3 × dt0.33 Btu/ft2h°F, t = °F B3h = 0.22 Surface area of support exposed to air = 0.4012 m2 B3v = 0.19

Estimate temperature profile without hot metal heat input • hr = ε × σ × (Tε4 − Tu4) ÷ (Tε − Tu ) • ε = 0.95, T = °R • σ = 5.67 × 10–8 W/m2K4 • Estimate to = 33.4°C = 92°F • tref = 273.16 K • hr = 6.1037 W/m2K • 1 W/m2 °C = 0.176 Btu/ft2h°F • 1.73 × 10–9 Btu/hft2R4 • hc = 2.26 W/m2K • ha = 8.37 W/m2K • Q o = 11.3 W • ti (°C) = 37.6281°C • (Q /A)a = 28.1 W/m2, due to larger surface • For inside, hc = 7.18 W/m2K • Q i = 11.3 W • Iterate estimate to until Q o = Q i • (Q /A)p = 1,245.44 W/m2 through small cross-section • Neglecting any heat input from nearby hot metal, the shell behind the support is 37.6°C Estimate hot shell temperature for insulated sections of reactor • Estimate tl = 36.2°C = 97°F • hr = 6.1892 W/m2K h has units of Btu/ft2h°F and t has units of °F. For convection from a horizontal plate (e.g., a davit support), the heat-transfer coefficient is hc = 0.22 (to – ta )0.33, where h has units of Btu/ ft2h°F and t has units of °F. For the heat transfer to air, radiation is also involved. The heat-transfer coefficient ha is the sum of the convection and radiation heat-transfer coefficients; i.e., ha = hc + hr . The radiation heat-transfer coefficient is hr = ε σ (to4 – ta4 – 8) ÷ (to – ta ), where ε = 0.95 and σ = 5.67 × 10 W/m2K4 – 9, or 1.73 × 10 Btu/h ft2R4. The emissivity, ε, in chemical plants is usually taken to be 0.90–0.95 because clean and bright surfaces are rapidly dulled by corrosion, chemicals, etc. In this case, 0.95 has been assumed; σ is the Stefan-Boltzmann constant. Calculations for heat transfer. While the area of the shell wall behind the davit support is well-defined for use in calculating the heat transmitted to and through the reactor wall, the top davit support is an irregular horizontal piece of metal. To simplify calculations, the support is assumed to be a rectangular piece with dimensions of 454 mm by 504 mm, with a 277-mm-diameter hole in it (see sketch in FIG. 7). The vertical

• Radiant heat transfer is much higher than conductive to atmosphere • hc = 2.77 W/m2K • ha = 8.96 W/m2K • (Q /A)a = 55.3 W/m2 • to = 194°C • ti = 194°C • For inside, hc = 3.31665 W/m2K • (Q /A)p = 55.3 W/m2 • Iterate estimate tl until (Q /A)a = (Q /A)p Estimate heat transfer from hot shell wall to cold davit support • Areas for heat transfer from various sources, m2: 0 From process gas = 0.00908 m2 0 From side metal = 0.00036 m2 0 From top/bottom = 0.008172 m2 • Add heat duties from all sources and check resulting metal temperature to atmosphere • Estimate to = 104.91°C • Q p = 6.91997 W • Q s = 10.8621 W • Distance to hole = 278.588 mm • Q t = 215.333 W • Distance to hole = 160 mm • Q b = 215.333 W • Total duty = 448.45 W • Iterate to until Q total = Q a • hr = 8.6174 W/m2K • hc = 6.31 W/m2K • ha = 14.9 W/m2K • Q o = 448.455 W • ti = 114.662°C • For the thinner support, the inside shell wall temperature is 4.2°C higher • Therefore, the thicker top support shell wall temperature will reach dewpoint first sections of the support are ignored, as is the surface area of the davit itself, which represents a much larger area for heat transfer to the atmosphere. For heat transfer through the support, a representative distance is also needed. The distance to the hole is taken to represent an average distance for heat transfer through the support to the atmosphere. Using these assumptions, the calculation results are expected to be conservative; i.e., they will show a higher inside shell wall temperature. The first calculation step is to estimate the temperature profile when considering only heat input from the process and neglecting heat input from the adjacent, insulated hot shell. Since the heat-transfer coefficients are related to the temperature difference, the temperature on the outside of the support must be estimated, and then the heat transfer to the atmosphere is calculated. The heat transfer to the atmosphere is then compared to the calculated heat transfer from the process to the reactor shell behind the support. When both heat quantities match, the estimated temperature profile is calculated. Hydrocarbon Processing | MARCH 2013

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Corrosion Control During this procedure, the sum of all heat inputs flowing through the support is compared to the sum of the heat loss to the atmosphere from the support. At an outside wall temperature of 106°C, at the selected representative distance, there is a heat balance. The resulting inside shell wall temperature is calculated to be 114°C. Once the corrosion process starts, corrosion To investigate why only the top davit support horizontal section (the thickest metal plate) was corroded, products building up on the metal surface the same calculation was made using a 20-mm-thick may act like a fouling or insulating layer. davit support horizontal section. Calculations with the thinner support showed that the estimated inside shell wall temperature was about 118°C, which was about 4°C higher than with the thicker plate. surface area available for heat transfer reduces the temperature Therefore, for any given reactor operating conditions, the difference for transmitting a specific amount of heat, compared thicker metal plate will reach the dewpoint before the thinner to the surface area of the shell behind the support. plate. A simulation is then made to check the expected dewThe second calculation step is to estimate the temperature of point temperature, cooling the reactor effluent composition to the insulated hot shell that surrounds the support. A similar prothe dewpoint at the operating pressure of 3.9 kg/cm2g. The recedure is followed, this time with another layer (the insulation) through which heat passes. The temperature of the insulated hot sulting dewpoint temperature is 109°C. shell is estimated to be 203°C. It is possible that the heat-transfer coefficient for rain falling The third calculation step is to estimate the total heat transfer on the bracket is 5 to 50 times that of the radiant, plus convecfrom all sources, from the support to the atmosphere. Besides heat tion transfer to still air. Additional calculations are performed transfer from the process to the support via the shell, heat comes to define the range of the inside metal temperatures that will refrom the hot shell from the sides, the top and the bottom of the sult. At five times the heat transfer of still air, the rain decreases support. The cross-sectional areas for each source differ due to gethe calculated inside metal temperature to approximately 71°C. ometry, and the distance from the sides to the hole is greater than At 50 times the heat transfer of still air, the calculated inside the distance from the process to the center of the top and bottom. metal temperature drops to approximately 49°C. As shown in FIG. 7, the inside temperature of the shell is about 39°C in this case. One interesting result is that, for heat transfer to the atmosphere from the bracket, the radiant contribution is much higher than the conductive contribution. Also, the larger

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Discussion of results. The average inside shell wall temperature for the whole horizontal davit support was calculated to be very close to the estimated dewpoint. Therefore, it was likely that, during some combination of actual operating conditions, the coldest spot behind the davit support bracket (the center) had reached the dewpoint and initiated condensation from the process gas. The process gas contained a significant concentration of water (1 mole per mole of EDC produced) and traces of unreacted HCl. Since HCl is easily absorbed in liquid water, the drops of water condensing at the cold spot will have a high HCl concentration, making them very corrosive to carbon steel. During periods of rainfall, the heat-transfer coefficient due to the rain is at least five times higher than the natural convection to air. Therefore, the metal temperature is calculated to drop significantly below the dewpoint. This will initiate the corrosion process if it does not begin during normal operation. Once the corrosion process starts, corrosion products (such as iron oxide) building up on the metal surface may act like a fouling or insulating layer, which will also lower the metal temperature behind the fouling. Catalyst fines in the process gas are present in this area of the reactor. It is possible that catalyst particles come into contact with the condensed droplets and become sticky and adhere to the metal surface. This wet catalyst can become a nucleus for further corrosion by absorbing water/HCl, while also acting as an insulating layer. This interpretation seemed to be supported by the energy dispersive spectroscopy analysis. The catalyst was copper chloride supported on an aluminium oxide base, while corrosion products were largely metal oxides, with some chlorides.

Corrosion Control The black deposit, which was furthest away from the hole, showed some aluminium and some chloride. It also showed some catalyst deposit on the cyclones and coils, which appeared to be a normal occurrence in most parts of the reactor. The red deposit, which was below the hole, gave a high concentration of oxygen, likely from the transportation of corrosion products with the condensate and their deposition when the condensate vaporized again. The catalyst quantity in this area was not much different from the black deposit, since the aluminium content was similar. The yellow deposit, closest to the hole, seemed to have a higher catalyst content, since the aluminium content was more than double that of the other samples. It also had high oxygen content, indicating corrosion products in this area as well. The shape of the corroded area, which mirrored the horizontal section of the support, was likely due to condensed droplets running down to a hot section of the shell, then vaporizing again, so that only the cold metal behind the support was affected. The coldest part of the horizontal section was the center, which was the furthest distance from the adjacent insulated hot sections. The corrosion was likely to have started here and then spread uniformly toward the edges of the cold region. Additionally, any signs of erosion can be explained by the action of the catalyst fines in the gas acting like sand blasting, once a hole in the shell was created and process gas began flowing out as a result of the pressure difference. The thermal calculations support the theory that a cold spot in the top davit support initiated dewpoint corrosion, which

led to the hole in the shell. Regarding possible erosion from the process gas in normal operation, the two sets of cyclones were symmetrically arranged, with inlets 180° from each other. Any set of flow conditions was expected to affect both sides of the reactor equally, while the shell damage was restricted to one location. Since it was not expected that the reactor operated outside of normal fluid velocities, dewpoint corrosion was likely the cause of the hole in the shell. Having verified the inadequacies of design and the root cause, the davit supports were shaved off and the shell was completely insulated. No additional problems were encountered in the reactor. Recommendations. To avoid cold spots and dewpoint cor-

rosion in such a reactor service, it is not recommended to have large, uninsulated parts welded to the shell of the reactor. Conversely, if these parts are present, they should be properly engineered at the design stage to avoid potential cold-spot issues in service. K. RAMESH is the vice president of mechanical and central engineering services at Reliance Industries Ltd. in Hazira, India. He has over 20 years of experience in all aspects of static equipment inspection, design, nondestructive testing, failure analysis, material selection, quality assurance, corrosion monitoring and control, and welding inspection. Mr. Ramesh holds a master’s degree (MTech) from the Indian Institute of Technology in Kharagpur, India. He is a certified inspector for API-510, API-570, API-653, AWS-CWI and ASNT Level-II RT/PT/MT/UT.

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Special Report

Corrosion Control T. HANRATTY, The Sherwin-Williams Co., Houston, Texas

Corrosion under insulation is a hidden problem Corrosion under insulation (CUI) is one of the costliest avoidable problems facing the hydrocarbon processing industry (HPI) today. CUI afflicts refineries—specifically, the steel piping, storage tanks, container vessels and other process equipment within the plants that are subject to extreme temperature fluctuations. Calcium silicate or mineral-wool insulation applied to the pipe or vessel can mitigate the thermal cycling effects. But the presence of seams, gaps or other discontinuities in the insulation layer makes them susceptible to infiltration by outside moisture or from the process environment itself. Understanding the mechanisms. CUI originates when water and contaminants infiltrate an insulated system that has certain water retention, permeability and wetability characteristics. Water sources include rainfall, cooling tower drift, steam discharge, wash downs and (because insulation is not vapor tight) condensation. Water may enter the system due to breaks in the waterproofing, inadequate system design, incorrect installation, poor maintenance practices or a combination of such factors. Once wet, the insulation system’s weather barriers and sealants trap the water inside, so the insulation remains moist. Next to the equipment surface, the insulation forms an annular space or crevice that retains the water and other corrosive media. For example, chlorides and sulfates that may be native to the insulation can accelerate the corrosion process. Substrates of either carbon steel (CS) or austenitic and duplex stainless steel (SS) are susceptible to CUI. In CS, CUI occurs in piping or equipment with a skin temperature in the range of 25°F to 350°F (-4°C to 175°C), where the metal is exposed to moisture over time under any kind of insulation. The rate of corrosion

varies with the specific contaminants in the moisture and the temperature of the steel surface. Waterborne chlorides and sulfates concentrate on the CS surface as the water evaporates. In austenitic and duplex SS, external stress corrosion cracking (ESCC) can occur, but the temperature threshold is higher, from 120°F to 350° F (50°C to 175°C). For ESCC to develop, sufficient tensile strength must be present. Here again, waterborne chlorides concentrate on the SS’s hot surface as water evaporates. Why CUI is a hidden problem. Once infiltration occurs, insulation and cladding conceal the progress of degradation to piping and equipment. Even with observation ports, less than 1% of the surface is visible, and those areas generally are not representative of the whole unit. Removal of just some insulation to complete a minor repair job can lead to the discovery of a degree of corrosion that is an unpleasant surprise and may require a facility shutdown to rectify. Typically, insulation is removed on a 15–20 year cycle, which makes diagnosing a problem in a timely manner less likely. Yet, because in existing facilities

pipes and vessels may have been put into service when coatings, insulation and refinery operating conditions were very different than they are today, it is all the more critical to promptly identify deterioration lurking on these substrates. The evolution of the CUI problem goes back to the 1970s. Before that time, little if any thermal insulation was applied to heated CS equipment and vessels below 300°F (150°C). But as energy costs increased, it became more cost-effective to apply insulation. Concurrently, newer processes came onstream, operating at higher and often cyclic temperatures while austenitic SS pipe and equipment became more common as well. Together, these developments dramatically increased the amount of insulation used in the HPI and set the stage for CUI to become a pervasive issue. Chronic CUI has also become a problem for reliability engineers. Major equipment outages, whether for periodic inspection and maintenance or due to a catastrophic failure, account for more operational disruptions than any other cause. Understandably, all of the major HPI companies became aware of and were concerned about the problem, with

Topcoating recommended Although it is common in the petrochemical and refining industries to use a shop-applied inorganic zinc (IOZ) coating as a primer on new CS piping (because it dries quickly and is cost effective), IOZ provides inadequate corrosion resistance in closed, sometimes wet, environments. At temperatures greater than 140°F (60°C), the zinc may undergo a galvanic reversal, where the zinc becomes cathodic to the CS. Shop-primed pipe will be finish-coated at the jobsite, depending on the service conditions needed. The NACE standard recommends topcoating the IOZ to extend its service life, and that it not be used by itself under thermal insulation in service temperatures up to 350°F (177°C) for long-term or cyclic service. In cases where pipe is previously primed with an IOZ coating, it should be topcoated to extend its life. Hydrocarbon Processing | MARCH 2013

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Corrosion Control one taking the lead in thoroughly analyzing the mechanisms of corrosion and developing test protocols for protective coatings to deal with it.

So, while in the past CUI was once largely a maintenance issue in existing refineries, the dominance of offshore new construction and conversions demanded new strategies for asset protection. Owners expect their vessels to have decadeslong life cycles without dry docking. And specification of coatings has become a high-level decision, given what is at stake.

205°C); and inert multipolymeric matrix coatings are best for temperatures from –50°F to 1,200°F (–45°C to 650°C). Coatings suitable for refinery applications should perform well in testing protocols based on both accelerated and real Protective coatings. Though valiant world scenarios involving typical CUI efforts to keep water out of insulated mechanisms. Chief among these is the systems can be made using different boiling water test, which is the gold standesign materials and configurations of dard for accelerated testing of heat-resistant coatings. To ensure a coating can offer superior resistance to thermal cyOnce infiltration occurs, insulation and cladding conceal cling, a steel panel is subjected the progress of degradation to piping and equipment. to thermal shock in a simulated immersion scenario. The Even with observation ports, less than 10% of the surface best results show no adhesion is visible. Removal of just some insulation to complete or blistering after 80 cycles of a minor repair job can lead to the discovery of a degree the test when applied at ambient temperatures. of corrosion that is an unpleasant surprise and may But these coatings have to require a facility shutdown to rectify. endure more than just corrosion. Manufacturers of second-generation CUI coatings recognize that products Given such customer demand, the must address the challenges stemming the equipment to be insulated, CUI is not usually kept at bay on the strength protective coatings industry has devel- from new construction. Now coatings of those measures alone. Industry guid- oped products that can accommodate must offer enhanced shop coating and ance, provided by NACE International, not only extended life cycle expectations, throughput properties; durability during The Corrosion Society, holds that im- but also the dramatic increase in steel fab- transportation and erection of fabricated mersion-grade protective coatings are rication work to produce the equipment modules; and worker friendliness in a shop environment. the best defense against CUI in both in- and vessels. This means more flexible coatings, sulated CS and austenitic and duplex SS. Insulated steel capable of trapping water Research and development is key. with harder films, that deliver measurable is considered to be under immersion at Ten years ago, the typical process operat- reductions in dry-to-recoat and dry-to210°F (99°C) or higher. ing temperatures of insulated equipment touch times; better impact resistance to NACE Standard SP0198-2010, “The were lower than they are today. Modern limit damage in transport from the shop Control of Corrosion Under Thermal facilities operate at temperatures as high to the jobsite; and lower volatile organic Insulation and Fireproofing Materials— as 400°F (205°C), where 300°F (150°C) compound emissions to promote worker A Systems Approach,” reflects the latest was more the norm previously. Although safety. With these attributes come a reinsights in CUI prevention and mitiga- most equipment doesn’t run at the high duced total cost of ownership and extendtion from the oil and gas industry, includ- end of the temperature design, spikes can ed service life for high heat applications ing the products and systems available to occur for various reasons and must be and equipment under insulation. Coatings that demonstrate these addicombat CUI that have a track record of taken into account when specifying the success. And many of the insights stem appropriate coating system. Thus, a cru- tional attributes will perform well in rigfrom the shift of oil and gas exploration, cial consideration when determining the orous testing protocols based on ASTM extraction and processing to offshore appropriate protective coating system to standards for abrasion and impact resisdeepsea locations around the world. use under insulation is the expected ser- tance, such as the falling sand, Taber abraThis shift has numerous implications vice temperature of the equipment or pip- sion and pencil hardness tests. These tests for those entrusted with protecting float- ing, especially when intermittent thermal are done under similar ambient temperaing production, storage and offloading cycling—from hot to ambient or hot to ture conditions to the ones used in the boiling water test. (FPSO) assets from corrosion, in general, less hot temperatures—is present. and CUI, in particular. The marine enviPresent commercially available coatings ronment combined with extreme process are engineered to perform at various tem- TIM HANRATTY is petrochemical business temperature fluctuations can accelerate perature ranges because one size does not development manager and a corrosion specification the corrosion rate of process equipment fit all. Phenolic epoxies are for tempera- specialist for The Sherwin-Williams Co. in Houston, He is a NACE-certified CIP Level 1, Level 2 and and CS process piping and vessels under tures of –50°F to 300°F (–45°C to 150°C); Texas. Level 3 Peer Review Coating Inspector. Mr. Hanratty insulation by a significant factor com- Novolac epoxies should be applied for was a member of the committee that developed NACE pared to onshore assets. temperatures of –50°F to 400°F (–45°C to Standard SP0198-2010. 52MARCH 2013 | HydrocarbonProcessing.com

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| Bonus Report SAFETY DEVELOPMENTS Safety programs are part of best practices implemented by hydrocarbon processing facilities and are central to improving plant and company profitability. More importantly, safety programs must be able to adapt to the needs of each facility, and they require support from top management down to factoryfloor personnel. Engineers work with an operator training simulator. See full story in the February 2013 Hydrocarbon Processing article, “Operator training simulators for brownfield units offer many benefits,” pp. 45–47. Photo courtesy of Honeywell Process Solutions.

Bonus Report

Safety Developments M. A. TURK and A. MISHRA, Invensys Operations Management, Houston, Texas

Process safety management: Going beyond functional safety Modern hydrocarbon processing facilities have become increasingly more complex. Likewise, the risks in managing greater capacity refineries and petrochemical complexes have increased. Ensuring the safety of employees, the environment and physical plant assets in the event of an unexpected process excursion cannot be overstated. The development of new techniques and technologies designed to improve operational safety has evolved to meet these challenges. Operating companies are increasing efforts to reduce the risk of catastrophic events such as the release of toxic, reactive or explosive chemicals that can damage the environment or plant assets, as well as, cause injury or death to employees and the general public.

ROAD TO IMPROVING PLANT SAFETY This journey begins with the development of the modern process safety management (PSM) systems and requirements. Efforts to improve plant safety were led by state-of-the-art functional safety systems. These systems enable the orderly shutdown of processing units when abnormal situations occur that are beyond the capabilities of the regulatory control system or operators to correct or to prevent a catastrophe. While functional safety has proven successful in reducing the probability of catastrophic events and recognizes the role of human factors, it does not explicitly address the key roles of management and business processes in maintaining operational integrity and profitable performance of process plants. In this context, what are the approaches that operating companies should take to go beyond functional safety to proactively measure, monitor and display a plant’s risk profile in near real time so that proper actions can be taken in a more timely manner to improve process safety performance? Why invest time and resources to go beyond the limitations of functional safety? To answer this question, we must discuss the pivotal concepts of safety-performance indicators and values (plant assets, the environment, the public and employees) at risk from potential catastrophic events. What are the best practices for establishing a PSM culture along with designing, implementing and maintaining a proactive PSM system to complement existing functional safety systems? HISTORY OF SAFETY MANAGEMENT SYSTEMS As industrialization and technology progressed in the early 20th century, the pattern of intermittent catastrophes began. In 1921, at the BASF plant in Oppau, Germany, explosions

destroyed the plant, killing at least 430 people and damaging approximately 700 houses nearby. This explosion occurred as blasting powder was used to breakup the storage pile of a 50/50 mixture of ammonium sulfate and ammonium nitrate. This procedure had previously been used 16,000 times without any mishap. In 1947, a fire and explosion in Texas City, Texas, on the Monsanto Chemical Co.’s S.S. Grandcamp while loading ammonium nitrate fertilizer killed over 430 people. There was no specific legislative response to these incidents.1 Interestingly, the US Center for Chemical Process Safety (CCPS), which provides leadership and infrastructure to promote and advance PSM, suggests process safety was born on the banks of the Brandywine River in the early days of the 19th century at E. I. du Pont’s black powder works. Recognizing that even a small incident could precipitate considerable damage and loss of life, du Pont directed the works to be built and operated under very specific safety conditions.2 Industry has a short memory; here is a brief list of several recent major industrial disasters with dire consequences: • 1984—Bhopal, India. A toxic material released caused 2,500 immediate fatalities and many other offsite injuries over time. • 1984—Mexico City, Mexico. An LPG explosion caused 300 fatalities (mostly offsite) along with $20 million in damages. • 1988—Norco, Louisiana. A hydrocarbon-vapor-cloud explosion resulted in seven onsite fatalities and 42 injuries, as well as over $400 million in damages. • 1989—Pasadena, Texas. An ethylene/isobutene explosion and fire caused 23 fatalities, 130 injuries and more than $800 million in damages. Such catastrophic safety incidents damaged the public and the environment. They also caused significant economic loss. In response, governments continue to enact legislation and impose fines focused on reducing the probability of future events. Likewise, operating companies formed safety-related consortiums that include suppliers of process automation technology. The goal is to identify automation solutions that can enable operating companies to avoid catastrophic safety events through early detection and correction. As evidenced by recent safety-related catastrophes, such solutions have not been entirely successful. The present state-of-the-art safety management includes safety studies (HAZID, HAZOP, risk analysis), safety instrumented systems (SISs) for fire and gas detection, and emergency shutdown, abnormal situation management applications, and operator guidance tools. As illustrated in FIG. 1, the first Hydrocarbon Processing | MARCH 2013

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Safety Developments step in implementing a functional safety system is the upfront analysis and conceptual design. It begins with a meeting with all stakeholders to determine possible hazards and hazard characteristics, and to establish the basic scope of the project. Work then proceeds to develop the detailed design for the SIS. The next steps involve: • Executing the process hazard analysis (PHA) and layers of protection analysis (LOPA) • Specifying the safety instrumented functions (SIFs) and preparing the safety requirements specification (SRS) reports • Developing the safety integrity level (SIL) verification worksheet and report. While these approaches to safety management have produced positive results in reducing the probability of potentially dangerous process upsets or failures, they are either static (e.g., HAZOP studies) or reactive (e.g., emergency shutdown systems) in nature. Their performance is also hampered by complacency. Time passing without an incident is not necessarily an indication that all is well. There is always a succession of failings that lead to an incident, as shown by the Swiss-cheese model (FIG. 2). If unchecked, all systems will deteriorate over time, and major incidents can occur when defects cross a number of risk-control systems concurrently. In effect, the “holes” in the Swiss-cheese model become larger. Without setting leading and lagging indicators for each risk-critical control system, it is unlikely that failings in these barriers will be revealed as they arise before all of the important barriers are defeated. Numerous recent high-profile incidents have heightened the awareness that organizations need to pay more attention to process safety. By definition, process safety is a blending of engineering and management skills focused on preventing catastrophic accidents and near hits—particularly, explosions, SIS safety lifecycle – Upfront activity (analysis, conceptual) Develop conceptual process design basic project scope

Typical process history Possible hazards Hazards characteristics P&ID (rough) Hazards classification Consequence matrix Tolerable frequency Initiating events Enable factors Probability of exposure Failure probabilities IPL credits Risk reductions Corporate SIS standards International standards Prescriptive subsystems SIS operation requirements SIS maintenance requirements P&ID (semifinal)

Identify PHA potential SIF

Basic project scope Process characteristics Process details Process flow diagram

PHA report

Execute consequence analysis LOPA

LOPA report

Develop safety function requirements

FDS, SIF matrix SRS Safety lifecycle chart SIF dossier

fires and damaging releases associated with the loss of containment of energy or dangerous substances such as chemicals and petroleum products. These engineering and management skills exceed those required for managing the workplace. As industrial infrastructures continue to age, the consequences of applying process safety incorrectly increases with escalating consequences, such as: • Damage to people, the community and environment • Litigation against corporations and individuals • Increased scrutiny by regulators and governments • Undermined investor confidence with resulting loss in stock price. In some cases, even when executives and managers have prioritized process safety, things still go wrong. Too often, organizations or individuals make process-safety decisions under pressure, or without proper context or sufficient information. What’s missing is the ability to provide plant personnel with real-time, proactive actionable information about the plant’s risk profile via continuous measurement, monitoring and visualization of key operating and safety-related parameters. Result: Potentially hazardous events can be averted without resorting to a plant trip or an emergency shutdown. This is the goal of PSM; it involves next-generation automation solutions aimed at making step-change improvements in safety performance. Such systems can provide a “safety early warning and hazard avoidance system.” This should be an essential component of the modern hydrocarbon enterprise. By way of definition, PSM is the application of management systems to identify, understand and control process hazards, thus preventing process-related injuries and incidents.3 The goal is to minimize process incidents by evaluating the whole process. PSM came into widespread use after the adoption of OSHA Standard 29 CFR 1910.119 Process Safety Management of Highly Hazardous Chemicals in 1992. PSM covers: • Process safety information • Employee involvement • PHAs • Operating procedures • Training • Contractors • Pre-startup safety reviews • Mechanical integrity • Hot work • Management of change Process design Other

Hazard

DCS Develop SIS conceptual design

Develop non-SIS conceptual design

Perform SIL verification

SIS detail design Non-SIS detail design

SIF, Non-SIF functions C&E charts SIS interlocks requirements SIS hardware concept SIS operator interface

Accident SIL verification report SIL verification worksheet

FIG. 1. Steps in the FEED of a safety instrumented system.

56MARCH 2013 | HydrocarbonProcessing.com

SIS

FIG. 2. Swiss-cheese model of how a hazard can propagate and become a harmful event.

Safety Developments used to show plant availability and optimizes operating condi• Incident investigation tions. Effective management of major hazards requires a proac• Emergency planning and response tive approach to risk management. Information to confirm that • Compliance audits critical systems are operating as intended is essential. Leading • Trade secrets. indicators that can confirm that risk controls are contining to Another definition of PSM is “the proactive and systemoperate is an important step forward in the management of maatic identification, evaluation and mitigation or prevention of jor hazard risks. chemical releases that could occur as a result of failures in processes, procedures or equipment.” 4 PSM is intended to ensure freedom from unacceptable risk due to: • Fire Since the advent of the modern • Explosion • Suffocation hydrocarbon age, petroleum refining and • Poisoning. petrochemical process operations have FIG. 3 shows where PSM fits into the overall conbecome increasingly complex and potentially text of operational integrity (i.e., keeping the process in the pipe), and how functional safety is a key elevery dangerous if not managed correctly. ment of PSM. Business case for PSM. A cost/benefit analysis

is at the center of decision-making on investments. To justify cost, it is necessary to determine if the magnitude of the value delivered justifies the cost in terms of time, effort and money. Investments in safety—functional safety systems, abnormal situation management applications, etc.—have been made largely to satisfy legislative requirements and to maintain the license to operate. There is no legislation that directly defines the requirements for a real-time PSM system or the penalties for not implementing one. Thus, investments in a PSM system may be made if it can be shown that it delivers a significant, tangible reduction in the risk of a catastrophic failure, and that it produces a measurable economic benefit for the plant. TABLE 1 summarizes estimated annual benefits associated with implementing a PSM system. For a 100,000-bpd petroleum refinery, operating for 330 days/yr at an average refining margin of $5/ bbl, the estimated annual PSM benefit is $2.85 million. In addition to the stated benefits from TABLE 1, the “incremental value-at-risk” can provide ongoing quantified measures of the economic impact from the PSM system.

DESIGN AND FRAMEWORK FOR A PSM SYSTEM It is important to find the right level of balance among the various possible safety indicators so that process-safety decisions accurately reflect the company’s desired operational risk profile. Although risk can never be eliminated, a variety of mechanisms can be put in place to balance desired safety outcomes with day-to-day business imperatives and pressures. Too often, many organizations rely heavily on failure data to monitor performance. Thus, improvements or changes are only determined after something has gone wrong. Often, the difference between whether a system failure results in a minor or catastrophic outcome is purely down to chance. The consequence of this approach is that improvements or changes are only determined after something has gone wrong. Discovering weaknesses in the quality of managing the process and control systems by having a major incident is too late and costly. Early warning of dangerous deterioration within critical systems provides an opportunity to avoid major incidents. Knowing that process risks are successfully controlled has a clear link with business efficiency, as several indicators can be

Measuring performance. The main reason for measuring process safety performance is to provide ongoing assurance that risks are being adequately controlled. Directors and senior managers need to monitor the effectiveness of internal controls against business risks. For petroleum refineries and petrochemical manufacturers, process safety risks are a significant aspect of business risk, asset integrity and reputation. Many organizations lack good information to show how well they are managing major hazard risks. This is because the information gathered tends to be limited to measuring failures, such as incidents or near misses. Those involved in managing process safety risks need to ask fundamental questions about their systems, such as: • What can go wrong? • What controls are in place to prevent major incidents? • What does each control deliver in terms of a “safety outcome”? • How do we know that the controls continue to operate as intended? Measuring performance before a catastrophic failure. According to James Reason, (major) accidents result when a series of failings within several critical risk-control systems materialize concurrently.5 Each risk-control system represents an important barrier or safeguard within the PSM system. A

Operational integrity Process safety đƫ!+,(! đƫ.+!//!/ đƫ-1%,)!*0ĥ/5/0!)/ Functional safety đ đ  đ
(.)/

Occupational safety đ.%,/ đ(%,/ đ((/

“Keeping the process in the pipe” FIG. 3. Role of PSM in supporting operational integrity. Hydrocarbon Processing | MARCH 2013

57

Safety Developments significant failing in just one critical barrier may be sufficient to give rise to a major accident. Continuously measuring and monitoring the actual real-time performance of these safety barriers ensures that operational integrity is not compromised due to degradation of barriers. Leading and lagging indicators are set in a structured and systematic way for each critical risk-control system within the whole PSM system. In tandem, they act as system guardians, providing dual assurance to confirm that the risk-control system is operating as intended or providing a warning that problems are starting to develop. Leading indicators are an active monitoring form focused on a few critical risk-control systems to ensure continued effectiveness. Leading indicators require a routine systematic check that key actions or activities are undertaken as intended. They can be considered as measures of process or inputs essential to deliver the desired safety outcome. The leading indicators identify failings or “holes” in vital aspects discovered during routine checks on the operation of a critical activity within the riskcontrol system.

Lagging indicators are reactive monitoring methods requiring the reporting and investigation of specific incidents and events to discover weaknesses within that system. These incidents or events do not have to result in major damage or injury or even loss of containment, providing they represent a failure of a significant control system that guards against or limits the consequences of a major incident. Lagging indicators show when a desired safety outcome has failed or has not been achieved. The lagging indicator reveals failings or “holes” in that barrier discovered following an incident or adverse event. The incident does not necessarily have to result in injury or environmental damage, and it can be a near miss, a precursor event or an undesired outcome attributable to a failing in that risk-control system. Several organizations and standards recommend applying leading and lagging metrics to understand the quality of the PSM system. Several examples are: • ISA 84.00.04—Recommended Practices for Guidelines for the Implementation of ANSI/ISA-84.00.01-2004 (IEC 61511 Mod) • CCPS

TABLE 1. Operator training simulator potential benefits Potential PSM benefits Assumptions

Value estimates

Maximum sustainable daily rate (bpd) = 100,000 Margin ($/bbl) = $5 Benefit type

Notes

Potential, $

Annual probability, %

Actual cash, $

$39,860

4%

$1,594

4%

$301

4%

$128

Loss avoidance 1. Catastrophic loss Human life

1 death + 5 serious injuries

$20,000

Cleanup

$660

Compensation to local businesses

$200

Fines Equipment replacement

Loss of earnings

Punitive environmental fine

$1,000

1 heater destroyed Damage to 2 adjacent heaters Damage to reactor Catalyst losses

$4,000

20,000 BPD of gasoline for 3 months

$9,000 $7,513

2. Reputation Loss of sales

5% of 70,000 bpd of fuel sales for 1 year

$6,388

Hiring and retaining staff

3% Loss of total staff of 1,500 employees; Hiring Cost = $25,000/new employee

$1,125 $3,190

3. License to operate Safety

Additional sensors and safeguards; updated HAZOP; revised ESD logic; additonal bunding

Environmental

Automated reporting system

$2,690 $500 $50,563

Loss avoidance subtotal

$2,023

Improved economic performance 0.5% increase in output due to:

Productivity benefits

$825

100%

$825

Shorter startups Fewer unplanned shutdowns Faster grade/throughput changes Better handling of process disturbances Direct cash benefits

58MARCH 2013 | HydrocarbonProcessing.com

$51,388

$2,848

Safety Developments • The Energy Institute (EI), formerly known as the Petroleum Institute. The common theme of these metrics is applying key performance indicators (KPIs) generated from the management of the process/functional safety equipment and the people and processes that are used in terms of their competence, leadership and risk-management capabilities. For example, the EI has published a Process Safety Management framework, developed by the energy industry, for use by various industry sectors.6 The framework is intended to be applicable worldwide, to all process industries such as power, petroleum, chemicals, refining, etc. The framework encapsulates learning from people with practical experience of developing and implementing PSM as part of an integrated management system. It clearly sets out what needs to be done to ensure the integrity of the operation and define what measures should be in place and how they are performing. Note: It is not intended to replace existing process safety or health, safety and environmental (HSE) management systems. The EI’s framework consists of three levels: focus areas, elements and expectations. The focus areas set out the high-level components of the PSM framework. Within each of the focus areas are a number of elements. Each element contains expectations defining what organizations need to do properly to meet the intent of each element. Details for EI’s PSM elements set four key operating aspects that organizations should do to ensure the integrity of the operations: • Process safety leadership 0 Leadership commitment and responsibility 0 Identification and compliance with legislation and industry standards 0 Employee selection, placement, competency and health assurance 0 Workforce involvement 0 Communication with stakeholders • Risk identification and assessment 0 Hazard identification and risk assessment

0 Documentation, records and knowledge management • Risk management 0 Operating manuals and procedures 0 Process and operational status monitoring, and handover 0 Management of operational interfaces 0 Standards and practices 0 Management of change and project management 0 Operational readiness and process startup 0 Emergency preparedness 0 Inspection and maintenance 0 Management of safety-critical devices 0 Work control, permit to work and task risk management 0 Contractor and supplier, selection and management • Review and improvement 0 Incident reporting and investigation 0 Audit, assurance, management review and intervention. FIG. 4 shows the proposed PSM framework—based on industry guidelines—and the associated components of a welldesigned PSM system to enable real-time measurement and monitoring of a plant’s risk profile. It provides actionable inforPSM

+ –

Designed risk

Process plant

Key performance FIG. 5. PSM control loop.

Plant 1

Plant 2

Solution design

Solution implementation

Services

LOE 1 Customer self assessment

Complacency

LOP

LOP

Plant 3

LOE 2 LOP

LOP

LOPs

KPI framework Corporate and site management

Operations, maintenance and engineering

FIG. 6. Asset-owner safety model.

Safety performance indicator calc. engine Industry guidelines

SPI

Integrated information and workflow platform PSM support applications/solutions

FIG. 4. PSM framework and components.

Plant Software

PSM framework

Dashboards leading and lagging KPIs

LOE KPI

Asset risk

LOPs KPIs

Asset risk

FIG. 7. Plant-safety model with KPIs and SPI. Hydrocarbon Processing | MARCH 2013

59

Safety Developments mation that can be used to prevent catastrophic events. Where an organization has an existing HSE or PSM system, it may be useful to benchmark against the framework or to carry out a risk assessment vs. the expectations of each element and identify any aspects of the existing system that may need enhancing. Implementing such a PSM system establishes the foundation of a PSM “control loop.” FIG. 5 illustrates such a control loop to prevent complacency from increasing the probability of a catastrophic event due to plant personnel ignoring leading and lagging indicators about degradation of protection levels provided by risk-control loops. During plant operations, systems are modified to adapt to the changing system needs. Systems and procedures can deteriorate over time, and system failures discovered following a major incident frequently surprise senior managers, who sincerely believed that the controls were functioning as designed. Used

effectively, process safety KPIs can provide an early warning that critical controls have deteriorated to an unacceptable level. Measuring performance to assess how effectively risks are being controlled is an essential part of an HSE system. This can be accomplished in two ways: • Active monitoring. It provides feedback on performance before an accident or incident • Reactive monitoring. It involves identifying and reporting on incidents to check that the controls in place are adequate, to identify weaknesses or gaps in control systems and to learn from mistakes. SPIs and incremental value-at-risk. After a set of KPIs

Percent Percent

60MARCH 2013 | HydrocarbonProcessing.com

Percent

Percent

Percent

have been adopted, the asset owner’s management is responsible for monitoring these KPIs and responding to deviations from their baselines. At higher management levels, the relevance of the KPIs associated with managing plant equipment can be lost. Therefore, it becomes necessary to translate the Corporate individual equipment level KPIs and their business impact into 15.1 49.5 4.95 value at risk, $ million plant-level safety performance indicators and its business impact. This concept can be extended to any number of facilities enabling upper management to understand the quality of PSM Inc. asset Inc. prod. Inc. prod. Plant SPI at risk, at risk, at risk, across the enterprise. $ million thousand bpd $ million Using the individual equipment KPIs, a new approach allows an asset owner to understand the overall safety state of the plant NA-P1 89% 10 30 3 and its economic impact on the business. In addition, this approach is tied to the existing LOPA and financial impact analysis. KPI metrics are gathered based on the asset owner’s manageNA-P2 90% 3 10 1 ment of the plant equipment, capability of employees and processes followed to manage process safety. Typically, 10–20 key metrics can be covered and include 1) management of safety-reEU-P3 96% 0.1 1 0.1 lated equipment (e.g., completion of periodic field-device proof tests associated with a distillation column), 2) competence of plant personnel (e.g., their level of training and skills testing), ME-P4 90% 2 8 0.8 3) adherence to established procedures (e.g., near-miss investigations) and 4) leadership (e.g., involvement of leadership in periodic, formal safety reviews). These metrics can originate AP-P5 100% 0 0.5 0.05 from management based on the layers of protection (LOPs) associated with the different lines of equipment, from at a LOP level (e.g., SIS) or at the line of equipment level (leadership). FIG. 8. Example of a corporate dashboard. The safety performance indicator (SPI) is an aggregation of the individual KPIs into a single number. The SPI can be calculated at the equipSafety performance Inc. rev. Inc. prod. Inc. asset ment level (equipment SPI) and at the indicator at risk at risk at risk Plant value plant level. FIG. 6 illustrates the owner at risk 89% $ 1 million 5,000 $ 10 million safety model for an enterprise’s global assets. This model can consist of plants Op. readiness Leadership Competency distributed over different geographic re100 100 100 80 80 80 gions. A plant is decomposed into lines 60 60 60 of equipment (LOE), which have LOPs 40 40 40 associated with the plant-safety model, as 20 20 20 0 0 0 shown in FIG. 7. Jan Feb Mar Apr MayJuneJuly Aug Sept Oct Nov Dec Jan Feb Mar Apr MayJuneJuly Aug Sept Oct Nov Dec Jan Feb Mar Apr MayJuneJuly Aug Sept Oct Nov Dec Safety device management Incident reporting Underlying the plant-safety model is a 100 100 safety related KPI framework; it addresses 80 80 60 60 the management of process safety related 40 40 to plant equipment, business processes, 20 20 0 0 and procedures used to manage the equipJan Feb Mar Apr MayJuneJuly Aug Sept Oct Nov Dec Jan Feb Mar Apr MayJuneJuly Aug Sept Oct Nov Dec ment and the capabilities of employees FIG. 9. Example of a plant-level dashboard. applying these processes and procedures.

Now Accepting

Challenging Projects Sure we can do the standard tower open/clean/inspect/close work but it’s those tough and challenging jobs that have helped us earn our stripes. We recently completed a revamp on one of the largest vacuum towers in the western hemisphere, much to our customer’s satisfaction. We’ve mastered a resection method that is an excellent and cost saving alternative when footprints are tight. Not long ago we tackled a 114-vessel project at a natural gas processing plant. Our team blinded, opened, cleaned, inspected and repaired all of the vessels in just 14,642 man-hours; less time than had been scheduled and well under budget. When performing any number of services, we don’t overlook anything from external pipe flanges, complicated vessel internals, feed/draw arrangements, section replacements, nozzle and strip lining installation/repair. We strive for “Zero Injury.” As for quality, well, that’s why our customers invite us back again and again and give us their annual maintenance/service contracts. We thrive on accepting challenges then exceeding expectations but we are just as agile with a single tower project as we are with plant-wide turnarounds. We are quick on our feet and can mobilize swiftly for emergencies. Challenge us today and we’ll have your towers productive tomorrow.

3012 Farrell Road, Houston, TX 77073 *…œ˜i\ÊÓn£°™x£°ÓxääÊÊUÊÊ>Ý\ÊÓn£°™x£°ÓxÓä www.tigertowerserivces.com Select 75 at www.HydrocarbonProcessing.com/RS

Safety Developments TABLE 2. Asset impact levels vs. asset value and production losses Level

Asset loss value

Production loss, bbl

S0

$10,000

0

S1

$100,000

1,000

S2

$1,000,000

5,000

S3

$10,000,000

15,000

S4

$100,000,000

50,000

S5

$1,000,000,000

100,000

Calculating the weighted KPI for a protection layer. The KPI for a LOP can be calculated as: Kj

e∑ i i KPI _ LOPj = K (w ) e∑ i i

(w ×KPI i )

where: KPI_LOP = Weighted average KPI of a layer of protection w = Weight of a KPI7 KPI = Key performance indicator related to plant, process, people (as applicable) K = Number of KPIs for an LOP I = Index for counting number of KPIs J = Index for counting number of LOPs. Calculating safety performance index for equipment.

Consider that a piece of equipment has a number of LOPs. From a safety perspective, the LOPs are of different importance and risk levels. From the LOPA, each layer has an associated risk-reduction factor. The weighted KPIs associated with the equipment can be aggregated and weighted, using the riskreduction factor associated with the LOP: SPI _ EQUIPj =

∑

Lj i

w _lopi × KPI _ LOPi

∑

L i

wi

=

∑

where: E = Number of pieces of equipment in a plant I = Index used to count the pieces of equipment in the plant EQ_RISK = Total mitigated risk for a piece of equipment 8 SPI_PLANT = SPI for the plant Estimated losses associated with LOE risk and plant. Based on the SPI, a safety performance state can be calculated. For example, the SPI can have ranges such as good (> 95%), warning (90% to 95%) and bad (< 90%). Associated with each LOE is an asset impact. For example, the asset impact may be defined as S0 to S5, as shown in TABLE 2. Incremental estimated asset value-at-risk is a safety performance adjusted metric (expected value) that can be calculated using the SPI, the safety performance state and the asset impact. For example, the incremental asset value-at-risk can be estimated as follows: 100% of the asset loss value-at-risk if the safety performance state is determined to be “bad”; 50% of the asset loss value-at-risk if the safety performance state is determined to be “warning”; 0% of the asset loss value-at-risk if the safety performance state is determined to be “good”: LOE: Estimated incremental asset value-at-risk: 0 if SPI > 95% = 0.5 defined asset impact if SPI 90% and 1 defined asset impact if SPI < 90%

95%

The plant-level incremental asset value-at-risk can be estimated by adding the estimated incremental asset values-at-risk for the LOEs with the facility. The plant-level incremental production value-at-risk can be estimated by adding the incremental production values-at-risk for the underlying lines of equipment: Plant: Estimated incremental asset value-at-risk: = ∑ LOE incremental asset value-at-risk Plant: Esimated incremental production capacity at risk:

Lj i

rrfi ×KPI _ LOPi

∑

L i

wi

⎪⎧⎪0 if Plant SPI > 95% ⎪⎪0.5×defined production capacity if Plant SPI ≥90% = ⎪⎨ ⎪⎪ and ≤95% ⎪⎪ ⎪⎩1×defined production capacity if Plant SPI < 90%

where: L = Number of layers of protection w_lop = Weight of a layer of protection (= RRF for the layer of protection) I = Index for counting LOP J = Index for counting number of pieces of equipment.

For a corporation with many plants, the incremental asset values-at-risk and the product values-at-risk can be aggregated as:

Calculating safety performance index for a facility.

Corporation: Estimated incremental production capacity at risk:

Consider that a facility has a number of LOEs. From a safety perspective, LOEs are of different importance/risk levels. From the LOPA, each LOE has associated with it a total equipment risk. The SPIs for the LOEs can be aggregated using the total risk factor calculated from the LOPA:

SPI _ PLANT =

∑

E i

1 ×SPI _ EQUIPi EQ _ RISK i E 1 ∑ i EQ _ RISK i

62MARCH 2013 | HydrocarbonProcessing.com

Corporation: Estimated incremental Asset value-at-risk: = =

Plant incremental asset value at risk Plant incremental asset value-at-risk

Dashboards. To display the SPI and related incremental asset value-at-risk and incremental production loss, dashboards can be used, as shown in FIGS. 8 and 9. The plant-level dashboard could display the plant safety-performance data and provide drill-down capability to the underlying KPIs for analysis of the underlying causes of identified risks. Once identified, corrective action plans can be defined and implemented in a timely manner to avoid costly catastrophic safety events.

Partner with the Best

Safety Developments Best practices and lessons learned. As proven with the name of the American Fuel and Petrochemical Manufacturers’ (AFPM’s) safety conference, i.e., the National Occupational and Process Safety Conference, the refining and petrochemical industries are clearly focused on PSM as a key component of their operational strategies. To support these operational strategies, there are nine steps or best practices to use when implementing and maintaining an effective process safetyperformance management system: Step 1. Establish the organizational arrangements/relationships needed to implement indicators. Step 2. Decide on the scope of the indicators. Step 3. Identify the risk-control systems and decide on the outcomes. Step 4. Identify critical elements of each risk-control system. Step 5. Establish the data collection and reporting system. Step 6. Review (benchmark against the IE PSM Framework or equivalent). Step 7. Deploy the KPI model and SPI calculations. Step 8. Educate management on the importance of PSM. Step 9. Establish management roles and actions for review of KPIs, SPIs, estimated asset value-at-risk and estimated production value-at-risk. LITERATURE CITED “A Canadian Perspective of the History of Process Safety Management Legislation,” 8th International Symposium: Programmable Electronic System in Safety-Related Applications, Sept. 2–3, 2008, Cologne, Germany. 2 Center for Chemical Process Safety website: http://www.aiche.org/CCPS/ Students/GetSmart.aspx. 3 Center for Chemical Process Safety website: http://www.aiche.org/CCPS/ Students/GetSmart.aspx. 4 H. J. Toups, LSU Department of Chemical Engineering, with significant material from SACHE 2003 Workshop. 5 Managing the Risks of Organizational Accidents, Ashgate Publishing Co., 1997. 6 Energy Institute, London, 1st Ed., December 2010. 7 A weight of 0 signifies that a KPI is not used. 8 This is equal to the sum of all the mitigated risks for an item of equipment. 1

MARTIN A. TURK, PhD is the director of Global Industry Solutions for the HPI for Invensys Operations Management at Houston, Texas. For most of his 40+ years of experience, Dr. Turk has been involved in engineering, consulting, sales and marketing activities related to process automation. These activities include process simulation, advanced control and information/automation system strategic planning. Dr. Turk is responsible for definition of industry-specific solutions for downstream petroleum refining and petrochemicals, participation in industry conferences and working with Invensys clients worldwide to identify and quantify automation opportunities in their manufacturing facilities that will provide them with significant returns on investments. He received his BS degree in chemical engineering from the University of Dayton and his PhD in chemical engineering from the University of Notre Dame. Also, he has published technical papers and made presentations at domestic and international seminars on a variety of subjects related to advanced automation solutions for the process industries.

With over 50 independent subsidiaries and more than 220 engineering and sales offices spread across the world, SAMSON ensures the safety and environmental compatibility of your plants on any continent. To offer the full range of high-quality control equipment used in industrial processes, SAMSON has brought together highly specialized companies to form the SAMSON GROUP.

AJAY MISHRA is the R&D program manager at Invensys. He helps define the detailed features and technology roadmaps for the Triconex branded safety & critical control products. Mr. Mishra holds a BSEE degree from the College of Engineering, Pune, India and an MBA from the UCLA Anderson School of Management. He has over 20 years of experience in safety and critical control systems in process control SIS, and railways systems including product development, project engineering, project management and product management. Mr. Mishra is a TÜV certified Functional Safety Engineer for hardware/software design (IEC 61508) and Safety Instrumented Systems (IEC 61511). A01120EN

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SAMSON AG · MESS- UND REGELTECHNIK Weismüllerstraße 3 60314 Frankfurt am Main · Germany Phone: +49 69 4009-0 · Fax: +49 69 4009-1507 E-mail: [email protected] · www.samson.de SAMSON GROUP · www.samsongroup.net Hydrocarbon Processing | MARCH 2013

63

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Bonus Report

Safety Developments R. LIMAYE, Praxair, Houston, Texas

Operator response to alarms is an important protection layer Operator response to alarms is important layers of protection (LOP). When implemented with good design, engineering and maintenance practices, an alarm can help reduce the safety integrity level (SIL) of a safety instrumented function (SIF). The ISA84 standard covers the functional safety life cycle requirements of safety instrumented systems (SISs) and has been recognized by the US Occupational Safety and Hazard Administration (OSHA) as a generally accepted good engineering practice that can be used to comply with the process safety management (PSM) standard 29CFR1910.119 for SIS. In the initial phase of the safety life cycle, hazard and risk assessments identify the risks associated with the process. The risk is reduced to a tolerable level as defined by the corporate standard by implementing safety functions that act as LOP. Layers of protection analysis (LOPA) is the most commonly used technique in the process industry, especially for selecting the amount of risk reduction required to be provided by each protection layer. If the risk is not reduced to a tolerable level using non-instrumented safety functions such as pressure safety valves (PSVs), mechanical stops and process design, then additional risk reduction is necessary by means of a SIF implemented in a SIS. Often, the performance requirement of a SIF can be reduced by one order of magnitude by implementing an operator response to an alarm as a protection layer. Reducing SIL requirements (e.g., from SIL3 to SIL2) offers huge capital and maintenance cost savings. The higher the SIL, the higher the cost of implementation and maintenance. The ISA18.2 standard describes the alarm management life cycle. However, the standard does not give enough guidance on the alarms that have been credited in LOPA for risk reduction. The standard talks about the alarm classification and defines “highly managed alarms” as a special class that has requirements similar to the alarm for which credit has been taken in LOPA for risk reduction. The rules for taking credit for an operator response to an alarm are defined in the ISA84 standard. Teams performing alarm engineering may not have familiarity with the requirements of ISA84; therefore, those requirements are outlined here. If alarm management and functional safety teams come up with separate lists and databases, it becomes difficult for operations to maintain and keep the lists up-to-date. The best practices for maintaining a common list as a master alarm database will be discussed here. The ISA 18.2 alarm management standard provides a life cycle approach to manage alarms starting from alarm philosophy and rationalization to operations and maintenance.

It states the mandatory and nonmandatory requirements for alarm engineering in different stages of the life cycle. The standard refers to ISA84 for the requirements of alarms related to process safety. In many organizations, the teams who perform the alarm management life cycle steps are different from those who work on SISs and may not be familiar to the requirements of ISA84, which imposes specific requirements on the alarms related to process safety. The LOP model identifies the process alarm as one of the protection layers in reducing the demand rate on SISs. Although casual verbal or written communication refers to alarms as a protection layer, the correct term is “operator response to alarm as protection layer.” LOPA. To illustrate the concept, consider a process example

consisting of a high-pressure (HP) knockout drum, as shown in FIG. 1. The product gas at high pressure and temperature exiting the reformer is cooled in a process gas cooler, which condenses the water in the gas. The condensate is removed in the knockout drum V-100. The liquid level in the knockout drum is controlled by a level control loop, LC-100. The normal operating range of 35%–50% has been established to maintain a liquid level blanket in the knockout drum. If the level goes high, the Product gas

Knockout drum V-100

HP product gas + condensate

Normal operating range

LAL 100 50

Low alarm

35 30

Low level trip

10

SP LT PV LC 100A 100

LT 100B

LALL 100

LT 100C

SIF 1

OP

V-101 Regulating valve LV-100

On/off valve XV-100 LP storage tank

FIG. 1. An example of the knockout drum process. Hydrocarbon Processing | MARCH 2013

65

Safety Developments separation will not occur, as the liquid will be carried over into the product stream. If the level drops too low, there is a risk of the HP gas entering the low-pressure (LP) system downstream of the control valve LV-100. Maintaining the liquid level blanket to avoid this hazardous event is very important. The low alarm limit for the pre-trip safety-related alarm is set to 30%. A SIF is implemented to prevent the HP gas from entering the LP system. If the level falls below the trip limit of 10%, then the on-off valve XV-100 closes. During hazard and risk assessments, all possible process risk scenarios are identified and documented. Hazard and operability (HAZOP) studies are one of the most commonly used techniques of performing process hazard assessments. The scenarios are risk ranked based on the risk matrix developed by the organization. The risk matrix typically defines different levels depending on the severity of consequence and likelihood of the hazardous event. The scenarios with risks higher than a predefined threshold level are considered for LOPA to ensure adequate LOP exist. For the process scenario described previously, an extract from the LOPA is presented in TABLE 1. The safety functions are assigned to LOP during LOPA. Two independent protection layers are identified in the LOPA example above. The first protection layer is the operator response to an alarm with a risk reduction factor (RRF) of 10. The RRF is also represented in terms of probability of failure on demand (PFDavg ). In this instance, PFDavg = 1/RRF = 1 X 10-1. As per ISA84, the maximum risk reduction that can be assigned to operator response to an alarm implemented in a basic process control system (BPCS) is 10. The second protection layer is a SIF with an RRF greater than 100. A SIF is a combination of sensor, logic solver and final control element. The logic solver is often referred to as the SIS. TABLE 1. An example of LOPA Hazard description Loss of primary containment of the product gas upon failure of level control. Loss of liquid blanket in KO drum V-100 and HP flammable gas entering the LP system through KO drum bottom with a possibility of explosion and fatality. Description

Probability (frequency per year)

Tolerable risk (defined by organization) [ 1 in 100,000]

1 X 10–5

Likelihood of initiating event (control loop failure) [1 in 10]

1 X 10–1

Probability of ignition

1

Likelihood of operator present near the vessel

1 X 10–1

Frequency of unmitigated consequence

1 X 10–2

The logic of a SIF is executed in the SIS. The amount of risk reduction provided by the SIF determines the SIL of the SIF. As shown in TABLE 2, there are four SILs defined in the ISA84 standard that offer different bands of risk reduction. In the LOPA example, the SIF must offer an RRF of greater than 100 to reduce the overall PFDavg to less than 1x10–5 and thus the SIF has a SIL requirement of SIL2. Without the alarm, the SIF must have an RRF greater than 1,000 (or PFDavg < 1 X 10–3), making it a SIL3 SIF. A SIL3 SIF requires considerably higher capital cost to implement. For example, it may require two ON/OFF valves in a double block and bleed arrangement, an additional pressure transmitter for pressure between the two block valves and a logic solver with much more stringent requirements of failures rates. Safety related alarm. An alarm is called a safety related alarm (SRA) when operator response to the alarm is used as a protection layer to reduce the overall risk of hazardous process events. This alarm is identified in LOPA and a risk reduction credit is taken for the operator response to the alarm. Other terms used to describe this alarm type are: critical alarm, safety-critical alarm and process safety alarm. The ISA18.2 standard talks about an alarm class called highly managed alarms (HMAs). The ISA84 life cycle requirements of SRA are similar to those of highly managed alarms, although many people do not like to use the term “highly managed alarms” to classify safety-related alarms. Considerations for using operator response to an alarm as a protection layer are summarized as: • The sensor used for the alarm system is not used for control purposes where loss of control would lead to a demand on the SIF • The sensor used for the alarm system is not used as part of the SIS • Limitations have been taken into account with respect to risk reduction that can be claimed for the BPCS and common cause issues • Risk reduction claimed is not more than a factor of 10 • There is sufficient time for the operator to take corrective action • Documented description of the response to the alarm (corrective action) is available, and rationalization has been performed • The operator has been trained to take preventive actions • Performance shaping factors have been considered • Human ergonomic factors have been considered • The test and maintenance requirements are the same as any other independent protection layer TABLE 2. SIL determination Demand mode

Independent protection layers IPL description IPL-1 IPL-2

Probability of failure –1

Operator response to low alarm in KO drum

1 X 10

SIF to close the valve XV-100 upon low-low level alarm

1 X 10–2

Frequency of mitigated consequence

66MARCH 2013 | HydrocarbonProcessing.com

SIL 1 2 3

–5

1 X 10

4

Target RRF

Target PFDavg –2

–1

> 10 to < = 100

–3

–2

> 100 to < = 1,000

–4

–3

> 1,000 to < = 10,000

–5

–4

> 10,000 to < = 100,000

> = 10 to < 10 > = 10 to < 10

> = 10 to < 10 > = 10 to < 10

Safety Developments • The sensor used for generating alarms should be tested at a proof test interval established in the safety requirements specifications • The person who performed the tests and any maintenance should be documented and archived for the records • Access control: Access to make changes to alarm parameters such as alarm setpoint, priority and filter constant are restricted, and the proper management of change (MOC) process is followed to make any changes once the system is put in service. Independency requirements. ISA84 clause 11.2.10 states that one should not share a device with a SIF and control function where failure of the device will cause the BPCS loop to place a demand on the SIS and simultaneously cause the SIF to fail in a dangerous state. Therefore, the same sensor used for generating an alarm cannot be shared with the control function in a BPCS and cannot be shared with the SIF implemented in the SIS. When independent sensors are available for each function, they can be configured in a fault tolerant mode to achieve higher reliability and to increase the diagnostic coverage. Each owner operator company is responsible for doing analysis to ensure their configurations are valid and can satisfy the independency criteria of ISA84. Standalone SRA. When the operator response to an alarm is identified in LOPA as a protection layer, and there are no other safety instrumented functions or control functions associated with the measurement, then the transmitter is wired to a BPCS, as shown in FIG. 2. SRA and control function. SRAs and control functions are both implemented in BPCSs. When the same process measurement is required for both the functions, separate transmitters should be used for alarm and control. Each transmitter should be wired to separate cards in the controller or preferably separate controllers of a BPCS. The two transmitters could be configured in a BPCS, as shown on the right of FIG. 3, to improve the availability, facilitate the maintenance and to improve the diagnostic coverage. The right side of FIG. 3 also illustrates that a software switch (HS-100x) is provided for the operator to manually change the source of input for the alarm, as well as for the control function. This allows taking one of the transmitters out of service for maintenance or proof testing. Depending on the functionality in the BPCS, a deviation alarm should be configured for the maintenance technician. If the difference between the readings of two transmitters is more than a pre-set threshold, a low priority deviation alarm is generated. The operator action for this alarm is typically to generate the maintenance work order for the instrument technician to check the transmitters and correct the situation. When the hand switch is used to switch the input to another source, a timer KS-100x should be started with alarm KAH100x. If the time in the switched input mode exceeds the preconfigured limit, an alarm should be generated. The preconfigured timeout limit to generate a warning alarm should be less than the minimum time to repair (MTTR) of the transmitter. If the time in the switched state exceeds the MTTR, then the SIF should initiate the action to put the process in a safe state.

SRAs, control functions and SIFs all need the same process measurement. When the available instrumentation is adequate to meet the independency criteria of each function, it is beneficial to wire it in a fault tolerant configuration to improve reliability. A typical implementation is shown in FIG. 4. In this example, the three differential pressure transmitters using independent taps for process connection are wired to a SIS through the safety certified current-loop isolator and repeater. The HART transmitters are powered by a SIS and the isolators have the capability to pass through the HART signals on each channel. SIL1 or SIL2 SIF is implemented in a SIS with a 2oo3 voting logic for the level input signals. Depending on the BPCS, the actual implementation may differ. FIG. 4 shows a generic representation where a middle of 3 selector block is used in a BPCS that has an SRA configured. Some BPCS have a standard 2oo3 block that can be used as well. When sharing the transmitters between a BPCS and SIS, several considerations should be taken into account: • Failure of any hardware or software outside the SIS should not prevent any SIF from operating correctly

LT 100B

LI 100B

SRA

LAL 100B

FIG. 2. P&ID representation of a stand alone SRA. SRA

LAL SRA 100B LI 100B

LC 100

LI 100B

LT 100B

LT 100A

LT 100B

LV-100

KAL 100B KS 100B HS DEV 100B

LAL 100B

KAL 100A KS 100A HS 100A

LC 100 LT 100A

LV-100

FIG. 3. SRA and control functions are both implemented in BPCSs.

LAL 100

SRA

LC 100

DEV

LALL 100 m of 3

m of 3 LI 100A

I

I

LY-100A LT 100A

LI 100C

LI 100B

I

I

LY-100B

SIF 1

I

I

LY-100C

LT 100B

LT 100C

XV-100

LC-100

FIG. 4. SRAs, SIFs and control functions are implemented to achieve high reliability. Hydrocarbon Processing | MARCH 2013

67

Safety Developments cess engineers. Such calculations become easy when the process model is available. Alarm response time is the difference between the time at which the alarm condition occurs and the time when the process starts responding in the direction to correct the alarm condition. It includes the sensor lag, BPCS lag, operator response time and any process lag. Process dead time is the amount of time it takes for the process to begin reacting after corrective action. Alarm response time = Sensor delay + BPCS delay + Operator response time + Process dead time The process safety time for alarm has to be greater than the alarm response time. These different time elements are shown in FIG. 5. Operator response time is impacted by things like human factors and ergonomics, which are collectively called performance shaping factors. As per the ISA18.2 feedback model of operator process interaction, the operator response time constitutes the following human interactions: Detect: The operator becomes aware of the deviation from the desired condition. The design of the alarm system and the Time considerations. Time available for the operator to take operator interface impact detection of deviation. corrective action is an important factor in alarm engineering. Diagnose: The operator uses knowledge and skills to interIn the alarm rationalization process, the alarm priority is impret the information and diagnose the situation when determinpacted by the severity of the consequence of inaction to alarms ing the corrective action to take in response. and the time available for the operator. Alarm philosophy Respond: The operator takes corrective action in response documents typically recommend the use of a rationalization to the deviation. matrix of consequence severity and maximum time to respond. Minimum time to respond is the quickest possible time Process safety time (PST) is the difference between the to allow an operator to go through the detect, diagnose and retime at which the unacceptable condition occurs (TCONDITION ) spond steps. It should be defined in the alarm philosophy docuand the time where the unwanted event occurs (TEVENT ). ment. It is not physically practical to take necessary corrective Process safety time = TEVENT – TCONDITION actions in less than this time. Three to 10 minutes is the most In the previous example, there are two protection layers. commonly used value as a minimum time to respond. The process safety time for an alarm is the time when the levIf the required operator response is faster than the minimum el reaches 30% until it goes to the trip setpoint of 10%. The time to respond, then no credit can be taken for the operator time can be calculated by dividing the volume of the knockresponse to alarm as a protection layer. This requirement is apout drum for the 20% of instrument range (the difference beplicable to not just the SRA, but also to any alarm configured tween the alarm setpoint and the trip setpoint) by the worst in the system. In such situations, various options should be case flowrate of condensate when the level control valve LVreviewed to allow sufficient time for operator response. In the 100 stays wide open. The process safety time for the SIF can above process example, the simplest option is to check if a low be calculated in a similar manner. It will be the time when the alarm setpoint (LAL-100) can be increased to get more process level reaches a trip point of 10% until it goes to 0%. safety time. Using a restriction orifice to limit the maximum Typically, the process safety time calculation is done by proflow could be another option. Maximum time to respond is just Alarm Note: Not to scale what it says, the maximum time available condition 35-50% Normal operating range Alarm occurs for an operator to respond. This does not 30% Low alarm setpoint 30 Trip annunciated 10% Trip setpoint condition mean the operator will take this much occurs time to respond. Incident Sensor delay 10 Maximum time to respond = Process Detect DCS delay Diagnose Respond safety time for alarm – (Process dead time Safe state + Sensor and BPCS delays) 0 RL RL In most cases, sensor delay and BPCS Time Operator response Dead time Dead time delay are very small and could be insignifSensor Process Process delay Read Logic Write icant. Depending on the application, proresponse response cess dead time could also be insignificant. Final control element action Therefore: Alarm response time SIF response time Maximum time to respond ≈ Process Process safety time—alarm Process safety time—SIF safety time for alarm. Mean time to respond is the actual FIG. 5. Various time elements in relation to process safety time. operator response time, somewhere beLevel %

• Failure of a BPCS component does not result in the initiating cause for the process hazard and the failure (or defeat/ bypass) of the SIF that protects against the specific scenario under evaluation • The probability of common mode, common cause or dependent failures has been adequately evaluated and determined to be sufficiently low; it is often recommended to use diverse measurement technology to reduce the common cause failure problems (a combination of differential pressure and guided wave radar transmitters is an example of using diverse technologies to measure the same process value) • The shared components are managed according to ISA84, including proof testing, access security and management of change • The sensor is powered by the SIS • The signal is transmitted to the BPCS by an optical isolator or other means to ensure that no failure of the BPCS affects the functionality of the SIS.

68MARCH 2013 | HydrocarbonProcessing.com

Safety Developments tween the minimum and maximum time to respond. In FIG. 5, the operator response time shown represents the mean time to respond. The detect, diagnose and respond actions shown in the graph represent the average time for each action. Alarm rationalization is the process to review potential alarms, using the principles of the alarm philosophy, to select alarms for design, and to document the rationale for each alarm. To determine the alarm priority, the rationalization process typically uses a matrix of maximum time to respond and the severity of the consequence of not attending to the alarm. SIF response time includes the sensor delay, analog input card scan time, logic execution, writing output to the final control element and the time it takes for the valve to close. Typically, valve closing time is most significant in these time elements. SIF response time should be less than half of the process safety time for the SIF. Managing databases. Usually, the functions of alarm management, process engineering and SIS engineering are performed by different individuals or teams in an organization, as each area requires different competencies and experience. It is important to have good coordination between these teams. The purpose of alarm documentation and rationalization is to publish a master alarm database, which needs to be maintained by operations to sustain the benefits of alarm engineering efforts done upfront. Process engineers typically maintain a database of normal operating limits, safe operating limits, design limits, engineering unit ranges and the maximum allowable working pressure/ temperature (MAWP, MAWT) for each process measurement. This database is expected to be maintained throughout the life of the plant. The PSM group and SIS engineers identify the safety-related alarms during LOPA. These alarms are listed as a protection layer with a risk reduction credit of up to 10 (PFDavg >= 10–1). Usually, someone ends up generating a list of SRAs that get passed on to the operations with an expectation of maintaining it throughout the life cycle. Managing and maintaining all these databases can become an overwhelming task for maintenance and operations organizations. If proper procedures are not in place, such lists and databases become outdated and the plant may be out of compliance. Coordination of these different groups to generate one single relational database goes a long way in maintaining the integrity of the information. The solution could be using a simple homegrown relational database or one of the commercial software products available in the market. In either case, establishing the work process that coordinates the effort of different teams is a key element. Alarm classification. Alarm classes are used to set common characteristics and requirements for managing alarms. It is a good practice to classify all alarms identified in LOPA as safety-related alarms. It is also recommended to identify these alarms on a P&ID with SRA written next to the alarm, as shown in FIGS. 2–4. In addition to alarms identified in LOPA, the diagnostic alarms on subsystems involved in SIFs should also be classified as SRAs. When a dangerous fault is detected in the subsystem (as a result of diagnostic or proof testing) which can tolerate a single hardware fault, the process can continue the safe operation

while the faulty part is being repaired. If the repair of the faulty part is not completed within the mean time to restoration assumed in the calculation of the probability of random hardware failure, then a specified action should take place to achieve or maintain a safe state. When the action involves the operator notifying maintenance to repair a faulty system in response to a diagnostic alarm, this diagnostic alarm may be a part of the BPCS but should be subject to appropriate proof testing and MOC along with the rest of the SIS. An alarm can be assigned to multiple classes. The objective is to create or view different types of lists, using the filters on alarm classes. The diagnostic alarms could be assigned a class maintenance related alarm (MRA) in addition to a SRA to indicate it is not a process alarm. ISA18.2 describes highly managed alarms (HMAs) as the classes of alarms that require more administration and documentation than others. The alarms identified in LOPA fall in this class. Depending on the alarm philosophy, the alarms identified in LOPA could be assigned to both SRA and HMA alarm classes. Maintenance and documentation. Sensors used for generating safety-related alarms should be treated as part of the SIS with respect to the proof test procedures and documentation required to be maintained as evidence of satisfactory completion of those procedures. The records that certify proof tests and inspections were completed require the following as a minimum of information: • Description of the tests and inspections performed • Dates of the tests and inspections • Name of the person(s) who performed the tests and inspections • Unique identifier of the system tested (for example, tag name, equipment number) • Results of the tests and inspection (for example, “as-found” and “as-left” conditions). Final thoughts. When operator response to alarm is used as

one of the protection layers to reduce the process risk below a tolerable level, it is important to pay attention to the requirements of the ISA84 standard. The sensor used for generating an alarm should not be shared with other functions, such as process control and safety functions. When different functions need the same process measurement and if multiple sensors are available to satisfy the independency criteria of each function, then the sensors should be wired in a fault tolerant configuration to improve the reliability of all the functions. Another important consideration is to ensure that the operator response time for the alarm is less than the process safety time. Coordination of the different teams in an organization to use and maintain a single relational database for alarm rationalization, documentation, safety-related alarms and normal operating limits goes a long way in ensuring the data integrity for the life of the plant. RAJEEV LIMAYE is director of control systems and instrumentation at Praxair’s global hydrogen business unit in Houston, Texas. He has worked in the process automation industry for over 25 years and holds a PE license in the state of Texas. Mr. Limaye is a certified functional safety expert and a member of the ISA84 standards committee. Hydrocarbon Processing | MARCH 2013

69

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Bonus Report

Safety Developments V. RAMNATH, Aker Solutions, Pune, India

Key aspects of design and operational safety in offsite storage terminals The Buncefield incident in the UK and the Jaipur fire in India stand out as major offsite storage terminal disasters in the European and Asian regions. In the Americas, a similar explosion rocked the Gulf terminal in San Juan, Puerto Rico, destroying 22 tanks and forcing the evacuation of hundreds of nearby residents. Storage terminal accidents continue to impose considerable costs in terms of human lives and health, property damage and public welfare. Multiple segments of the public—particularly citizens, citizens’ groups and the media—are likely to become more interested in chemical safety and chemical release risk reduction once they become aware of the potentially large consequences if proper proceedures are not followed. On December 11, 2005, at the Hertfordshire oil storage terminal in Buncefield, UK, an unconfined vapor cloud explosion (UVCE) occurred that eventually overwhelmed 20 large storage tanks (FIG. 1). Fortunately, there were no fatalities; however, 60 people were injured and there was large-scale damage and disruption to local businesses and residents.

Process safety improvement. These incidents highlight the

need for an exhaustive design and safety review of practices while designing and operating offsite storage terminals to minimize the risk of fatalities and property damage. Process safety and risk reduction can be effectively achieved by applying multiple layers of protection (FIG. 3), which are known in the industry as “onion layers.” Each protection layer, when called for, can function independently to prevent or miti-

Recent events. A major accident involving a gasoline vapor

cloud explosion followed by an ignition occurred at Indian Oil’s petroleum, oil and lubricants (POL) terminal on October 29, 2009, in Jaipur, India (FIG. 2). Buildings in the immediate neighborhood were heavily damaged. The effect of this accident was felt up to 2 km from the site.

FIG. 2. Fire at the Indian Oil storage terminal in Jaipur, India.

Emergency response Embankment

Mitigation layers

Relief devices SIS Alarms, operators BPCS

Prevention layers

Process

FIG. 1. Damaged tanks at the Buncefield disaster.

FIG. 3. Protection layers. Hydrocarbon Processing | MARCH 2013

71

Safety Developments • Conduct a critical quantitative risk assessment (QRA) of the facility, which should stress worst case scenarios • Consider fuel storage terminals as hazardous locations • The control room should be located remotely from potential leak sources as is practical; the QRA will determine the blast pressure structural design requirements for control rooms Process safety and risk reduction can be • Water tanks and the water pump house should effectively achieved by employing multiple be located a safe distance from potential leak sources and the tankage area layers of protection, which are known in the • Locate buildings and structures in the upwind industry as “onion layers.” Each protection direction • Avoid congestion in the storage terminal area, layer can function independently to prevent with the location of individual facilities determined or mitigate the undesirable event. by the QRA • All buildings not related to terminal operation should be located outside the plant area, including the canteen or any other area where sparks and open flames may exist prevention layers is to reduce the frequency of the undesired • When a tank storage terminal includes pipeline operations event; whereas, mitigation layers take action in the event of in the same location, the control rooms for both the tank termifull or partial failure of the prevention layers. nal and pipeline division should be located in the same operaIn the context of storage terminals, several key aspects can tional building be implemented that play a role in prevention or mitigation of • Locate the emergency exit gate away from the main gate uncontrolled releases. and in a location that is always available for use • Consider radar-level gauges as minimum instrumentation Prevention layers. For terminal design, engineers should: • Leak-detection devices must be installed on the bottom • Ensure that the receiving tank has ultimate control of of all tanks tank filling, allowing it to safely stop or divert a transfer with• Install cathodic protection to prevent deterioration of the out depending on the actions of a remote third party tanks through corrosion • The first body valve on the tank should be a fail and fire • To prevent an overflow, tanks should have headspace marsafe remote operated shutoff valve (ROSOV) on the tank nozgins that enable the filling line to be closed off in time. zle inside the dike, operable remotely from outside the dike as For process operations, engineers should: well as from the control room • Develop procedures for periodic testing of overfill pro• All other operational valves must be outside the dike area tection • The piping layout inside the tank dike area should ensure • Ensure effective communication within and between opeasy accessibility for any operation erations, maintenance and contractors • Structural design should be followed to ensure installa• Standard operating procedures (SOPs) should be pretion of sleeves to eliminate emissions from slotted guide poles pared that not only give what the procedures are, but also why on floating roof tanks. they are needed. • Use double seal systems for floating roof tanks where ap• Regular inspections of pipelines, including thickness surpropriate, based on the nature of the material being stored, the veys and pipeline support system analysis should be carried tank size, throughputs, location considerations and meteorology out and recorded • Carry out a hazardous operations (HAZOP) study of the • Closed-circuit television systems should be installed, facility during engineering and at decided frequencies during covering tank farm areas and other critical areas operation. • Near-miss reporting systems should be in place • Collate incident data on potential failures and operationProtection layers al failures and share the information on risks with the industry. • Develop incident investigation mechanisms for failures and safety malfunctions. Prevention Mitigation For basic control systems, engineers should: • The control system needs to have inventory checks to ensure that the receiving tank has adequate empty volume to Process F&G system, Emergency SIS receive material operation relief system response • Setpoints of high-level trips and alarms requiring operator action should allow sufficient time for action to be taken to Alarms Basic Bunds/dikes Terminal deal with the developing situation. control design system For alarms, engineers should: • High-level and high-high-level alarms should occur from FIG. 4. Classifying the layers of protection. independent sensors gate the undesirable event. These protection layers have the common goal to control and/or mitigate risk from the facility. Broadly, these layers, as shown in FIG. 4, can be classified into prevention and mitigation layers. The objective of the

72MARCH 2013 | HydrocarbonProcessing.com

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• High-level alarms should trigger facility operators to initiate corrective actions, resulting in a shutdown of the pumping system and isolation of the tank. For safety instrumented systems (SIS), engineers should: • The overall system for tank filling control should be of high integrity, with sufficient independence to ensure timely and safe shutdowns to prevent tank overflow • When changes or modifications to an SIS are planned, the changes should be subject to a management-of-change process identifying and addressing any potential safety implications from the modification • A functional safety assessment should be performed on each system, typically at the design stage before the system is commissioned. Mitigation layers. For fire and gas detection and relief systems, engineers should: • Review the classification of places where explosive atmospheres may occur and then evaluate the siting and/or protection of response facilities, such as fire fighting pumps and emergency switch devices • Apply measures to detect hazardous conditions arising from loss of primary containment • Install gas detectors in bunds • The thermal safety valve should be provided at the operating manifold (outside of the dike) • Hydrocarbon detectors need to be installed near all potential leak sources • Install a rim-seal fire detection and protection system in all floating roof tanks of the terminal

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• Each facility needs to maintain adequate firefighting equipment at each pump station and breakout tank area. The equipment should be plainly marked as firefighting equipment and located so that it is easily accessible during a fire. For embankment, engineers should: • Ensure secondary containment, like dikes and bunds, are constructed from concrete and are leak-proof; the lining requirements must prevent soil or ground contamination • Install systems that prevent or limit the passage of stored materials via drains to the normal effluent processing route. For emergency responses, engineers should: • Locate push buttons on motorized valves outside the dike, where it’s easily accessible by the operator • Terminal emergency switch devices should be located in the control room as well as in other strategic locations • Emergency procedures should be written and made available to all facility personnel, outlining the actions to be taken during a major incident • Mock drills should include the full shutdown system activation. • A system should exist to warn neighboring industries of impending danger and arrangements should be made to enable their assistance, if necessary. VINOD RAMNATH works for Aker Solutions’ process department in Pune, India. He has experience in the detailed engineering of process plants, offsite design, quantitative risk assessment, safety integrity studies and process plant commissioning/debottlenecking. Mr. Ramnath started his career with Reliance Industries. Hydrocarbon Processing | MARCH 2013

73

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Ziegler-Natta catalysts are primarily used in polypropylene (PP) production. These catalysts are very sensitive to various poisons and their activity varies according to the nature and level of catalyst poisons. Among others, typical propylene polymerization catalyst poisons include acetonitrile, arsine, carbon dioxide, carbon monoxide (CO), carbonyl sulfide, cyclopentadiene, ethylene oxide, oxygen, palladium, phosphine, moisture, methanol and propylene oxide. Catalyst poisons usually are present as impurities within feedstreams like propylene, ethylene and hydrogen. Each of the poisons has a varying degree of influence on the catalyst activity. Though their general behavior is known, it is always difficult to quantify the losses due to individual poisons within commercial-size plants. Among the difficulties, accurate measurement of trace impurities in ppb levels remains the biggest challenge. Experience shows that there are great difficulties in offline sampling as well as online sampling. Changes in process conditions during polymerization is also another factor. It is not always possible to normalize catalyst productivities against changing process parameters. CO is one of the strongest poisons for Ziegler-Natta catalysts. If not treated properly at the source unit, CO is present with the propylene feed as a contaminant and it can reduce catalyst activity drastically (FIG. 1 and FIG. 2). As per the published literature, a concentration of about 6 ppb (wt%) CO reduces catalyst activity by approximately 5%. Few PP process technologies specify feed specification as low as 20 ppb (wt%) in propylene feed to achieve the guaran-

teed figures of catalyst activity. If the feed-treating unit is not designed carefully to knock down the CO levels to the safe limit, the facility’s bottom line will shrink due to excessive catalyst usage and cost. Commercial examination. This study from a commercially

operating unit shows the importance of using a proper guard against possible feed contaminants acting as catalyst poisons. The unit has two different primary sources of propylene supply. Source A has CO treatment beds (typically, a sulfur removal bed followed by a CO removal bed) while Source B doesn’t have any guard beds to protect against CO (FIG. 3). Normally both sources feed propylene to a PP unit. In this case, one of the sources is under a shutdown and an alternative Source C is used to meet the propylene demand. Water Oxygen Methyl acetylene Methanol H 2S COS CO CO2 C4 dienes

Water

Arsine

Oxygen

Propadiene

Methyl acetylene

Acetylene

Relative poison strength

0

Methanol

20

40

60

80

FIG. 2. Relative poison strength.

H2S COS

Propylene Source A

CO CO2

Sulfur removal

CO removal

Dryer

C4 dienes Arsine Propadiene

Propylene Source B

ppm for 20% catalyst activity reduction

Acetylene 0

50

100

150

FIG. 1. Catalyst activity reduction due to different poisons.

To PP plant

Dryer

200

FIG. 3. Propylene source configuration. HYDROCARBON PROCESSING | CATALYST 2013

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Catalyst activity, tons PP/kg catalyst

CATALYST

28 26 24 22 20 18 16 14 12 10

Source B+C Source A shutdown

Source A+B

1

residence time and at the same H2 concentration (same polymer grade) in a gasphase reactor. FIG. 4 reveals that catalyst activity is approximately 25% lower when untreated propylene feed Source A is online. This translates into an additional cost of approximately $5 million/yr due to excessive catalyst consumption. FIG. 5 illustrates the effect that varying 25 26 27 28 29 30 31 CO levels in propylene feed Source A have on catalyst activity for a period of nine days on a real-time basis. The figure shows that the CO level is increasing while catalyst activity is dropping. A loss of approximately 45% catalyst activity is seen when CO levels go up from approximately 90 ppb to approximately 250 ppb. FIG. 6 portrays a relationship in a commercial reactor between levels of CO in a propylene feed vs. catalyst activity when other process parameters are kept constant. A simple extrapolation yields a further activity rise by 10% to 13% if there is no CO contamination in the propylene feed. Source A+B+C

Source A startup

2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 Date of month

FIG. 4. Effect of untreated propylene source on catalyst activity.

Solution. Since the source of the CO contamination was known

FIG. 5. Effect of CO in propylene on catalyst activity. FIG. 4 shows the effect on catalyst activity when Source A (without CO guard beds) goes offline and then comes online again. The daily average catalyst activity is plotted for an entire month when the polymerization unit is running at the same

and the average CO level was confirmed by the online CO analyzer, it was easier to calculate lost revenue. Various options were evaluated against the cost and feasibility of carrying out the change. In FIG. 7, it was proposed to partially fill the existing dryer on Source B with CO-removal catalysts. This proposal would save on potential capital and construction costs associated with installing a new bed, associated piping and required plot space. All operational aspects and regeneration were carefully evaluated and documented accordingly. The end result was a new mixed bed with a molecular sieve and CO-removal catalysts. After commissioning, an immediate increase was observed in catalyst activity.

CATALYST NEWS NEW STEP-OUT CATALYST FOR ULSD AND HYDROCRACKER PRETREAT

Haldor Topsøe has announced an improved catalyst preparation technology. In hydrocracking pretreatment and highpressure ULSD services, the catalyst exhibits 40% higher activity than its predecessor. The higher activity can be used to: • Achieve longer cycles at the same feed rate • Process tougher feeds • Increase conversion to achieve lower product sulfur • Increase throughput. The technology offers an improved production technique for hydroprocessing catalysts. It combines previous technology with a catalyst preparation step developed by Topsøe’s research department. The combined effect from merging the two technologies has led to a metal slab structure that is characterized by an optimal interaction between active metal structures and the catalyst carrier. As researchers from Topsøe demonstrated in the 1980s, the activity of Type II sites is very strongly influenced by this interaction. The technology affects C–78

CATALYST 2013 | HydrocarbonProcessing.com

the way in which the metal slabs are bound to the carrier, resulting in the activity of both the direct sites and the hydrogenation sites being increased significantly. These new catalysts exhibit high stability with the added benefit of a step change in hydrodesulfurization and hydrodenitrogenation activity. NEW CATALYSTS CAN HELP AUTOMAKERS MEET CARB LEV III EMISSION REGULATIONS

The California Office of Administrative Law recently approved the more stringent LEV III automotive emissions regulations for new passenger vehicles proposed by the California Air Resources Board (CARB). BASF’s direct ozone reduction catalyst systems can help automakers meet these increasingly strict LEV III low-emission vehicle regulations. When applied to heat exchange surfaces, such as car radiators, the BASF system effectively turns them into “smog eaters.” The catalyst reduces ground-level ozone in the air that passes over coated surfaces by converting ozone molecules into oxygen molecules instantly upon contact.

CATALYST

50-200 wt ppb CO

Current

< 20 wt ppb CO

Proposed

Molecular sieve

CO removal catalyst Molecular sieve

FIG. 7. Present and proposed scheme for using existing bed with

two types of material. 0.14

25 20

FIG. 6. Relationship between CO and catalyst activity in a

commercial reactor.

As shown in FIG. 8, CO at the outlet of the bed decreased from approximately 120 ppb to below 20 ppb. A quick analysis based on the two month operation revealed a possible return on investment in less than a year. Keep it clean. To realize full activity from polymerization cata-

lysts, feed streams need to be as clean as possible from all contaminants and catalyst poisons. A very low level of even a single catalyst poison can reduce catalyst activity drastically. In line with the “keep it clean” philosophy, the feed treatment unit must be carefully selected to treat all possible contaminants. The money invested can give a quick return and the payback period can be completed in as little as six to nine months. HANIF POORKAR is a senior process engineer for Tasnee in Al-Jubail, Saudi Arabia, where he has been employed for more than eight years. He has mainly worked with different polypropylene process technologies throughout his 16 year career. He holds a degree in petrochemical engineering from Dr. Babasaheb

“This is the first commercial product that destroys harmful, ground-level ozone already in the air,” said Nick Leclerc, product manager for BASF Corp. “The California Air Resources Board recognizes it as an effective ozone-reduction strategy and offers credits toward LEV III certification for car makers using the technology.” Over three million vehicles are using this technology. GRACE ACQUIRES NOBLESTAR CATALYSTS ASSETS

W. R. Grace has completed its acquisition of the assets of Noblestar Catalysts, a Qingdao, China-based manufacturer of fluid catalytic cracking catalysts, catalyst intermediates and related products used in the petroleum refining industry. “Qingdao is a leading economic center in China,” said Qingdao Bureau of Commerce Vice Director General Cong Yan during a ribbon cutting ceremony. “We welcome foreign investment, especially from companies like Grace, which can help develop our fast-growing petrochemical industry while also acknowledging environmental and safety concerns.”

0.12 Bed taken online

0.10

15

0.08 0.06

10 5 0 3/21/2012

Improved catalyst mileage

Catalyst yield Inlet CO ppm Outlet CO ppm

0.04 0.02 0.00

3/22/2012

3/22/2012

3/22/2012

3/22/2012

3/22/2012

3/23/2012

FIG. 8. Catalyst activity before and after mixed bed commissioning. Ambedkar Technological University in Lonere, India, and an MBA degree from Karnataka University in India. HAMAD AL-SHBRAIN is the polymer process engineering manager for Tasnee in Al-Jubail, Saudi Arabia. He is a chemical engineering graduate with more than 16 years of experience. SAAD AL-HARBI is an operations manager at Tasnee in Al-Jubail, Saudi Arabia. He is a chemical engineering graduate and has worked for Tasnee for more than 11 years. ABDULLAH AL-SAEED is a chemical engineering graduate from King Saud University in Saudi Arabia. He currently works as an operations manager for Tasnee in Al-Jubail, Saudi Arabia.

“The successful acquisition of Noblestar’s assets in Qingdao is another milestone in Grace’s long relationship with China,” said Grace’s CEO Fred Festa. “Our goal is for customers to look to Grace for innovative technology and industry-leading technical service, as well as a globally integrated manufacturing network that aligns with the world’s demand.” Grace expects to make additional investments at the Qingdao site for environmental, safety and manufacturing upgrades. “We have been happy and proud to be a business partner of Grace’s refining technologies business for years and we are excited to continue a business relationship with Grace in the future,” said Chao Cui, CEO of Noblestar Catalysts. Grace first established a presence in China when it founded Grace China Ltd. in 1986 as one of the first foreign-owned companies to do business in the country, through its can sealants plant in Shanghai. Currently, Grace operates five manufacturing facilities, three sales offices and two technical service centers in mainland China, including its Asia-Pacific regional headquarters in Shanghai. HYDROCARBON PROCESSING | CATALYST 2013

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Stimulate the heart of your hydroprocessing unit ImpulseTM, the catalyst technology that combines the stability you recognize with the activity you need 4JOHMF
TPVSDF
*40

t
*40

t
0)4"4
 www.axens.net Select 51 at www.HydrocarbonProcessing.com/RS

AXENS

THE PERFORMANCE IMPROVEMENT SPECIALISTS Axens is recognized as a worldwide technology benchmark for clean fuels production, conversion solutions, aromatics and olefins production and purification. The combination of the technology and services with the catalysts and adsorbents manufacturing and supply business is an efficient organization that handles market needs as a single source. Currently, Axens offers a complete product range of hydrotreating and hydroconversion catalysts from naphtha and gas oil to residue applications and continues to launch new catalysts to meet high conversion and mild hydrocracking unit’s objectives and to produce ultra-low sulfur diesel (ULSD) while maximizing refinery profits.

HRK, HDK AND HYK SERIES HYDROCRACKING CATALYSTS Axens’ commercial hydrocracking catalyst suite upgrades a wide range of heavy feedstocks to produce the desired slate of products while meeting ultimate quality targets. It relies on a combination of catalysts derived from HRK, HDK Series (both for pretreating section) and HYK Series (for hydrocracking section) depending upon operator conversion targets. The combination of HRK, HDK and HYK Series enables to squeeze more middle distillates from heavy ends while reaching high conversion levels.

SULFUR RECOVERY CATALYST Over the last 30 years, continuous Axens R&D efforts contributed to a drastic change in the recovery of sulfur. Axens pioneered the use of titanium dioxide catalysts to boost the efficiency of the Claus unit, and our low temperature catalysts allow major energy and costs savings on the Claus tail gas treatment units, and even a revolution in the way these units are designed. Axens brings essential support to its customers and offers a broad range of innovative products in the field of sulfur recovery ranging from Claus alumina (CR) or boosted alumina (CR-3S) for improved performance, pure titanium dioxide catalysts (CRS 31, CRS 31 TL) and BTX management catalyst (CSM 31) to Claus tail gas hydrogenation catalysts. Our low temperature Claus tail gas treatment (TGT) CoMo catalysts (TG 107 and TG 136) allow lowering the operation temperature while maintaining TGT unit performances. As a result, energy consumption and related CO2 emissions are reduced and increase catalyst service life is increased.

IMPULSE CATALYST TECHNOLOGY With the improvement in fuel product specifications and increased demand for middle distillates, hydrotreating catalyst technology has become crucial to the refining industry. Axens has recently introduced onto the market Impulse™, a new catalyst technology leading to a complete, high performance range of hydrotreating catalysts combining stability and high level of activity. Impulse catalysts allows for even more difficult feedstock to be processed without cycle penalty. They offer, higher flexibility, maximum throughput with higher end boiling point and longer cycles. Impulse achieves increased activity by maximizing the amount of catalytically most active mixed Mo (Ni)/Co sites.

SPONSORED CONTENT

CONTACT INFORMATION 89, bd Franklin Roosevelt - BP 50802 92508 Rueil-Malmaison - France [email protected] www.axens.net

HYDROCARBON PROCESSING | CATALYST 2013

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building blocks love strong foundations Today’s petrochemical industry provides the building blocks for a wide range of materials. As the global leader in catalysis, BASF provides a strong foundation of product and process innovations across the petrochemical value chain. The result is a broad petrochemical catalyst and adsorbent portfolio backed by dedicated customer and technical service and enabled through the strength of BASF - The Chemical Company. At BASF, we create chemistry for a sustainable future.

Request your free chemical value chain wall poster at www.catalysts.basf.com/catalog Select 70 at www.HydrocarbonProcessing.com/RS

BASF

BASF—SERVING THE HYDROCARBON PROCESSING INDUSTRY FOR OVER 120 YEARS From our invention of the contact process for sulfuric acid production in 1888 to the industrial catalytic process for ammonia in 1913 to commercialization of the three-way catalytic converter in 1974, BASF has played a vital role in shaping the history of industrial catalysis. As the global leader in this field, we continue to revolutionize the hydrocarbon processing industry and advance chemical applications of all kinds, helping drive our customers’ success. As a catalyst company, BASF can only be successful by working with strong partners. We provide a comprehensive and cost-competitive portfolio of leading catalytic technologies, which is constantly improved through innovation. We use highly efficient platform technologies and high-throughput catalyst-screening methods. This, in combination with dedicated R&D projects in all regions, enables us to develop innovations quickly, in close collaboration with our customers. At BASF, we understand that the total value of our solutions must extend beyond the product itself to include commercial and technical support, global supply chain management and responsive customer service. We’ve built a commercial and technical team that has more than 300 years of combined

First definition of catalysis by Jöns Jakob Berzelius

Phinesse™ introduced as part of the REAL program offering lower rare earth

BASF develops the “contact process” for producing sulphuric acid

Rare Earth Alternative (REAL) Solutions introduced to cope with high rare earth oxide price environment

Adsorbents Chemical Catalysts ■ Custom Catalysts ■ Environmental Catalysts ■ Polyolefin Catalysts ■ Precious Metal Services ■ Refinery Catalysts ■ ■

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Alwin Mittasch (BASF) finds the industrial catalyst for production of ammonia by systematic tests of several thousand catalyst formulations

experience in the development and application of catalyst technologies across the entire range of chemical processes. This expertise is further strengthened by a global team of customer service providers, global centers of manufacturing excellence, and the capability of the global BASF supply chain. BASF is committed to partnering with its customers to understand the challenges faced by the global hydrocarbon processing community and we are proud to continue our proven tradition of introducing sustainable solutions to the market. The result is a broad catalyst portfolio backed by dedicated customer and technical service and enabled through the strength of BASF—The Chemical Company. At BASF, we create chemistry for a sustainable future.

Fritz Haber, Carl Bosch, and Alwin Mittasch implement the industrial scale ammonia process

BASF and Dow announced the world’s largest commercial-scale propylene oxide (PO) plant and the first based on the innovative hydrogen peroxide to propylene oxide (HPPO) technology

Franz Fischer and Hans Tropsch develop the catalytic hydrogenation of carbon monoxide to liquid hydrocarbons (gasoline)

BASF received Frost & Sullivan Award for Technology Leadership in recognition of Distributed Matrix Structures (DMS) and NaphthaMax® FCC catalyst innovation

Eugene Houdry develops fluid catalytic cracking (FCC)

The three-way catalytic converter for removing pollutants from gasoline engine exhaust is commercialized by BASF

1st refining catalyst application (Platinum reforming catalyst for Sinclair Oil Catforming process)

1st FCC Catalyst produced

Contact Us Americas [email protected] Asia Pacific [email protected] Europe, Middle East & Africa [email protected] www.catalysts.basf.com

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“Understanding surface morphology gives me an edge in my lab and in my garden.”

Meet Josiane Ginestra: Vegetable Grower. Catalyst Maker. Customizing the surface of soil or a catalyst is equally critical when you want performance. No one understands that better than Josiane Ginestra, CRITERION’s leading alumina scientist. Her 15+ years working within Shell’s catalyst division has led to the development of a new generation of resid hydro-conversion catalysts. Today you’ll find Josiane overseeing the Canadian oil sands refining processes, continuing to help develop custom solutions by reevaluating (and resurfacing) existing ones.

Leading minds. Advanced technologies.

www.CRITERIONCatalysts.com Select 55 at www.HydrocarbonProcessing.com/RS

CRITERION

THE NEXT BIG CHALLENGE UNLOCKING VALUE FROM UNCONVENTIONALS DAVID SHERWOOD, Senior Principal Scientist, Criterion Catalysts and Technologies

Unconventional heavy resources such as tar sands, shale oil and heavy crude oil have yet to see extensive use in the marketplace because of the processing challenges that they introduce. Meanwhile, the reserves of conventional, easy-to-process crude oils continue to decline. One of the next big challenges for the catalyst industry is to enable these ultraheavy oils to be processed with equivalent ease to conventional crudes. The issue with these unconventional resources is not only that they are heavy, viscous and often acidic, but that they also contain elevated amounts of contaminants such as nickel, vanadium and nitrogen. Although the industry is used to handling these elements, unconventional resources typically have them at unprecedented high levels. They also contain contaminants, such as arsenic, that have not historically been an issue in refineries. In addition, unconventional heavy crudes contain fewer valuable distillates and their processing leaves more less-valuable residue, which means that a greater conversion ability is required to transform the residues into the desired product slate. This ties in with a second challenge faced by the catalyst industry: how to increase the overall conversion of the residues of all crudes into more valuable distillates. Worldwide, the demand for fuel oil is projected to decrease, whereas the demand for clean transportation fuels is projected to increase. We need to get much more out of the bottom of the barrel for both conventional and unconventional crudes.

SPONSORED CONTENT

Major focuses for Criterion’s residue upgrading teams are the development of catalysts that can handle the extremely high levels of contaminants from the residues of unconventional crudes and increasing the production of valuable distillates from all crudes. The role of catalysis in residue upgrading cannot be overstated. At Criterion, we have had a very strong focus on this for almost 30 years, secured several patents and become a world-leading supplier of ebullated-bed catalysts. We will continue to target our research at developing catalysts that can facilitate the processing of heavier feeds and help to raise residue conversion levels without sacrificing cycle length.

CONTACT INFORMATION 910 Louisiana Street, 29th Floor Houston, TX 77002 Phone: +1 713 241 3000 [email protected] www.criterioncatalysts.com

HYDROCARBON PROCESSING | CATALYST 2013

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Select 76 at www.HydrocarbonProcessing.com/RS

GRACE CATALYSTS TECHNOLOGIES/ADVANCED REFINING TECHNOLOGIES

LEADING WITH THE BROADEST PORTFOLIO OF FCC AND HYDROPROCESSING CATALYSTS Only Grace Catalysts Technologies (Grace) and Advanced Refining Technologies (ART) offer refiners the broadest portfolio of high technology, state-of-the-art FCC catalysts and additives and hydroprocessing catalysts. Driven by our world-class research and development, together Grace and ART have the manufacturing flexibility and technology to customize products designed specifically to optimize FCC and hydroprocessing operations.

FCC: INVESTING IN THE GROWTH REGIONS Grace is the global leader for FCC catalysts and additives, offering solutions-oriented approaches backed by a broad, highly differentiated portfolio and industry-leading technical service. Grace’s research leadership and flexible manufacturing system support value-added technology tailored to meet customers’ current and future needs. Management of rare-earth inflation has been a major focus of the RT business over the last several years. Due to our ongoing commitment to Research and Development, Grace was able to launch eight no or low rare-earth FCC catalyst and additives in early 2011 and saw rapid market acceptance of the new technologies. The market has since experienced a similar rapid decline in rare-earth price, in response to new supply of the light rare-earth elements, and weak demand as a result of substitution efforts in both FCC and other industries. FCC customers are beginning to shift back to higher rare earth in their catalyst formulations; however many customers choose to remain on the low rare-earth formulations due to the realized yield benefits. In May 2012, Grace announced its intent to form a joint venture with Al Dahra Agricultural Company to build and operate a fluid catalytic cracking catalysts and additives plant in the Middle East. The plant in Abu Dhabi would be the first FCC catalysts and additives plant in the region, and is an important step for Grace to reinforce the reliable and timely distribution of FCC catalysts and additives to refineries in the region. There are expected to be 16 FCC units built in these regions in the next five years, which would increase the catalyst opportunity in the region by approximately $150 million. In May 2012 Grace announced the acquisition of Noblestar Catalysts Co., Ltd, a manufacturer of fluid catalytic cracking (FCC) catalysts, catalyst intermediates, and related products used in the petroleum refining industry located in Qingdao, China. This acquisition provides Grace with immediate, local manufacturing capacity to better serve our refining customers within China and North Asia.

SPONSORED CONTENT

HYDROPROCESSING: FOCUS ON R&D LEADS TO FOUR NEW CATALYSTS FOR 2013 Advanced Refining Technologies (ART), the joint venture between Chevron Products Company and Grace, was created to develop state of the art hydroprocessing catalysts. ART combines Grace’s material science, manufacturing, marketing and sales strength with Chevron’s extensive experience operating its own refineries and leadership in technology, design and process licensing. ART continues to grow and plan for the future, having recently introduced four new catalysts. For distillate hydrotreating, there is 545DX for enhanced ULSD run length and 425DX for enhanced ULSD performance at low pressure. For fixed bed resid hydrotreating units, ICR 173 is a new deep MCR and S conversion catalyst, and HCRC™ (High Catalytic Resid Conversion) Technology delivers higher catalytic and lower thermal conversion for ebullating bed resid hydrocracking units. In addition, ART is expanding and upgrading its plants to produce these new catalysts and future ones that are currently under development. These products are being designed to help refiners meet the challenges of the expected growth in demand for diesel and other low-sulfur fuels, while continuing to further upgrade the bottom of the barrel. In addition to its manufacturing facilities in the United States and Japan, ART serves the Middle East/Arabian Gulf region through a key toll manufacturing partnership with Kuwait Catalyst Company, in which it also owns an equity stake.

CONTACT INFORMATION 7500 Grace Drive Columbia, MD USA 21044 Fax: +1 410.531.4540 www.grace.com; www.artcatalysts.com

HYDROCARBON PROCESSING | CATALYST 2013

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Stepping up performance – next generation BRIM™ technology W WW.T OPSOE.CO M

Are you looking to step up plant performance? Topsøe’s next generation BRIM™ catalysts offer refiners the opportunity to increase performance through an increase in catalyst activity. Using the original BRIM™ technology Topsøe has developed several new catalysts, resulting in higher activity at lower filling densities. The next generation BRIM™ catalysts display -

high dispersion high porosity high activity

We look forward to stepping up your performance!

Select 103 at www.HydrocarbonProcessing.com/RS

HALDOR TOPSØE

CATALYSING YOUR BUSINESS Through half a century’s dedication to heterogeneous catalysis, Topsøe has developed and strengthened its position as a leading market player in catalysts, and technologies for process design. Topsøe’s markets include oil refineries, chemical plants and the energy sector, where the catalysts and technologies ensure smoothrunning and cost-efficient operations with optimal production results.

HYDROPROCESSING WORLDWIDE Topsøe has developed process design and catalysts for virtually all areas of hydroprocessing and the catalysts and technologies are in operation in plants worldwide. Topsøe’s hydroprocessing expertise offers integrated solutions including reactor internals, grading material, catalysts, process design and detailed reactor engineering. The supply of catalysts and technology offers clients a single point of expertise and responsibility. In the design of new hydroprocessing units, Topsøe’s research and test facilities offer clients testing opportunities including detailed feedstock and process analyses, which form the basis of tailor-made solutions.

TOPSØE’S REFINING COMPETENCIES Through extensive hydroprocessing research and development Topsøe offers – a broad hydroprocessing catalyst portfolio and tailor-made technologies for revamps and grassroots units meeting all specific needs of the refiner – in-depth knowledge of hydroprocessing reactor fluid dynamics and in-house developed designs for reactor internals ensuring efficient catalyst usage – more than 20 years of experience with graded bed catalyst design based on particle size, shape, void and catalytic activity for pressure drop abatement

RESEARCH BASED CATALYSTS AND TECHNOLOGIES A fundamental understanding of catalyst behaviour at the nano scale enables Topsøe to continuously develop new and improved products to meet clients’ needs. One recent development was Topsøe’s BRIM™ catalyst preparation technology, which has led to a whole new generation of unmatched activity hydrotreating catalysts with great stability.

RENEWABLES FUEL Topsøe has developed hydroprocessing catalysts and technology for processing a wide range of renewable feedstocks to gasoline, jet and diesel. Feedstocks include vegetable and animal oils, fatty acid methyl esters, waste oils and greases, tall oil and other forest waste products, algae and plastics. These feeds can be converted to transport fuels, either in stand-alone plants or by co-processing with normal refinery feedstocks.

RELATED INDUSTRIES Topsøe’s refining experience extends to related industries offering solutions for hydrogen supply, sulphur management and NOx emission. Efficient hydrogen technology and catalysts from Topsøe ensure optimised processes with low energy consumption to capacities from 5,000 to more than 200,000 Nm3/h hydrogen. Topsøe’s WSA and SNOX™ technologies remove sulphur and nitrogen oxides from flue gases, recover the sulphur oxides as concentrated sulphuric acid and reduce the nitrogen oxides to free nitrogen. The SNOX™ process is particularly suited for purification of flue gas from combustion of high-sulphur petcoke and other petroleum residues such as heavy fuel oil and tars as well as sour gases. Topsøe’s SCR (Selective Catalytic Reduction) DeNOx process is the most efficient process for removing nitrogen oxides from gases and is suitable for treating off-gases from a wide range of different industries and applications including fossil-fuel and biomass fired utility boilers, gas turbines, oil refining and chemical plants, stationary diesel engines and waste incinerators.

MARKET EXPERIENCE Topsøe has extensive market experience with all aspects of hydrotreating ranging from naphtha to heavy residue. More than 200 hydrotreating units have been licensed using Topsøe hydrotreating technology of which a large number are designed for production of ultra-low sulphur diesel with less than 10 wt ppm sulphur. Topsøe has more than 180 references in operation or projected for the production of ultra-low sulphur diesel having less than 50 wt ppm sulphur, corresponding to 5 MMBPD. 150 of these references use catalysts produced with Topsøe’s BRIM™ technology.

SPONSORED CONTENT

CONTACT INFORMATION Nymoellevej 55, DK-2800 Lyngby, Denmark Phone: +45 4527 2000 Fax: +45 4527 2999 [email protected] www.topsoe.com

HYDROCARBON PROCESSING | CATALYST 2013

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For seven decades we've helped feedstock processors enhance profits with with responsive—and responsive—and responsible—recovery responsible—recovery and refining and refining of PGMs of PGMs from from spent spent hydrocracking hydrocracking catalysts. catalysts. Tell us what we can do for you at sabinmetal.com Tell us what we can do for you at sabinmetal.com

Profitable hydrocracking/hydrocarbon processing starts and ends with Sabin Metal worldwide Select 68 at www.HydrocarbonProcessing.com/RS

SABIN METAL GROUP OF COMPANIES

PRECIOUS METALS RECOVERY AND REFINING FOR GLOBAL INDUSTRY The Sabin Metal Group of Companies recovers and refines PGMs (platinum, palladium, ruthenium, rhodium), plus rhenium, gold, silver, and other precious metals from spent hydrocarbon, chemical, and petrochemical processing catalysts with zeolite, soluble and insoluble alumina, silica-alumina, and carbon supports. The Sabin Metal Group is composed of five independent organizations including Sabin Metal Corp., Scottsville, New York, considered the most sophisticated facility of its kind for safely processing precious metal-bearing materials; Sabin Metal West, a specially equipped facility for sampling large lots of precious metal-bearing spent hydrocarbon processing catalysts. This refinery employs electric arc furnace (EAF) technology which helps maximize recovery of precious metals, and also incorporates a unique “low dust” continuous sampling system for accurate sample derivation and total environmental safety and compliance. Sabin Metal Europe B.V., a technical service division based in Rotterdam, works with hydrocarbon, chemical, petrochemical, and nitric acid processors in Europe, Africa, and the Middle East, to recover and refine precious metals from spent catalysts and nitric acid production equipment and facilities. Sabin International Logistics Corp. (SILC), is a licensed hazardous waste, hazardous materials, and general commodities transporter providing global transportation and logistics support for spent precious metal-bearing catalysts and other materials. The company operates its own fleet of trucks, and is also a Permitted and Licensed Freight Broker. SMC (Canada) Ltd., the McAlpine Mill in Cobalt, Ontario, Canada, offers capabilities and processing technologies to extract highest possible metal values from residual materials generated in refining, smelting, and milling operations. The Sabin Metal group of companies is the largest domestically owned, independent precious metals refining organization in North America. The company’s recovery/refining facilities and sales/service offices are located in strategic countries around the world. Sabin’s gold, silver, platinum, and palladium are accepted on NYMEX/COMEX (Chicago Mercantile Exchange); Sabin’s platinum and palladium are also accepted for delivery on the London/Zurich market and by the London Platinum and Palladium market (LPPM). The organization is entering its seventh decade of working with a worldwide customer base by providing added value services along with the peace of mind that comes from working with an environmentally responsible precious metals refiner. There are many criteria to consider when evaluating, selecting, and working with a precious metals refiner to recover precious metals from spent hydrocarbon processing catalysts. Essentially, your refiner must provide highest possible returns for PGMs from spent catalysts with rapid processing turnaround time. Another key issue—perhaps most important—is that the refiner complies with applicable environmental standards concerning process effluent disposal or atmospheric discharge at its facilities.

SPONSORED CONTENT

Sabin’s continuous catalyst sampling system helps assure maximum returns of precious metals from spent hydrocarbon processing catalysts.

CONTACT INFORMATION 300 Pantigo Place, Suite 102 East Hampton, NY 11937 Telephone 631-329-1717 Fax 631-329-1985 [email protected] sabinmetal.com

HYDROCARBON PROCESSING | CATALYST 2013

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No Harm,

NO FOUL Protect your catalyst bed from being fouled with new MacroTrap® XPore 80 guard bed media MacroTrap® XPore 80 guard bed media offers increased macroporosity which ultimately increases filtration capacity. MacroTrap® XPore 80 media’s proven technology traps particulate contaminants for pressure drop stability and improved reactor cycle performance – so catalyst bed fouling is sidelined. Don’t let contaminants foul your reactor. Protect your catalyst investment with MacroTrap® XPore 80 guard bed media from Saint-Gobain NorPro, the leading name in guard bed media.

Magnification of Macropores

Saint-Gobain NorPro www.norpro.saint-gobain.com Ohio, U.S.A. Tel: +1 330 677 3552 [email protected] Steinefrenz, Germany Tel: +49 6435 9657 0 [email protected]

Select 80 at www.HydrocarbonProcessing.com/RS

SAINT-GOBAIN NORPRO

CERAMIC MEDIA TECHNOLOGY FOR IMPROVED FIXED BED REACTOR PROCESSING Saint-Gobain NorPro has served the refining, petrochemical/chemical, environmental, and oil and gas production industries for more than 100 years, providing technology-driven ceramic solutions for process and manufacturing challenges. The company offers product innovations for fixed bed reactor processing, heat and mass transfer applications, drilling and exploration. We are the leading supplier of custom catalyst carriers, bed support media and proppants. Ceramics play a key role in catalysis, and Saint-Gobain NorPro continues to build on its impressive collection of engineered ceramic media and shapes.

GUARD BED MEDIA Processing a variety of feedstocks presents unique challenges for today’s refiners. Maintaining desired unit cycle length is critical. As a result, the catalyst must be protected from contaminants that can cause fouling, leading to increased pressure drop. Saint-Gobain NorPro leads the industry with its MacroTrap® guard bed media—technology that “traps” contaminants. Use of the MacroTrap media leads to improved operational stability. The key to MacroTrap® guard bed media filtration capability is the tortuous path created by the internal macropore structure that efficiently “traps” particulates inside the body of the ceramic media, thus extending unit performance. MacroTrap® XPore 80 is the latest in guard bed media protection. MacroTrap® XPore 80 media has increased macroporosity with >80% internal void capacity, resulting in greater “trapping” capacity.

BED TOPPING MEDIA Saint-Gobain NorPro’s recent innovation in bed topping media is NGBT® (Next Generation Bed Topping). The NGBT® media’s shape offers substantial gains in void fraction and surface area when compared to conventional sphere-shaped media, Raschig rings or other competing media. The immediate benefit of NGBT® media is enhanced flow distribution, which in turn facilitates uniform distribution of feedstock.

BED SUPPORT MEDIA Bed support and bed topping media work to support the catalyst bed and also act as hold down layers. Denstone® bed support media continues to be the industry standard for well over 60 years. The new non-spherical media, known as Denstone® deltaP®, improves operating performance through reduced pressure drop, and offers cost savings through fewer layers of media required—which in turn increases catalyst loading in the reactor.

SAINT-GOBAIN Saint-Gobain NorPro is a wholly-owned subsidiary of Compagnie de Saint-Gobain, a multinational corporation with headquarters in Paris. Saint-Gobain transforms raw materials into advanced products for use in our daily lives, as well as developing tomorrow’s new materials.

SPONSORED CONTENT

CONTACT INFORMATION Phone: 330-673-5860 [email protected] www.norpro.sant-gobain.com

HYDROCARBON PROCESSING | CATALYST 2013

C–93

Bulgaria: Cat Tech Services (South Eastern Europe) Ltd. Vetrino, 9220, 25 Georgy Rakovski Str. Phone: +35 9527 46770 Fax: +35 9527 46770 Select 65 at www.HydrocarbonProcessing.com/RS

Refining Developments L. MCDERMOTT, Applied Instrument Technologies, Inc., Upland, California; and A. MALIK, Phillips 66, Roxana, Illinois

Improve benzene control The US Environmental Protection Agency (EPA) issued the Mobile Source Toxics (MSAT) Rule on February 26, 2007 (40 CFR Parts 59, 80, 85 and 86). This rule required refiners to reduce the average benzene concentration in the gasoline pool to 0.62 vol% or less by January 1, 2011. Refiners producing higher levels could reduce their concentration by purchasing credits, but they are capped at a maximum actual average concentration of 1.3 vol%. The MSAT ruling has led many refineries to upgrade processing to meet the new benzene limits. Because the majority of benzene in blended gasoline (by some estimates 70%–85%) originates from reformate, reducing the benzene levels in the reformate offers the most direct means to meet specifications for finished gasoline.1 The refinery. The Phillips 66 Wood River refinery in Roxana, Illinois, is a 365,000-bpd facility. In 2008, the refinery (a ConocoPhillips refinery at the time) evaluated options to meet the MSAT ruling. The plant decided to install a reformate splitter column to treat reformate produced by the refinery’s catalytic reformer. This technique is a post-fractionation process. Benzene is concentrated in the heart cut of the splitter containing 20%–30% of benzene. The benzene can be removed from the blending pool either by saturation or by additional concentration via extraction to produce a sellable high-purity benzene product.

To meet the required cycle time for the broad dynamic range of required measurements, four PGC systems would be required. • Cost. By using one FTIR system with two measurement flow cells (one for high-concentration analysis streams and one cell for measurement of low-concentration streams), the FTIR system cost was lower than that to install four PGCs. Additional cost savings for the FTIR system were calculated based on lower maintenance and utility requirements. To measure, control and optimize the reformate splitter column, Phillips 66 installed an online multistream FTIR analyzer to monitor six streams of the unit. FIG. 1 is a diagram of the splitter column, highlighting the analyzed streams. The light-reformate feed, heavy-reformate feed, side-draw feed to the benzene extraction unit, lower side-draw, tower bottoms and splitter top streams are analyzed by a two-cell FTIR system installed in a shelter near the unit. The FTIR system was specified to monitor the benzene, toluene and butane concentrations. The analysis slate was modified to add the analysis of distillation points, aromatics

Selection criterion. To optimize performance of the benzene-

reduction process and to ensure that the reformate complied with new blending pool specifications, the Wood River project team evaluated several analyzer technologies for real-time analysis. The refiner selected two technologies: process gas chromatography (PGC) and Fourier transform infrared (FTIR) spectroscopy. The technology selection phase focused on: • Ability to make the required analytical measurements. Both the PGC and FTIR systems were capable of measuring benzene and toluene at the required levels. • Site familiarity with technology and ability to support systems. The Wood River refinery has numerous PGC systems installed throughout the refinery, and it also had installed spectroscopic near-infrared (NIR) analyzers on the gasoline blenders since 2001. The site recently added online NIR analysis of the diesel blending stream. • Cycle time. Spectroscopic analyzers like the installed FTIR system have typical cycle times of 1–2 minutes per stream.

Splitter tops Light feed

RDC-1 feed

Toluene concentrate

Heavy feed

Heavy reformate product

FIG. 1. Benzene-extraction unit indicating analyzed streams. Hydrocarbon Processing | MARCH 2013

95

Refining Developments and research octane number (RON) for the toluene concentrate stream. Measuring butanes was not developed because of the difficulty in reliable sample handling and also the determination that there was limited value in using the butanes data in the unit control scheme. Analyzer system. A dual-stream, extended-range mid-infrared process FTIR analyzer was installed on the splitter column. The six incoming streams were plumbed to the analyzer shelter. Sample conditioning systems were installed on the exterior wall of the shelter to remove particulate materials and water from the streams. The conditioning system also provided temperature control of the sample. A stream switching manifold is used to select the streams to be analyzed. The analyzer system is linked to the distributed control system (DCS) through a Modbus TCP connection. The DCS system is used to activate streams, allowing the system to only measure required streams. The analyzer automatically cycles through selected active streams. During routine operation, the two cells are alternatively analyzed. While one analysis cell is being measured, the second cell is flushed Analyzer shelter Class 1, Group D, Division 2, T2c

Modbus

Heavy reformate Light reformate feed

RDC feed Splitter tops

Sample preconditioning enclosure

Toluene concentrate

Sample preconditioning enclosure

Heavy reformate feed

Plant DCS system FTIR analyzer – optical and electronic enclosure Laboratory LIMS FTIR dual-cell enclosure

Lab FTIR analyzer Wash/validation skid Wash

Val

out with the next stream to be analyzed. The system was configured with a pneumatically actuated optical switch controlled by the main analyzer control program, which automates system data collection and analysis, cell selection, validation, automatic background collection and data reporting. An automated sample-cell wash and validation system was installed to confirm proper system performance, and it allows for automatic sample-cell washing in the event of window fouling. FIG. 2 is a block diagram of the analyzer system. Calibration. A laboratory FTIR system was installed in the

Wood River laboratory shortly after the decision was made to install the online analyzer. During the time that the process analyzer was being built and installed, all routine samples for the six streams were collected and taken to the lab for analysis by conventional analyzers to determine the property values of interest. The samples were also scanned on the lab FTIR system. When the operators ran the samples on both the lab and conventional analyzer systems, the laboratory information management system (LIMS) number was entered, allowing for easy integration of lab data with the FTIR scans. The FTIR system logged collected data and sample information into a file that was used in PC software to create a calibration database. The laboratory system uses the same optics, interferometer and sample analysis cell as used in the online system. The extremely high accuracy of FTIR-based systems allows for seamless transfer of data and models between systems; it ensures that the models developed based on samples scanned on the lab system can be used directly online. The FTIR spectra were collected from 6,000 cm–1 to 1,000 cm–1 (1,667 nm–10,000 nm) using a 0.5-mm pathlength transmission cell. Representative calibration spectra for the six streams are shown in FIG. 3. The spectra for the different streams are colored as indicated in the legend. The large differences in the spectra for the different streams are due to the large compositional differences for the six streams. Initially, it was thought that it would be possible to combine data from several streams to minimize the modeling effort. A principal components analysis (PCA) was performed on the combined data set to investigate any relationships between the streams. FIG. 4 is a score plot of the first two factors. The six sample streams

FIG. 2. FTIR installation diagram. Heavy feed RDC-1 Toluene concentrate Heavy reformate Splitter tops Light feed

Factor 2 (10.5%)

2

0

-2

-5

FIG. 3. FTIR spectra for calibration samples.

96MARCH 2013 | HydrocarbonProcessing.com

0 Factor 1 (80.5%)

FIG. 4. Score plot for calibration samples.

-5

Refining Developments TABLE 1. Calibration table summary Concentration range Stream

Property

APC Use

Units

Min

Max

Samples

Factors

SEE

SECV

R2

BRU heavy feed

Benzene

Monitor

vol %

2.65

7.28

60

3

0.192

0.208

0.978

BRU heavy feed

Toluene

Monitor

vol %

12.2

21.1

58

4

0.63

0.7

0.914

BRU rdc-1 feed

Benzene

Monitor

vol %

25

40.5

85

4

0.77

0.81

0.944

BRU rdc-1 feed

Toluene

Control

vol %

0.02

2.12

76

4

0.121

0.125

0.946

BRU toluene concentrate

Benzene

Monitor

vol %

0.02

1.85

93

4

0.07

0.083

0.92

BRU toluene concentrate

Toluene

Monitor

vol %

35

67.45

102

4

1.49

1.54

0.918

BRU toluene concentrate

RON #

Control

ON

99.65

113.5

100

5

0.736

0.787

0.916

BRU toluene concentrate

Distillation 50%

Control

°F

235.7

257.5

98

3

1.05

1.1

0.941

Toluene

Control

vol %

3

37.65

108

3

1.35

1.39

0.972

BRU heavy reformate product BRU splitter tops

Benzene

Control

vol %

0.23

2.04

56

4

0.1

0.107

0.97

BRU light feed

Benzene

Monitor

vol %

10.95

18.5

90

3

0.5

0.52

0.926

BRU light feed

Toluene

Monitor

vol %

20.76

29.76

87

4

0.736

0.776

0.82

RDC feed

2.5

Distillation point predicted, 50%

Toluene predicted, %

2.0

255

250

1.5

245

1.0

240

0.5

0.0 -0.5 0.0

0.5

1.0 1.5 Toluene measured by lab, %

2.0

110

30

Toluene predicted, %

35

RON predicted

108

106 104 102

100 102

104 106 108 RON measured by lab, %

245 250 Distillation point measured by lab, %

110

112

255

260

Heavy reformate

40

112

100

240

FIG. 7. Correlation between model-estimated and lab-determined 50% distillation point for toluene concentrate.

Toluene concentrate

114

235 235 235

2.5

FIG. 5. Correlation between model-estimated and lab-determined toluene for RDC feed.

98 98

Toluene concentrate

260

114

25 20 15 10 5 0

0

5

10

15 20 25 Toluene measured by lab, %

30

35

40

FIG. 6. Correlation between model-estimated and lab-determined RON for the toluene concentrate stream.

FIG. 8. Correlation between model-estimated and lab-determined toluene for heavy reformate.

can be seen to be different enough in score spacing to discourage efforts to combine sample streams into any grouped or common models. Calibration models were developed for each stream, and the results were significantly better than the results obtained when data from similar streams was combined. Calibration models were developed using chemometric modeling software.a This software uses the principal compo-

nents regression (PCR) algorithm and incorporates advanced data-filtering techniques to reduce or eliminate interferences such as those caused by ambient carbon dioxide and moisture, baseline effects and possible pathlength variations. The software was evaluated to investigate additional data processing and calibration capabilities. The spectral regions that are saturated or which contain little or no relevant information for Hydrocarbon Processing | MARCH 2013

97

Refining Developments the property of interest, were excluded from the calibration. All calibrations were developed using random subset cross validation. Models were developed generally following the recommendations of ASTM E-1655, Standard Practices for Infrared Multivariate Quantitative Analysis. A minimum number of “outlier” samples were identified and removed from the calibration sets. Several of the removed outliers were based on incorrect sampling; others were identified as misidentified samples. TABLE 1 summarizes the results of modeling. For all measurements examined, a strong correlation (R2) was attained between the spectral response and the primary values, and the Splitter tops

2.5

Benzene predicted, 50%

2.0 1.5 1.0

0.5

0.0 0.2

standard error of cross validation (SECV) essentially matched the standard error of estimate (SEE). All models used a significant number of calibration samples, ranging from 56 to 108 samples. The number of factors used for all models varied from 3 to 5. The optimum number of factors was selected by evaluating the predictive residual error sum of squares plots, loadings and regression vectors calculated to avoid “over-fitting” of the data. Streams listed as “control” in TABLE 1 are used for advanced process control (APC) of the splitter column. FIGS. 5–9 show the strong correlation (fit) obtained for several of the key control properties used.

0.4

0.6

0.8

1.0 1.2 1.4 Benzene measured by lab, %

1.6

1.8

2.0

FIG. 9. Correlation between model-estimated and lab-determined benzene for splitter tops.

CPPAC Cen trif u gal P um p Pac

2.2

Control. The FTIR analyzer provides product stream analysis that is used in an APC application to maximize recovery, and to reduce energy at the benzene-removal unit (BRU) column. The controlled variables of this application are: • Percent benzene in the overhead liquid product • Percent toluene in the feed to the extraction unit • RON of the toluene concentrate stream • 50 % Distillation point of the toluene concentrate stream • 5% Distillation point of the toluene concentrate stream. The other calibrated properties are used for monitoring the system and are not actively used in the APC logic. However, they are available to operators and control engineers through the DCS system. All data are archived in the plant historian. A single FTIR system has been installed to analyze six streams on a reformate splitter column. Calibration models were developed and installed to monitor the concentration of benzene and toluene, as well as additional properties, including the RON and several distillation points. The accuracies attained using the FTIR system are comparable to those attained using the conventional analytical techniques. The system was installed to meet the MSAT deadline, and it has enabled the Wood River refinery to meet the regulated MSAT specification for blended gasoline and to optimize operation of the splitter column. LITERATURE CITED Palmer, R. E., “Options for reducing benzene in the refinery gasoline pool,” 2008 NPRA Annual Meeting, San Diego. 2 McDermott, L., “Cetane benefits,” Hydrocarbon Engineering, July 2003, Vol. 8 No. 7, pp. 69–73. 3 “Cutting through the haze: How will refiners meet EPA’s new benzene standards?, ”Benchmark Newsletter, 2007. 4 Uvland, K. A., Oil and Gas Process International, 1997 5 Vickers, G. H., “On-line determination of naphtha properties to control a refinery process using near infrared spectroscopy,” IFPAC Conference, 2001. 1

Centrifugal Pump Pac is a complete program designed for centrifugal pump engineers. This program provides commercially available pump designs to best suit given operating conditions and retirements, and it revises existing pump curve data for new pumping conditions. Price: $695 (Order # S140)

a

NOTES Calibration models were developed using chemometric modeling software available from Applied Instrument Technologies.

DR. ATIQUE MALIK holds a BS degree and PhD in chemical engineering from the University of Leeds. He has worked in the area of model predictive control for various companies. Dr. Malik has reviewed publications for the Journal of Process Control and has been associate editor for Control Engineering Practice. He is the team Lead for APC at the Wood River Phillips 66 refinery.

GULF P U B L I S H I N G C O M PA N Y

+1 (713) 520-4426 [email protected] www.gulfpub.com/soft

98MARCH 2013 | HydrocarbonProcessing.com

LARRY MCDERMOTT holds a BA degree in chemistry, an MS degree in marketing and technological innovation and an MS degree in management engineering. He has worked in the field of process analyzers based on spectroscopic systems for over 20 years. He is the technical applications manager for Applied Instrument Technologies, focusing on chemometric modeling of refinery process streams.

Clean Fuels C. POUWELS, Albemarle Catalysts, Amsterdam, The Netherlands, K. BRUNO, Albemarle Corp., Houston, Texas

FCC catalyst design evolves to maximize propylene

Understanding the mechanisms. In the refining industry, several cracking processes are applied to cleave carbon-carbon bonds: hydrocracking, thermal cracking and, most importantly,

catalytic cracking. Hydrocracking will not be discussed further as it is of nearly no importance for the production of propylene. Cracking is either an acid catalyzed or a thermal, not a catalytic process. Though FCC is predominantly a catalytic process, some thermal reactions also take place. Thermal cracking proceeds through a free radical process and involves three steps: initiation, propagation and termination. At high temperatures, radicals are formed in the initiation step. These radicals react further whereby ethylene is formed and new radicals, leading predominantly to ethylene, followed by methane and propane. Catalytic cracking, however, makes use of acid sites in the FCC catalyst and involves carbocations intermediates. Moreover, in FCC, a higher degree of branching is found, giving evidence to isomerization reactions. FIG. 1 compares the products of thermal and catalytic cracking of n-hexadecane at 500 °C.1 In thermal cracking very little branched products are found. Looking back in FCC history, some remarkable changes occurred when going from amorphous cracking catalysts to zeolite-based cracking. This step change in the FCC catalyst technology led to a drastic increase in gasoline yield. Those are all improvements that were highly desired from the FCC process. But at the same time, the gasoline research octane number (RON) dropped dramatically, which was associated with the lower olefin content of the gasoline. These yield and quality shifts were attributed to hydrogen transfer (HT) reactions, by which initially formed olefins are converted to more stable paraffins. And while gasoline olefins 140 Moles product/100 moles cracked

The growing demand for propylene has intensified interest in maximizing propylene from refinery fluid catalytic cracking units (FCCUs). Refiners in Asia and the Middle East are setting the pace, with numerous new FCC units planned or already onstream to take advantage of this opportunity. Each of these new units includes the most modern technology to achieve record high propylene and conversion. On one hand, new records are set in the heaviness of feeds posing challenges to maintain activity; conversely, some FCCUs are specifically designed for low sulfur (hydrotreated) gasoil (GO) feeds that impose a real challenge to meet the heat balance requirement of the FCCU. Achieving record high propylene yield and conversion from a wide range of feed quality offers considerable challenges to the catalyst design, as well. High accessibility (diffusion), advanced zeolite and low rare earth (RE) technologies are critical components of modern catalyst designs required to meet the desired objectives. Also, the feed composition impact and process variables on the yields and heat balance are significant and thus requires a good understanding of the chemistry to design the right FCC catalyst for individual FCCUs. Naphtha cracking played a dominant role in manufacturing propylene for a long time, but the primary product is ethylene. The growth rate of propylene demand is outpacing that of ethylene. More recently, the FCC process is often a better solution for refiners to invest in. The FCC process has become a major refinery process to generate propylene. With many new units announced and several units already in the process of engineering and construction, the importance of the FCC process to meet the growing propylene demand will increase. These new FCC units often operate at higher severity and up to the edge of what is possible today. The highest reactor temperatures and very high levels of shape selective ZSM-5 additives are almost never sufficient to meet new propylene targets, which seemingly hover at a plateau due to over-cracking, activity dilution and the absence of key catalyst design features. Many refiners are facing this plateau and have difficulty lifting the yield of propylene. A good understanding of the cracking mechanisms involved is required, as is a focus on the quintessence of FCC catalyst design to maximize propylene yield and the mechanisms of propylene generation.

Thermal cracking quartz chips Catalytic cracking amorphous SiO2/Al2O3

120 100 80 60 40 20 0

C1

C2

C3

C4

C5

C6

C7 C8 C-number

C9 C10

C11

C12

C13

C14

FIG. 1. Comparison of products observed in thermal and catalytic cracking of n-hexadecane at 500°C. Hydrocarbon Processing | MARCH 2013

99

Clean Fuels were reduced, light olefins also decreased with the change to zeolite-based cracking (TABLE 1). HT is a bimolecular reaction, whereby hydrogen is, for instance, transferred from naphthenes to olefins, producing paraffins and aromatics. Brønsted acid sites of the Y-zeolite play a crucial role in these reactions. Catalyst manufacturers have several tools to optimize the characteristics and functionality of the Y-zeolite, such as unit cell size (UCS), silica-to-alumina ratio (SAR) and content of RE and sodium. For the generation of propylene, other shape selective zeolites exist and are commonly applied. For example, the invention of ZSM-5 created another step change in FCC catalyst technology. Since then, other proprietary advanced zeolite technologies have been developed for maximum propylene applications.12 However, because it is well-documented public knowledge that ZSM-5 additives are effective to create light olefins, including propylene, the fundamental chemistry discussed here will use ZSM-5 to help illustrate the mechanisms of propylene production. Zeolite, such as ZSM-5 with a pore entrance smaller than of Y-zeolite, drives different catalytic reactions. The smaller pores restrict the access of branched and cyclic hydrocarbons into its interior and allow straight-chain and mono-methyl paraffins and olefins to enter, hereby generating predominantly propylene and also butylenes and ethylene. The reactants are particularly in the C6 and C7 range, but other gasoline molecules with a carbon atom number between 5 and 10 can also be converted by ZSM-5 to some extent. FIG. 2 shows clear evidence of the increased production of C3 , C2 and C4 , and the strong reduction in C6 and C7 components. FCC gasoline can be typically classified into five types of hydrocarbons: paraffins (P), isoparaffins (iP), olefins (O), naphthenes (N) and aromatics (A). Taking a closer look at the different types shows that isoparaffins and olefins with six or seven carbon atoms are mostly cracked by ZSM-5 (FIG. 3). When an olefin is cracked by ZSM-5, two smaller olefins are produced. The generation of C6 and C7 olefins in the primary cracking steps of FCC is therefore a valuable mechanism that needs managing. TABLE 1. The change from amorphous cracking catalyst to zeolite-based cracking leads to lower olefins Conversion

Gasoline, vol%

Propylene, vol%

Silica-alumina gel

75.5

47.5

8.5

REHY

85.5

61.0

5.9

The numbers on the bars in FIG. 3 show a reduction in the specific hydrocarbon type. The decrease in light isoparaffins is the largest, but this is speculated to be due to reduced formation of light isoparaffins in the presence of ZSM-5 and other reactions, including cracking to light olefins. The light olefins, however, contribute the most to propylene make, as olefins are more reactive and they crack to form two smaller olefin molecules. In general, gasoline-range olefins are the primary reactants for propylene. The molecules that can be readily cracked by ZSM-5 are in the gasoline boiling range and are often referred to as light-olefin precursors. Though isoparaffins are good light-olefin precursors, the small pores of ZSM-5 restrict access to only those paraffins with a methyl branch. Longer branches or multiple branches are consequently undesirable. As Y-zeolite catalyses isomerization reactions, these reactions should be carefully controlled. The right-hand end of the x-axis in FIG. 3 shows benzene and methylbenzene, which are two of the few gasoline components to increase slightly with the use of ZSM-5. FCC manufacturers have a variety of tools available for designing the optimal catalyst for maximum propylene applications, as these different catalyst components affect cracking, HT and isomerization. Y-zeolite is the component most talked about. Its acidity is high and contributes to a high level of cracking. Though Y-zeolites can be varied in terms of UCS, RE content and SAR, their HT activity is relatively high, and they also possess some isomerization activity. Lower RE content is thus strongly preferred for lowest HT. ZSM-5 zeolite is also very high in acidity and thus has high cracking activity, though the cracking rate is highly determined by the size and shape of the reactants. Its high SAR is also responsible for isomerization. Conversely, ZSM-5 has oligomerization activity, whereby propylene can be consumed. Finally, matrices play another crucial role in FCC catalysis. Matrices are predominantly known as the components responsible for pre-cracking large molecules before they enter the zeolites. A wide range of matrices that vary in functionality, such as bottoms conversion capability and metal capturing power, can be applied. With their differences in functionality, the acidity is different and can be rated as low to moderate compared to Yzeolite. Their cracking activity varies from low to high, while the HT is typically lower than that of Y-zeolite, as is the isomerization rate. FIG. 4 compares two different matrices, with different 5

18 16

FCC catalyst FCC catalyst + ZSM-5

14

Yield, wt% on feed

Yield, wt%

12 10 8 6 4

3 2 1

2 0

FCC catalyst FCC catalyst + ZSM-5

4

C1

C2

C3

C4

C5

C6 C7 C-number

C8

C9

C10

C11

C12

FIG. 2. The effect of ZSM-5 on products distributed by carbon number.

100MARCH 2013 | HydrocarbonProcessing.com

0

-2.8 iP6

-2.1 O6

-1.7 -2.0 -1.5 iP5 iP7 O7 C-numbers and component type

0.3 A7

0.5 A6

FIG. 3. Gasoline components that are most and least affected by ZSM-5.

Clean Fuels

Catalyst design considerations. With the understanding

of the chemistry and influence of operating conditions, what catalyst design should be applied for optimum propylene? In FCC, there is no generic design that fits all, since all FCC units TABLE 2. Comparing FCC key components

ZSM5-Zeolite

Acidity

Cracking

HT

Isomerization

Very high

High*

–

High

Y - Zeolite Matrices

High

High

High

Moderate

Lowmoderate

Low-High

Low to moderate

Low

*Size and shape of molecules determines cracking rates

TABLE 3. The effect of reactor temperature on propylene yield Unit 1 Reactor temp., °C

Unit 2

522

543

519

530

C , wt%

Base

+1.6

Base

+1.0

C4 /(C3+C4 ), -

0.58

0.56

0.59

0.57

= 3

*FCC catalyst with high ZSM-5 content

17

0.40

15

0.30

13

0.20 C3=: Matrix 2 C3=: Matrix 1 HTI: Matrix 2 HTI: Matrix 1

11 9 40

50

60 70 Conversion, wt%

80

HTI

Optimizing unit conditions. The FCC process is very versatile and has the flexibility to be operated in different modes which fit specific yield slates and meet market demands. To achieve maximum propylene, the FCC unit is typically operated in a high-severity mode, whereby a high conversion is achieved through high reactor temperature and high cat-to-oil ratios. These parameters can be optimized by the operators with relative ease. Other conditions that positively influence the propylene yield are low hydrocarbon partial pressure and/or high steam partial pressure, use of naphtha recycle and the choice of a hydrogen-rich feed. The latter options, however, are not always easy to control. While process conditions can elevate the propylene yield by a few percent, the choice of the catalyst or catalyst system can have a much larger impact on propylene, on the order of 6 wt%–8 wt% additional propylene on feed. A high-severity operation is most favorable for propylene production. When unit limitations permit, the highest reactor temperature is commonly applied. Increasing reactor temperature, however, also enhances thermal cracking reactions, with consequently more dry-gas production. A ratio that is commonly looked at in FCC is C4/(C3 + C4). A ratio of 0.68 denotes a high degree of catalytic cracking, whereas values near 0.50 indicate a high contribution of thermal cracking.3 (These values are applicable to FCC catalysts not designed for maximum propylene.) The effect of temperature on additional propylene make is dependent on the starting condition. In maximum propylene applications, reactor temperatures are already quite high and the amount of ZSM-5 is substantial. Under those conditions, the effect of an increased reactor temperature of 10°C modestly enhances propylene yield by about 0.8 wt%–0.9 wt%, as shown in TABLE 3. It is generally accepted that cracking processes have higher effective activation energy than bimolecular HT reactions, so raising the reactor temperature increases the tendency for gasoline olefins to crack to light olefins rather than to undergo HT.4 The impact of temperature on propylene make is evident in a thermal cracking process. Also, high severity FCC-type catalytic processes operate under much higher temperatures than FCC in order to maximize propylene.5, 6, 7 The effect of feedstock is important. Though a refiner may not have a wide choice of options, any considerations are valuable. It is generally accepted that cleaner feeds with higher hy-

drogen content are easier to crack and increase propylene yield through higher conversion. Unwanted are feeds that are high in aromatics, since these components cannot be converted to gasoline precursors and thus not lead to propylene. A potentially attractive option to consider, though, is the use of recycle streams of specific components which are valuable as reactants for propylene production, such as light naphtha recycle or the addition of products of C4/C5 oligomerization. These can either be added in the main riser or, even better, in a second riser when present. An interesting case for demonstration purposes was presented in 2007, in which Fisher Tropsch (FT) wax was cracked in a pilot plant.9 Blends in different ratios of equilibrium catalyst and ZSM-5 additive olefins and octane maximization are applied. High conversion rates are obtained due to the high crackability of FT wax. While in common FCC operation the FCC catalyst contributes to the highest conversion. With the use of pure ZSM-5 additive, conversion levels above 90 wt% are achieved with a propylene yield of 18 wt%, as shown in TABLE 5.

C3= yield, wt%

HT power. The one lowest in HT (as shown on the second Yaxis) leads to the highest C3=. Next to HT, several other secondary reactions take place, such as cyclization and aromatization reactions. Moreover, propylene molecules can undergo oligomerization reactions, which means it is possible that propylene can be consumed and its yield decreases.2 Longer residence times are detrimental for maximum propylene yields. In summary, several factors must be controlled for maximum propylene yield: 1. Maximize generation of gasoline precursors, which are predominantly C6 and C7 olefins and iso-paraffins 2. Minimize HT, which consumes valuable olefins 3. Control isomerization reactions to form methyl branches 4. Minimize other unwanted secondary reactions, such as cyclization, aromatization and oligomerization.

0.10 0.00 90

FIG. 4. Comparing the effect of different matrices on HT and propylene yield. Hydrocarbon Processing | MARCH 2013

101

Clean Fuels and operations are unique. Important is a good understanding of the specific unit, its constraints and the bigger picture of the type of operation, for instance, with respect to cleanliness or contamination of feed. ZSM-5 additives are a very effective solution for increasing propylene yield in the FCC unit. This approach gives the refiner a great deal of flexibility, as additive usage can be adjusted according to changes in propylene demand and to optimize operation within unit constraints such as wet-gas compressor loading. As explained before, ZSM-5 generates propylene by selectively cracking olefins in the gasoline boiling range. As the amount of ZSM-5 additive in the catalyst inventory increases, the incremental yield of propylene produced per percent of additive decreases. Propylene yield reaches a plateau once the ZSM-5 crystal concentration in the catalyst inventory reaches around 10%. The diminishing effectiveness of ZSM-5 at higher concentrations occurs largely because olefins in the gasoline become depleted. It is the base FCC catalyst design and technology, not ZSM5 concentration, that define the maximum propylene yield that can be achieved in the FCC unit. FIG. 5 shows the amount of propylene produced by a variety of ZSM-5 additives when used with four different host FCC catalysts. The yield response of TABLE 4. Example showing effect of hydrocarbon partial pressure on propylene yield RxT, °C

553

RxP, barg

551

2.1

2.8

Steam, mol% on feed

71%

60%

C3=, wt% on feed

11.6

9.9

TABLE 5. Cracking FT-wax in FCC pilot plant leads to high propylene yields Feed properties Type

Operations

FT-Wax

Unit

API, °

43.1

RxT, °C

SG

0.81

CTO

S, ppmw

2.5

Catalyst

Pilot plant 538 4 ZSM-5 additive olefins and octane maximization

Sim Dist, 5%, °C

368

Conv, wt%

92

Sim Dist, 95%, °C

474

C3=, wt%

18

the ZSM-5 changes according to the properties of the host catalyst, and at high concentrations of ZSM-5, the propylene yield reaches a plateau level that is determined solely by technology inherent in the base catalyst. The pore size of zeolite is too small for access by feed molecules, and so FCC catalysts contain some alumina matrix components to pre-crack the feed into smaller components. The product of this primary cracking by the matrix is very olefinic, and generates the gasoline olefin precursors that are then cracked into light olefins. We can estimate the relative activity of FCC catalyst for generating secondary reactions, using the hydrogen transfer index (HTI) for catalysts tested under constant conditions with the same feed. The HTI can be defined in several ways. One of them is the ratio of isobutane over total C4s. Catalysts with a lower HTI generate fewer secondary reactions, and more gasoline olefins are preserved for cracking into light olefins. The four base catalysts tested in FIG. 5 demonstrate the effect of HTI on propylene generation, with propylene yield increasing as the HTI of the base catalyst is reduced. Further testing under constant conditions demonstrates that suppressing HT is the key to maximizing propylene. In FIG. 6, multiple catalysts were mixed with varying amounts of ZSM-5 additive. The relative propylene shows a strong correlation with the HTI of the base catalyst. As suggested by the correlation obtained in FIG. 6, suppressing HT reactions is the key to maximizing propylene yield, by maximizing the availability of olefin precursors for cracking. There are several ways that catalyst properties can be modified to suppress HT reactions. It is acid sites density and the close proximity of these acid sites in the zeolite that promote HT. One approach is to reduce the number of these acid sites by lowering the zeolite UCS. A lower zeolite UCS is typically achieved by reducing the RE content. A consequence of reducing catalyst zeolite content or UCS is a reduction in catalyst activity for conventional catalyst technologies. Low RE technology is critical to enable both maximum propylene yield and overall performance at a lower UCS, or equivalently lower RE.10 Specialized low RE systems deliver minimum HT without the conventional loss in activity and therefore profitability11 Key features for success include a high zeolite SAR for maximum stability; advanced zeolite technology for propylene; strong, selective matrices for cracking with minimal HT; and a 3.5

9

SCT-RT test data, resid 3.0

7 Δ C3=

6 5 4

log (1 + %ZSM-5)

⌬ propylene yield, wt%

8

2.5 2.0

R2 = 0.95

3 2 1 0

1.5

Decreasing base catalyst HTI 0

2

4

6

8 10 12 ZSM-5 crystal, wt%

FIG. 5. ⌬ propylene yield vs. ZSM-5 crystal.

102MARCH 2013 | HydrocarbonProcessing.com

14

16

18

20

1.0 0.26

0.28

0.30 0.32 0.34 0.36 Base catalyst HTI at 72 wt% conversion

FIG. 6. Base catalyst HTI influences propylene yield shift.

0.38

0.40

Clean Fuels

0.40

RxT Rx-Press

0.25

64

S

CCR

Ni

V

C3=

543

1.4

0.931 20.6

0.38

4.7

4.7

9

10.5

528

1.7

0.921

22.1

0.32

3.9

4.4

6.7

9

RFCC-3

545

1.8

0.912 23.5

0.3

4.4

8.4

2.2

10.5

RFCC-4

544

1.4

0.93

20.5

0.5

6.0

10

10

10.8

FCC-1

545

2.5

0.899 25.8

0.1

0.1

11.8

FCC-2

555

2.6

0.908 24.3

0.16

0.1

10.7

0.30 0.25 0.20 0.15

Low accessibility High accessibility

0.10 0

Max

RE on zeolite

FIG. 8. High accessibility reduces HTI. 1.0 0.8 24% higher relative activity 0.7 0.6 0.5

81% higher relative activity

0.4 0.2 0.12

62

API

RFCC-2

Increasing propylene

0.3

0.20 60

SG

RFCC-1

Relative zeolite activity

Ecat HTI

0.30

0.15

TABLE 6. AFX applications demonstrated

0.9

MPC catalyst Competitor A Competitor B

0.35

dium and sodium. The reduction in HT obtained from high catalyst accessibility means that RE can be higher, providing an activity advantage compared to a low accessibility product. A critical aspect in the maximum propylene FCC catalyst design for residue FCC is an optimization between minimizing HTI for maximum propylene and retaining sufficient activity for acceptable catalyst consumption and conversion, at an optimized RE level. The maximum propylene yield is closely correlated to the inherent HT activity of the catalyst. Zeolite stability begins to decline more rapidly than the HTI index as the RE content of zeolite approaches zero. The activity advantage provided by high catalyst accessibility increases as the target propylene yield also increases, as additional zeolite can be added to the catalyst particle to compensate. As the propylene target increases and the design HTI decreases, a maximum zeolite constraint is reached, and the activity advantage provided by a high accessibility structure increases rapidly (FIG. 9).

HTI

catalyst architecture with unsurpassed high mass transfer or diffusion character, tantamount to the highest catalyst accessibility. Another method of reducing HT is to produce catalyst particles with a highly accessible pore structure that allows very rapid diffusion of oil molecules in and out of the catalyst particle. A high accessibility catalyst allows the olefinic products of primary cracking to escape from the catalyst particle more quickly. HT is a bimolecular reaction that requires the reactants to be in close proximity of a pair of acid sites; so, by reducing the residence time of olefins within the catalyst particle, the rate of HT is reduced. A second benefit of high accessibility is increased preservation of primary cracking products. This reduces liquefied petroleum gas (LPG) selectivity but increases LPG and gasoline olefinicity. In a constrained system more LPG can then be produced by tailoring the zeolites, including those designed for propylene production, used in the catalyst. One company’s worldwide equilibrium catalyst database shows a clear example of the effect that high catalyst accessibility technology has upon HTI. The company routinely tests equilibrium catalysts (Ecats) collected from FCC units all over the world in a standard micro reactor test called the fluid simulation test (FST). This test is conducted with constant feed quality and reaction conditions, and allows direct comparison of catalyst yield selectivity. This company’s high accessibility catalyst technology exhibits a lower HTI compared to two major competitive catalyst technologies, as illustrated in FIG. 7. Based on these principles, the company developed a maximum propylene catalyst (MPC). The MPC is a new concept in FCC catalyst design specifically developed for operation at high propylene yield. A unique feature is the high accessibility structure used to reduce HT relative to conventional FCC catalysts at the same RE level. The FST test data in FIG. 8 illustrates reduction in HTI that is achieved by high accessibility compared to equivalent formulations prepared using a conventional low accessibility technology. For VGO applications, the additional reduction in HT allows maximum preservation of gasoline olefins for cracking to smaller olefins, delivering the highest possible propylene yield. In residue FCC, maximization of propylene poses additional challenges. To maximize propylene a low zeolite UCS is required, and this is achieved by reducing RE. However, at the higher levels of contaminant metals found in residue FCC operation, some RE is necessary to stabilize the zeolite crystal structure and provide resistance against deactivation by vana-

66

68 70 72 Ecat FST conversion, wt%

74

FIG. 7. HTI of various catalyst technology platforms.

76

78

80

0.16

0.20 0.24 Hydrogen transfer index

High accessibility Low accessibility 0.28

0.32

FIG. 9. Activity advantage of high accessibility increases with propylene yield. Hydrocarbon Processing | MARCH 2013

103

Clean Fuels Maximum propylene applications in practice. Laboratory evaluations have confirmed the understanding of the complex and competing reactions involved in propylene maximization. Many customer laboratories have tested MPCs from multiple suppliers in several applications, ranging from very light feedstocks to heavy residues. Tests have been conducted at pilot scale as well as bench scale. One example of a very clean feed application is illustrated in FIG. 10, where a third party benchmarked five competitive MPCs. It confirmed superior propylene yield of the high accessibility, low HT MPC catalyst. MPCs are preferred for challenging applications at or close to the plateau level of propylene, in high pressure operations 14

Propylene, wt%

13 12 11 10

MPC technology Other company technology Competitive technologies

9 8 74

76

78

80 Conversion, wt%

82

84

86

FIG. 10. Benchmarking catalysts for maximum propylene by third party testing.

and in resid FCC units processing heavy residue. They are also preferred for grassroots modern FCC units, which are all in some way operating on the edge of what is possible. LITERATURE CITED Greensfelder, B. S., et. al, Ind. Eng. Chem: 41:2573, 1949. 2 Feignet, F., Echard, M., and T. Gauthier, “Pilot plant testing for resid to propylene,” 8th FCC Forum. 3 Nieskens, M., et. al, “The assessment of carbenium-ion and radical cracking in fluid catalytic cracking,” Akzo Workshop: The MON of FCC naphtha, 1987. 4 Venuto, P. B. and E. T. Habib, Fluid Catalytic Cracking with Zeolite Catalysts, Marcel Dekker, New York, 1979. 5 Letzsch, W. S., “Petrochemical building blocks from heavy oils” PTQ, Summer 1999. 6 Ma’adhah, A., “Integrating refining and petrochemicals,” Hydrocarbon Engineering, June 2003. 7 Angevine, P. J., “Zeolites to Maximize FCC Olefins, Technology, Licensing and Commercialization,” ChemIndix 2007, SPS/05. 8 Plain, C., “R2R (RDS-RFCC): The Ideal Partnership,” BBTC, Barcelona 2008. 9 Lappas, A. A., et. al, “Production of liquid biofuels in an FCC pilot plant using waxes produced from biomass to liquid process,” Ind. Eng. Chem, 2011. 10 Yung, Y. and K. Bruno, “Low RE catalysts for FCC operations,” PTQ, Q1 2012. 11 Arriaga, R. and K. Bruno, “Applying low RE technology in demanding FCC operation,“ NPRA Annual Meeting, March 2012, San Diego, California. 12 Bruno, K. and Y. Yung, “Choosing the advanced option,” Hydrocarbon Engineering, September 2010. 1

CAREL POUWELS graduated as a chemical engineer from Delft University of Technology in 1987. Upon graduation, Mr. Pouwels joined Albemarle in its application research division. He now serves as the company’s global FCC specialist for resid applications. KEN BRUNO received his PhD in chemical engineering from the University of Notre Dame. He is Albemarle’s global applications technology manager for FCC.

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